Hydrocarbon Processing 10 2016.pdf

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PROCESS CONTROL AND INFORMATION SYSTEMS Use process knowledge management systems to accelerate innovation Optimize a CDU using process simulation and statistical modeling

BUSINESS TRENDS Refining outlook: Asia-Pacific

PROCESS ENGINEERING Build a diesel fuel performance additive, the right way The effect of various parameters on tray point efficiency

HITTING TOP QUARTILE MEANS Reclaiming the dead money buried in your operation

Emerson.com/Reliability

The Emerson logo is a trademark and a service mark of Emerson Electric Co. © 2016 Emerson Electric Co.

OCTOBER 2016 | Volume 95 Number 10 HydrocarbonProcessing.com

S–55

34 SPECIAL FOCUS: PROCESS CONTROL AND INFORMATION SYSTEMS 35 Optimize a CDU using process simulation and statistical modeling methods







41 The history, and possible future, of model-less multivariable control





A. Kern

45 Use process knowledge management systems to accelerate innovation





J. Bird, D. Seillier and E. Piazza

S. Hall and P. Mahoney

51 Measure naturally occurring radioactive material in polypropylene plants



D. Williams

DEPARTMENTS 4

Industry Perspectives



8

Business Trends



15

Industry Metrics



17

Global Project Data

96 Innovations 99 Marketplace 100

Advertiser Index

101 Events

ADVANCES IN SULFUR MANAGEMENT—SUPPLEMENT S-56 Stricter standards in sulfur treating

102 People



COLUMNS

B. Andrew

S-64 Low-sulfur projects dominate the downstream construction landscape

L. Nichols

PROCESS ENGINEERING AND OPTIMIZATION 71 Build a diesel fuel performance additive, the right way—Part 1



N. Kasiri, P. Jouybanpour and M. Reza Ehsani

MAINTENANCE AND RELIABILITY 83 Monitor medium-voltage switchgear in refineries









B. Jazayeri

93 Use dynamic simulation to maximize plant operating performance



M. A. Alós

GAS PROCESSING SUPPLEMENT GP-1 Technology and Business Information for the Global Gas Processing Industry Cover Image: In close cooperation with the Reliance Industries team, Barco Inc. worked out a total visualization solution, including system design, software and hardware, configuration and integration, reliability and operator comfort, for Reliance refinery operations. Photo courtesy of Barco Inc.

The 2016 Top Project award nominees are out! Reassess and redirect your approach to a long-term career







J. Murray

PROCESS CONTROL AND INSTRUMENTATION 88 Lessons learned in commercial scale-up of new chemical processes

Editorial Comment

19 Reliability 21

Automation Strategies

23

Project Management

27

Codes and Standards

K. E. Litz, K. Edison and J. Rankin

79 The effect of various parameters on tray point efficiency





7

G. G. Pipenger

75 Remove sulfur and nitrogen from liquid hydrocarbons with absorption process





Identify challenges in alarm management Global trends in energy savings: Are emerging technologies the solution? Overcome technical difficulties in field pumps ordered to API 610/ISO 13709

31 Viewpoint

New challenges demand innovative solutions

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Industry Perspectives

P. O. Box 2608 Houston, Texas 77252-2608, USA Phone: +1 (713) 529-4301 Fax: +1 (713) 520-4433 [email protected]

PUBLISHER

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EDITOR/ASSOCIATE PUBLISHER

Lee Nichols [email protected]

EDITORIAL

What is the future of the downstream HPI? What is the state of downstream processing capacity? What regions and sectors are seeing growth, and which are stagnant, or possibly shrinking? Capital-intensive investments are being made in every sector and region of the hydrocarbon processing industry (HPI). These investments are ensuring that global demand for petroleum products will be met in the future. At present, Hydrocarbon Processing’s Construction Boxscore Database is tracking over 2,100 projects around the world. These projects represent more than $1.6 T in total capital expenditures through 2030. The editors of Hydrocarbon Processing have provided their insight into the major trends affecting the global downstream oil and gas industry. Their views are expressed in Hydrocarbon Processing’s HPI Market Data 2017. Available now, the report provides detailed information on topics such as: • New and active construction projects around the globe • Latest developments in the refining, petrochemical and gas processing/LNG industries • Forecast spending for capital, maintenance and operating expenditures • Supply and demand of transportation fuels by region and country • The effect of low oil prices on the downstream industry • Market dynamics and trade flows of refined products. Hydrocarbon Processing’s HPI Market Data report has been produced for more than 40 years, and provides the HPI professional with comprehensive data to make informed, strategic decisions and recognize new opportunities in the global HPI. For more information on Hydrocarbon Processing’s HPI Market Data 2017, visit HydrocarbonProcessing.com and click “Market Data.” 600

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300 200 100 0

Africa

Asia-Pacific

Canada

Europe

Latin America Middle East

FIG. 1. Total active projects by region. Source: Hydrocarbon Processing’s Construction Boxscore Database.

4 OCTOBER 2016 | HydrocarbonProcessing.com

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Editorial Comment

LEE NICHOLS, EDITOR/ASSOCIATE PUBLISHER [email protected]

The 2016 Top Project award nominees are out! It’s that time again! Hydrocarbon Processing has announced the nominees for its annual Top Project awards. Using Hydrocarbon Processing’s Construction Boxscore Database, the editors of Hydrocarbon Processing have identified nine projects that are anticipated to heavily impact the global or regional downstream industry. All of the nominees will contribute significantly to the hydrocarbon processing industry, whether through contributing capital expenditures, satisfying domestic or adding to regional demand, diversifying product offerings, or the resurgence in refining and/or petrochemical processing capacity. These nine projects span the globe and represent over $60 B in total capital expenditures. The winners of this prestigious award over the last two years include: • Refining o 2014—Saudi Aramco and Total Refining and Petrochemcial Co.’s (SATORP) Jubail Refinery

o 2015—SOCAR’s Turkey Aegean Refinery (STAR) • Petrochemicals o 2014—Saudi Aramco and Dow Chemical’s SADARA Petrochemical Complex o 2015—Sasol’s Ethane Cracker and Derivatives Complex This year’s refining nominees represent nearly 800 Mbpd of new refining capacity by the end of the decade. All of these projects are located in non-OECD countries and represent a total capital investment of nearly $40 B. The four petrochemical nominees span four different regions, have a total cost of more than $23 B and represent more than 10 MMtpy of additional petrochemical products production by 2020. The choice is now up to you! Beginning October 1, readers of Hydrocarbon Processing can make their choices known in an exclusive online poll. The winners will be revealed in Hydrocarbon Processing’s December issue, with a subsequent ceremony to be held at the Hydrocarbon Processing Awards in June 2017.

Top refining project nominees Operator

Project

Location

Capacity

Bharat Petroleum Corp. Ltd.

Kochi Integrated Refinery expansion

Kochi, India

120 Mbpd (expansion)

Saudi Aramco

Jazan Refinery

Jazan, Saudi Arabia

400 Mbpd

Petroperu

Talara Refinery expansion and modernization

Talara, Peru

30 Mbpd (expansion)

Nghi Son Refinery and Petrochemicals

Nghi Son Refinery and Petrochemicals Complex

Nghi Son, Vietnam

Kuwait National Petroleum Co.

Clean Fuels Project

Mina Abdullah/Mina Al-Ahmadi, Kuwait

200 Mbpd 64 Mbpd (expansion)

Top petrochemical project nominees Operator

Project

Location

Capacity

Carbon Holdings

Tahrir Petrochemicals Complex

Ain Sokhna, Egypt

1.36 MMtpy

Petronas

Refining and Petrochemical Integrated Development (RAPID)

Pengerang, Johor, Malaysia

7.7 MMtpy

Dow Chemical

Oyster Creek PDH Unit

Freeport, Texas, US

750 Mtpy

Turkmengas

Kiyanly Petrochemical Complex

Kiyanly, Turkmenistan

480 Mtpy

INSIDE THIS ISSUE

8 Business Trends.

Hydrocarbon Processing continues its series on the global refining industry. Part 2 provides a detailed overview of the refining sector and capacity construction outlook for the Asia-Pacific region.

31 Viewpoint.

Rebecca Liebert, President and CEO of Honeywell UOP, discusses how the future refiner must emulate an entrepreneurial spirit— one driven by world-class research and development, innovative cost elimination, commercial excellence and modern control and automation systems.

34 Special Focus.

Process control and information systems are an integral part of refinery and plant operations. Advanced process control solutions provide hydrocarbon processing companies with inventive ways to manage difficult and complex operational and reporting applications. The special report addresses concerns regarding human-machine interfaces, as well as innovations in control devices, hardware and software.

S-55 Sulfur.

Around the world, legislation mandating decreased emissions and lower levels of airborne pollutants is coming into effect. In response, refiners are implementing operational and processing changes to reduce sulfur levels in transportation fuels. Hydrocarbon Processing’s Sulfur Supplement includes companies taking on this sulfur challenge, as well as those that handle this element every day.

75 Process Engineering.

Although adsorbents have yet to be broadly adopted in the treatment of refinery liquid hydrocarbon streams, the field is very active. As capacity, adsorption kinetics and selectivity continue to improve, widespread adoption may soon be seen. Hydrocarbon Processing | OCTOBER 2016 7

| Business Trends Hydrocarbon Processing continues its series on the global refining industry. Part 1 provided a look at the present state of the refining industry, new project developments, demand outlooks for the refining sector and the move to low-sulfur transportation fuels. Part 2 focuses on the refining sector and capacity construction outlook for the Asia-Pacific region. This detailed overview analyzes major trends in the region and offers data on planned refinery capacity additions, upgrades and grassroots facilities. Photo courtesy of Toyo Engineering.

LEE NICHOLS, EDITOR/ASSOCIATE PUBLISHER [email protected]

Business Trends Global refining overview—Part 2 At the time of publication, the Construction Boxscore Database was tracking nearly 775 total active refining projects around the world (FIG. 1). According to the International Energy Agency’s (IEA’s) Medium-Term Oil Market Report 2016, global oil demand will increase to nearly 102 MMbpd by 2022. The majority of oil demand growth will be in non-OECD Asia and the Middle East. By 2022, the IEA forecasts that global refining capacity will reach nearly 105 MMbpd (FIG. 2). The following is an overview of the refining sector and capacity construction outlook for the Asia-Pacific region. Part 3 of Hydrocarbon Processing’s global refining overview will be published in November. Asia-Pacific. The region has witnessed incredible growth over the past several years, and is forecast to account for the majority of oil demand growth through the rest of the decade. As OECD Asia countries have witnessed stagnant or decreasing oil demand, non-OECD Asia has led the charge in global consumption over the past few years. Asia’s oil refining capacity has increased by over 8 MMbpd over the past decade. According to BP data, total refining capacity in Asia has expanded from 24.3 MMbpd in 2005 to over 32.5 MMbpd in 2015. The majority of new regional refining capacity will be located in non-OECD countries—primarily China, India, Indonesia, Vietnam, Thailand and Malaysia. However, the slowdown in China’s economy is paralyzing the region’s refining margins. China’s overcapacity has transformed the country into a net exporter of refined fuels. In turn, Chinese refined fuels, primarily diesel and gasoline, have flooded the regional market. Coupled with that phenomenon, new Middle Eastern refineries have started exporting high-quality refined fuels into Asia, which has added to the glut. The supply overhang has impacted Asian re-

fining margins, as well as forced the region’s refiners to cut run rates to maintain profitability. The hope is that reduced utilization rates and seasonal maintenance will help draw down refined fuel stockpiles and increase margins in the long run. China. The country’s oil consumption has expanded by over 5 MMbpd in the past 10 years. Total oil consumption has reached nearly 12 MMbpd, making China the second-largest oil consumer on the planet, after the US. Although China is importing large amounts of crude oil, it is not a reflection of the robust China of a few years ago. The Chinese economy is slowing, and much of the additional imported crude oil is going toward its strategic petroleum reserves program, teapot refinery quotas and to replace lost production. The flood of oil in China has changed the country’s refined products dynamics, and has inadvertently turned the country into a major diesel exporter. As the country moves toward a more service-oriented economy, the need for diesel fuel to power heavy machinery

and heavy-duty trucks is slowing. Overcapacity and slowing industrial buildout have created an oversupply of diesel, which led the country to become a net diesel exporter in 2014. In 2015, Chinese diesel exports jumped 75%, reaching over 300 Mbpd in 2016. This trend will continue to negatively affect refining margins in Asia. Conversely, China’s demand for gasoline has skyrocketed. The need for additional gasoline to fuel passenger vehicles continues to increase. The country’s refining network was designed to produce diesel, so Chinese refiners have been adjusting their refineries to maximize output of gasoline. By mid-2016, however, it appeared that the country had overproduced. Crude processing rates topped 11 MMbpd in June, which was a record for the country. Gasoline supplies are beginning to stockpile. The surplus fuel is also being exported into the Asian market. If domestic gasoline demand does not increase, or if Chinese refiners do not cut run rates, then gasoline could become the country’s new diesel.

250

200

150

100

50

0

Africa

Asia-Pacific

Canada

Europe

Latin America

Middle East

US

FIG. 1. Total active refining projects by region, September 2016. Source: Hydrocarbon Processing’s Construction Boxscore Database. Hydrocarbon Processing | OCTOBER 2016 9

Business Trends The country’s refining network has increased substantially over the past decade. According to BP, China’s refining capacity increased from 7.6 MMbpd in 2005 to 14.3 MMbpd in 2015, and it is expected to increase an additional 1.5 MMbpd by 2020. Major refinery projects in China are listed in TABLE 1. Due to the surge in new refining capacity, the country has delayed many refining projects for up to a year or longer. Some of the major delays include the construction of the Kunming, Jieyang and Zhanjiang refineries, as well as the expansion of the Huabei refinery. Lastly, over the past year, the world has witnessed the rise of Chinese teapot refineries. Although these refineries tend to be less complex than their nationally owned counterparts, teapot refineries account for one-third of China’s total domestic refining capacity. In 2015, expansions in teapot

refining operations increased the independent refining sector’s total capacity to nearly 4.5 MMbpd. In 2015, China loosened restrictions on the teapots’ ability to secure crude oil from the international market. Now that Chinese independent refiners can utilize crude oil in lieu of lowquality fuel oil, non-state refineries are expected to boost run rates. This action will likely add to the fuel supply glut already being witnessed domestically, as well as in other countries in the Asia-Pacific region. Since the majority of Chinese independent refiners lack infrastructure to export their products to the global market, their refined products will be sold primarily to the domestic market. With increased refined product output, Chinese teapot refinery production will ultimately eat into the market share held by state-owned entities. Chinese teapots’ growing crude

TABLE 1. Major refinery projects in China Company

Project

Capacity, Mbpd

Completion

Sinopec Sinopec

Tianjin refinery

240

2020 or after

Zhanjiang refinery

300

2019/2020

Sinopec

Hainan refinery

100

2020 or after

Sinopec

Luoyang refinery

160

2020

CNPC/PetroChina

Huabei refinery

100

2017

CNPC/PetroChina

Anning refinery

260

2016

CNPC/PetroChina/PDVSA

Jieyang refinery

400

2021

PetroChina/Rosneft

Tianjin refinery

200

2020

CNOOC

Huizhou expansion

200

2017

Zhejiang Petrochemical

Dayushan Island complex

400

2020 or after

PetroChina/Qatar Petroleum

Taizhou refinery

300

2021 or after

Note: Data from the US EIA and Hydrocarbon Processing’s Construction Boxscore Database

2021 2015

Africa Latin America Middle East Other Asia China FSU Europe North America 0

5

10

15

20

FIG. 2. Global refinery capacity additions by region, MMbpd, 2015–2021. Source: IEA.

10 OCTOBER 2016 | HydrocarbonProcessing.com

25

processing market share may force stateowned refiners to either find additional export markets or cut run rates. India. The country is emerging as the globe’s new oil demand center, with burgeoning consumption providing huge potential for downstream oil and gas growth. With a GDP ranked in the top 10 globally and a large, growing population, India has continuously seen increases in demand for energy in all forms. According to IMF’s World Economic Outlook 2016, India is forecast to be the fastest-growing economy over the next decade. India’s domestic refining capacity sits at approximately 4.6 MMbpd. The country’s refining network has more than tripled in capacity over the past two decades. The country’s refining capacity is adequate to meet consumption rates. However, due to the projected growth in demand, additional refining capacity is needed. Both public and private refiners are planning, or have already commenced, major refinery expansions, upgrades and grassroots facility constructions. According to India’s 12th and 13th Five-Year Plans, domestic refining capacity is expected to reach approximately 366 MMtpy by the early 2020s. Major Indian refining programs include: • Hindustan Petroleum Corp. (HPCL) is investing nearly $8 B to increase refining capacity by two-thirds, as well as meet new fuel standard requirements. These investments will help HPCL hit its refining capacity target of 500 Mbpd by 2021. • Reliance is investing more than $18 B in refining and petrochemical capacity projects. This investment includes the $4.5-B petcoke gasification project, which will allow the company to eliminate approximately 6.5 MMtpy of petcoke. • Essar Oil will invest nearly $280 MM in upgrades to its Vadinar refinery over the next 2–3 years. These projects include upgrades to the facility’s naphtha hydrotreater unit, isomerization unit and continuous catalytic reformer units, along with the construction of additional sulfur recovery units. • Chennai Petroleum is conducting a feasibility study to boost capacity ninefold at its Nagapattinam

Business Trends refinery in Tamil Nadu. Total capacity will increase from 20 Mbpd to 180 Mbpd. This project will be carried out in addition to Chennai Petroleum’s $500-MM delayed coker and crude oil pipeline project at its Manali facility. • IOCL plans to invest $600 MM to upgrade its recently commissioned Paradip refinery in Odisha to meet new Bharat Stage 6 (BS-6) fuel standards. The company will also invest $6 B to expand domestic refining by 30% by the early 2020s. • Bharat Petroleum Corp. Ltd. (BPCL) is investing $4 B to expand and modernize its Kochi refinery. • Numaligarh Refinery, a subsidiary of BPCL, is planning a $3-B expansion project at its Numaligarh refinery in Assam. The project’s scope calls for the tripling of capacity from 60 Mbpd to 180 Mbpd. If built, the refinery would satisfy fuel demand in the northeast portion of the country. • IOCL, BPCL, HPCL and Engineers India are conducting a feasibility study on a 1.2-MMbpd refinery on India’s west coast. The initial capacity of Phase 1 will be approximately 800 Mbpd and cost nearly $15 B. In total, Indian refiners are investing more than $30 B in additional refining projects through the early 2020s. Capital expenditures are expected to be even higher due to new regulations to curb air pollution and produce Euro 6-standard fuels by 1Q 2020. India is skipping the implementation of Bharat Stage 5 (BS-5) and moving directly to BS-6 standards. The BS-6 regulation is being imposed four years ahead of schedule and calls for a 68% reduction in nitrogen oxide emissions. Indonesia. At present, the country lacks adequate refining capacity to satisfy the growing demand for refined products. The country’s refining network is in desperate need of expansions and upgrades to meet booming demand for petroleum products. To increase domestic refining capacity, Indonesia has developed the Refinery Development Master Plan (RDMP). This plan will make way for new downstream investments, primarily from heavy downstream players, to expand and modernize Indonesia’s ailing refining network.

The RDMP’s goal is to raise the country’s domestic refining capacity from 1 MMbpd to 2.3 MMbpd. Indonesia will accomplish the RDMP by upgrading five of its major refineries, and through the construction of grassroots facilities. The refinery upgrades, which will cost approximately $5 B each, will allow the country’s refineries to process heavier, less expensive crudes into high-quality products. The refineries involved in the RDMP are the

Balongan, Cilacap, Dumai, Plaju and Balikpapan facilities. The ambitious program would not be possible without the help of Saudi Aramco and Sinopec. The RDMP also includes the construction of grassroots facilities. These facilities include the Tuban and Bontang refinery projects. The RDMP is scheduled to be completed by 2025. Vietnam. The country has seen its oil consumption rates rise substantially over the past decade. At present, it has only one

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Business Trends operating refinery, located at Dung Quat. The 176-Mbpd plant is unable to satisfy domestic demand for refined products. As a result, the country is dependent on fuel imports. Vietnam has ambitious plans to completely eliminate the refined fuels supply gap. This program is outlined in the nation’s 2020–2025 development plan. The country’s initial plan called for the development of six large-scale refinery projects: • Dung Quat refinery expansion • Nghi Son refinery and petrochemical complex • Vung Ro refinery and petrochemical complex • Nhon Hoi refinery and petrochemical complex (abandoned) • Nam Van Phong refinery • Long Son refinery and petrochemical complex. These projects have the potential to add 1.36 MMbpd of new domestic refining capacity at a total cost of more than $50 B. This would reverse the country’s status from a net importer of refined products to a net exporter by 2020. However, several

variables may act as a deterrent to these plans. These variables include the need to secure a massive amount of crude oil feedstock and financial backing, overcapacity concerns that could lead to a glut of fuel in the Asia-Pacific region, the need to secure export supply contracts with other nations, and the threat of future government and environmental regulations. The massive amount of new domestic capacity could be overkill. Many analysts question where the surplus refined products will go. The region is already being flooded by excess diesel supplies from China. If Vietnam builds more refining capacity, can it compete with countries like China for market share? The country will also need to construct a large amount of new infrastructure to support the new refineries and export refined fuels. The uncertainty in global oil markets has already forced the cancelation of the $20-B Nhon Hoi refinery and petrochemical complex. The project would have seen the construction of a 400-Mbpd refinery and an olefins and aromatics plant with production of up to 5 MMtpy.

Regardless of the demise of the Nhon Hoi complex, the country is poised to become a major refined fuels producer by 2021. Vietnam’s second refinery is expected to begin operations in late 2017. The $7.5-B, 200-Mbpd Nghi Son refinery’s completion date has been delayed from its initial completion date of mid2017. The delay in completion means the country will continue to rely on refined products imports through 2017. Malaysia. The country is developing numerous projects under its Economic Transformation Program (ETP). Launched in 2010, the ETP’s goal is to transform Malaysia into a developed country by 2020. One of the major investments in the ETP is the construction of the Pengerang Integrated Petroleum Complex (PIPC). The PIPC includes the $27-B Refining and Petrochemical Integrated Development (RAPID) project. RAPID will include the construction of a 300-Mbpd refinery, a petrochemical complex with a total capacity of 7.7 MMtpy and an LNG regasification terminal. The project’s scheduled startup date is 2019.

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Business Trends Thailand. In response to increasing demand, the country has announced multiple expansion and upgrade projects to boost domestic refining capacity. Costs could eclipse $3 B by the early 2020s. The projects include the expansion and upgrade of Thai Oil’s refinery in Sriracha, and Bangchak Petroleum Public Co.’s Energy, Efficiency and Environment (3E) Improvement project in Bangkok. Thai Oil’s Sriracha expansion project aims to add 125 Mbpd of refining capacity to the already operational 275-Mbpd refinery. The company aims to make a final investment decision on the project in 2017. Thai Oil has already awarded contracts to upgrade the facility. The residue upgrading project will allow the refinery to produce Euro 5-standard fuels. The project’s study phase is expected to be completed by the end of 2016. Meanwhile, Bangchak Petroleum’s 3E project will increase the Bangchak refinery’s crude distillation capacity, debottleneck units to increase run rates, change out catalysts and add a cogeneration power plant.

South Korea. With over 3.1 MMbpd, South Korea has the sixth-largest refining capacity in the world. Nearly 80% of the country’s refining capacity resides in three giant complexes. Together, these refineries represent approximately 2.3 MMbpd of refining capacity, and all are among the top five largest refineries in the world; only Reliance’s Jamnagar complex in India and PDVSA’s Paraguana refinery complex in Venezuela have higher capacities. Refining capacities for South Korea’s “big three” are listed below: • SK Energy Ulsan refinery complex: 840 Mbpd (third-largest worldwide) • GS Caltex Yeosu refinery: 775 Mbpd (fourth-largest worldwide) • S-Oil Ulsan refinery: 669 Mbpd (fifth-largest worldwide). Like much of Asia, South Korean refiners experienced strong refining margins in 2015 as oil prices dropped by more than 50%. This caused a run-up in the country’s refining utilization rates to over 90%. The increased production allowed South Korea to maintain a dominant position in the Asia-Pacific fuels export market. How-

ever, this trend is changing as the region is being flooded by additional fuels from China and other countries. Some South Korean refiners, such as GS Caltex and Hyundai Oilbank Co., have already cut refinery run rates. Regardless of the decrease in refining utilization, South Korea is still investing in its downstream sector, with a focus on petrochemical and refining expansion projects. The most notable project is SOil’s Residue Upgrading Complex Project (RUCP). The project is being constructed in unison with the facility’s olefin downstream complex (ODC). The project is part of the company’s strategic growth initiative, which includes refining and petrochemical integration. RUCP will convert heavy fuel oil into high-value-added gasoline and olefins. The RUCP and ODC will act as an integrated complex. RUCP will supply its production as feedstock to the olefins plant. Both projects are expected to begin operations by 3Q 2018. Next month. Part 3 of this overview will appear in November.

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Hydrocarbon Processing | OCTOBER 2016 13

MIKE RHODES, MANAGING EDITOR [email protected]

Industry Metrics

An expanded version of Industry Metrics can be found online at HydrocarbonProcessing.com.

Global refining margins, 2015–2016* 20

10

5

7

Cracking spread, US$/bbl

Aug.-16

July-16

June-16

May-16

April-16

Mar.-16

Feb.-16

Jan.-16

Dec.-15

Aug.-16

July-16

June-16

May-16

April-16

Mar.-16

Aug.-16

Joly-16

June-16

May-16

April-16

Mar.-16

Feb.-16

Jan.-16

Dec.-15

Nov.-15

Sept.-16

Aug.-16

July-16

June-16

May-16

Mar.-16

30

Dubai Urals

20 10 0

Prem. gasoline Jet/kero

Gasoil Fuel oil

Sept.-16

Aug.-16

July-16

June-16

May-16

April-16

Mar.-16

Feb.-16

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Aug.-16

Aug.-16

July-16

June-16

May-16

April-16

Mar.-16

Feb.-16

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Aug.-15

-10 -20

Sept.-15

Aug.-15

Light sweet/medium sour crude spread, US$/bbl

Oct.-15

Singapore cracking spread vs. Oman, 2015–2016*

Brent dated vs. sour grades (Urals and Dubai) spread, 2015–2016* 8 6 4 2 0 -2 -4

April-16

Source: EIA Short-Term Energy Outlook, September 2016.

Gasoil Fuel oil

Feb.-16

Cracking spread, US$/bbl

-10 -20

Prem. gasoline Jet/kero

Jan.-16

2017-Q1

0

Dec.-15

2016-Q1

10

Nov.-15

2015-Q1

30 20

Oct.-15

2014-Q1

Stock change and balance, MMbpd

Supply and demand, MMbpd

6 5 4 3 2 1 0 -1 -2 -3

40

Aug.-15

2013-Q1

Sept.-15

Rotterdam cracking spread vs. Brent, 2015–2016*

World liquid fuel supply and demand, MMbpd Forecast

Prem. gasoline Jet/kero Diesel Fuel oil

Cracking spread, US$/bbl

July-15

Oil prices, $/bbl

A S O N D J F M A M J J A S O N D J F M A M J J A 2014 2015 2016

2012-Q1

Nov.-15

60 50 40 30 20 10 0 -10 -20

W. Texas Inter. Brent Blend Dubai Fateh Source: DOE

Stock change and balance World supply World demand

Oct.-15

US Gulf cracking spread vs. WTI, 2015–2016*

Selected world oil prices, $/bbl

100 98 96 94 92 90 88 86 84 82 2011-Q1

Feb.-16

Production equals US marketed production, wet gas. Source: EIA.

Jan.-16

60

Japan Singapore

Dec.-15

J A S O N D J F M A M J J A S O N D J F M A M J J 2014 2015 2016

US EU 16

70

Aug.-15

0

80

Sept.-15

20

2 1 0

Nov.-15

Monthly price (Henry Hub) 12-month price avg. Production

Oct.-15

3

40

90

Sept.-15

4

Aug.-15

60

Utilization rates, %

100

5

Gas prices, $/Mcf

Production, Bcfd

Global refining utilization rates, 2015–2016*

6

80

120 110 100 90 80 70 60 50 40 30 20

Sept.-15

Aug.-15

0

US gas production (Bcfd) and prices ($/Mcf) 100

WTI, US Gulf Brent, Rotterdam Oman, Singapore

15

Margins, US$/bbl

US gasoline demand contributed to falling inventories. Operational problems at several refineries and the potential impact of a tropical storm strengthened margins. Asian margins recovered slightly amidst firm demand and falling inventories ahead of autumn maintenance. Improvements seen in European gasoline crack spreads allowed refinery margins to recover, despite the oversupply environment.

* Material published permission of the OPEC Secretariat; copyright 2016; all rights reserved; OPEC Monthly Oil Market Report, September 2016. Hydrocarbon Processing | OCTOBER 2016 15

The new WDG-V. Impressing even the world’s most demanding combustion manager.

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LEE NICHOLS, EDITOR/ASSOCIATE PUBLISHER [email protected]

Global Project Data New project announcements have hovered in the low teens for several months. This trend is due to several factors, which include an oversupplied fuels market, economic uncertainty and reduced funds to expand capacity due to the drop in oil prices. Although new project announcements have decreased over the past several months, announced global downstream project investments are

near an all-time high. Hydrocarbon Processing’s Construction Boxscore Database is tracking over $1.6 T in announced projects around the world. The majority of this new capacity is located in non-OECD countries in the Asia-Pacific and Middle East regions, as well as petrochemical and gas processing/LNG capacity additions in the US.

$222 B

Canada

$168 B

$332 B

Europe $275 B

US

$430 B

$140 B

Middle East Africa

$77 B

Asia-Pacific

Latin America

Total announced downstream project investments by region, 2016–2030 27

17

18

27

26 20

18

18

7% Africa 18% US

21 15

13

12

12

13

Aug.- Sept.- Oct.- Nov.- Dec.- Jan.- Feb.- Mar.- April- May- June- July- Aug.- Sept.15 15 15 15 15 16 16 16 16 16 16 16 16 16

Boxscore new project announcements, August 2015–present

30% Asia-Pacific 21% Middle East 3% Canada 13% Europe

8% Latin America Market share analysis of active downstream projects by region

Detailed and up-to-date information for active construction projects in the refining, gas processing and petrochemical industries across the globe | ConstructionBoxscore.com Hydrocarbon Processing | OCTOBER 2016 17

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Reliability

HEINZ P. BLOCH, RELIABILITY/EQUIPMENT EDITOR [email protected]

Reassess and redirect your approach to a long-term career It was not too long ago that a prospective employee could expect a 30-year career within an oil refinery or a petrochemical plant. Overall, the qualifications for such jobs were dependability, integrity and the desire to learn. It helped to maintain a focus on being better than average. For most of us, being in the upper 50% of professional performers was not difficult. Our goal was not necessarily to display a passion for our job; instead, it was to carry out our assignments well enough that our absence would have been noted. Advice for job holders. Chances are that not much has changed in this assessment. For someone who is presently employed, this philosophy is still the general roadmap for maintaining employment. However, our approach and actions may have to be redirected if we work at a plant that is endangered by corporate takeovers, consolidation, automation or failure to keep up with top-quartile performers. In this case, it might be best to obtain a more detailed roadmap, figuratively speaking. Instead of being a good instrument technician, we should strive to learn the interaction of electronics, pneumatics and computer-guided automated processes. We may have to learn more about the logic that drives the meshing of software and hardware. That means that we must actively pursue and make use of adult education opportunities. For example, could acquiring a foreign language skill set us apart from a slightly more complacent competitor? Advice to job seekers. Suppose that

you are among the young people that are preparing themselves for the job market. Industry employment opportunities seek new entrants; after all, people retire and must be replaced. Suppose that you are enrolled in a community college and are taking a course in automated drafting of

electrical schematics. Would it not be advantageous to be in a position to explain, during a job interview, to what the various symbols on your schematic refer? Along those lines, and all things being equal, would it not also be advantageous to list on a resume, “Fluency in Spanish as a secondary language?” It never hurts to develop widely varying interests. No one has ever been dismissed from a job interview for knowing the difference between oils and greases, or between rolling element bearings and sleeve bearings. Conversely, others may not be eligible for employment if they spent years in a machine shop and never bothered to understand the limitations of one lubricant relative to the other, or the maintenance requirements and lubricant application details of each. Potential advice for all. During a job

search, it pays to know what the company produces and how the processes work that yield the product. An entry-level applicant is not expected to be an expert. However, it is expected that you have shown interest in how things are being made, what raw materials are being used, what the end use is for the products that leave the plant, and whether those products are shipped in wooden crates, cardboard boxes or railroad tank cars. Know what the interviewer does and why you are being interviewed: All companies—in particular, within the hydrocarbon processing industry (HPI) segment—seek applicants who are focused, dependable and teachable. When you are asked about your goals, or where you see yourself in 10 years, try to explain how you intend to add value, not just collect a paycheck. Never give the interviewer an immature answer by saying that your goal is to secure the available position. Let us revert back to the possibility of a job holder becoming a job seeker in

a struggling economy. Today, while still gainfully employed operating a filling or packaging machine for a motor oil producer, the groundwork for a wider range of future jobs should be laid. Ample time exists to become familiar with the fact that these machines have bearings, and that those bearings require lubrication. Suppose that time had been spent observing and understanding the maintenance details of these carton setups, and oil dispensing and quality-control monitoring machines or devices. Consider how those proactive efforts might open doors in a potential search for employment, because virtually all maintenance involves lubricants and lubricant application. Face the facts: Good jobs go to valueadders. Don’t put all your energy into an expensive college education, only to graduate near the bottom of the class. Become a person who moves from the level of being inquisitive to the knowledge level, and from there to the next level—wisdom. The process of becoming a value-adder can begin somewhere between the inquisitive and knowledge levels. Set a time budget; write the numbers and targets on a piece of paper and be specific. Resolutely distinguish between your needs and your wants. Our needs are finite, while our wants are limitless and have no boundaries. HEINZ P. BLOCH resides in Westminster, Colorado. His professional career commenced in 1962 and included long-term assignments as Exxon Chemical’s regional machinery specialist for the US. He has authored over 650 publications, among them 19 comprehensive books. Mr. Bloch holds BS and MS degrees in mechanical engineering. He is an ASME life fellow and maintains registration as a professional engineer in New Jersey and Texas. Hydrocarbon Processing | OCTOBER 2016 19

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Automation Strategies

LARRY O’BRIEN, VICE PRESIDENT ARC Advisory Group

Identify challenges in alarm management ARC Advisory Group recently conducted a survey on practices and trends in alarm management in the process industries. The objective was to learn how end users, suppliers, consultants and system integrators are approaching the often challenging issue of migrating existing alarm management applications. Alarm management, in general, continues to be a significant issue in process plants, driven largely by the need to conform to current standards and best practices, such as ISA 18.2, EEMUA 191 and IEC 62682. The primary goal of these standards and practices is to develop a continuous improvement approach to alarm management, and to ease the alarm burden on operators. As older alarm management systems become obsolete, end users must migrate to new applications. In many cases, users are taking advantage of this opportunity to improve their alarm management philosophy and implement some of the newer aspects of these solutions, such as dynamic alarms that can change in lock-step with the dynamically changing state of the plant.

graphics, as this standard is very popular in North America. However, it is clear that many users outside of North America also follow the standard. Close to 20% of respondents indicated that they follow the IEC 62682 standard, which closely mirrors ISA 18.2. ARC asked respondents to briefly describe the three primary challenges they faced regarding alarm management project implementation. These challenges are outlined below.

Survey data. ARC received more than 170 responses to the survey. Close to half of the total respondents were end users, while consultants represented over 19% of respondents. Suppliers represented a relatively small portion of total respondents (17.5%). Other respondents included original equipment manufacturers (OEMs) and skid-mounted equipment manufacturers, and system integrators. While not all respondents answered all survey questions, a general alignment in the responses— whether from end users, suppliers or third parties—was noted. On an industry basis, the bulk of responses came from the oil and gas sector (> 24%), while petrochemicals and bulk chemicals accounted for 19% of responses. Regionally, most respondents were from North America and Western Europe, which collectively have the largest installed base of advanced alarm management applications. Most survey respondents recently implemented a new alarm management and rationalization project at their company or facility, many on a company-wide basis. Clearly, strong activity in alarm management and rationalization exists. ARC believes that this effort will only escalate in the next few years as many end users face the need to migrate from older alarm management platforms, while other users that have yet to implement advanced alarm management solutions will embark on new projects. More than 35% of respondents disclosed that they were applying minor upgrades to existing applications. The remainder were fairly evenly distributed among those that are implementing brand new projects, those that are migrating to a new solution from a new supplier, and those that are migrating to a new solution from the same supplier. Most respondents (72%) indicated that they adhere to the ISA 18.2 standard. This is consistent with the survey demo-

Challenge #2: Lack of subject matter experts. Secondary challenges can also include human issues such as buy-in, ease of use and basic trials such as time and resource allocation. However, we are beginning to see more specific technical and implementation challenges as secondary issues, including alarm philosophy development challenges, configuration issues, developing or redefining key performance indicators (KPIs), ease of use and database issues. Personnel issues also begin to become more specific and clearly defined, such as finding sufficient subject matter experts (SMEs), developing common work processes and procedures across the enterprise, and overcoming resistance to change by operators and other personnel.

Challenge #1: Securing management buy-in and resource allocation. Many of the primary challenges listed

deal with human issues, such as obtaining “buy-in” from operators and management, and finding the appropriate amount of time, resources and training to effectively complete the project. Cost and funding issues are also prevalent throughout the responses. Actually performing the alarm rationalization aspect of the project was listed as a primary challenge.

Challenge #3: Alarm rationalization and consistency. ARC still observes cost and resource concerns repeated as tertiary challenges, but more specific personnel-oriented and technology issues, such as keeping alarm rationalization up to date, the management of change, and the implementation of dynamic alarming, remain relevant. Achieving consistency in alarm management while dealing with disparate sources of data was also pinpointed as a key challenge. LARRY O’BRIEN is the vice president of process automation at ARC Advisory Group, and has more than 20 years of experience working in the automation and consulting business. Mr. O’Brien has co-authored numerous reports at ARC, including the “Collaborative Process Automation System 2.0,” the “Distributed Control System Market Size and Forecast,” the “Control System Migration Survival Manual,” and the “Automation Supplier Provided Services Market Size and Forecast.” He also served three years as global marketing manager at the Fieldbus Foundation. Hydrocarbon Processing | OCTOBER 2016 21

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Project Management

DOUG HUTTON AND JUAN GÓMEZ-PRADO KBC Advanced Technologies, Manchester, UK

Global trends in energy savings: Are emerging technologies the solution? Despite low energy prices, energy conservation continues to be an important component of business strategy in the oil refining and chemical industries. Driven globally by the need to increase operating margins or meet government mandates (e.g., natural resources in Saudi Arabia, or European emissions standards), operating companies are seeking new solutions to improve energy performance. Energy efficiency is a mature area of operational optimization in refineries and petrochemical complexes. Many of the tools, techniques and technologies are familiar to the industry. A consequence of this familiarity may be the perception that all of the usual avenues for improvement have been exhausted. This perception is leading to a common question in many energy efficiency discussions about what new technologies are available for reducing energy use.1 While this appears to contradict the drive for no-/low-cost improvements, it does recognize that saving energy is still important. The reality is that the bulk of “new” technologies have been known for many years and have yet to achieve the breakthroughs expected of them. At the same time, experience shows that, across the refining and petrochemical industries, significant energy reduction (and, therefore, margin improvement) can still be achieved. This article discusses two trends in energy management: 1. The drive for no-/low-cost improvements 2. The upsurge in interest in new technologies. Energy management systems. Analyzing the results of more

than 30 years of energy studies highlights a number of trends: • The best energy performance in the refining sector has not improved in absolute terms in the last 20 years. However,

overall average performance has improved, even if the best performance has not. • New refinery units are generally not optimized for energy performance on a standalone basis. • New units are seldom optimally integrated from a site-wide perspective. • General reluctance exists to implement new technologies for energy improvement in refining; in contrast, acceptance of new technology has been stronger in petrochemicals. These trends have been consistent even through the preceding period of high energy prices and low investment costs, so they are not driven by the present low energy prices. In a sense, low energy prices are not changing the industry’s fundamental approach to energy performance. The industry is, however, putting off investment and focusing instead on sustaining the resulting benefits of no-/low-cost improvements. Specific areas receiving significant attention include: • Operational savings 0 Benchmarking against industry best practices to set performance targets 0 Software tools to attain and sustain top performance (e.g., metric tracking, utility optimization, advanced process control, real-time optimization) • An energy management system (EMS)2 to implement the processes and organization to manage and continually improve energy performance. A combination of constrained capital budgets and increased drive to improve operating margins means that there is a good business case for developing an EMS. The successful implementation of energy management results in the capture of low-cost opportunities and sustained benefits.

TABLE 1. Elements of an effective EMS Personnel

Methodologies

Technology

Senior managers Energy team • Corporate • Onsite Plant team • Plant manager • Plant engineers • Supervisors • Operators Technical support team Utilities support team

Operations • Online performance monitoring • Online utility system optimization Technical support • Opportunity identification and investment planning • Operational scenario planning • Production planning support • Design Maintenance practices Knowledge management • Capability development • Technology awareness • Documentation • Innovation

Energy metric tools • Tracking applications • Unit monitoring • Equipment monitoring Opportunity identification • Simulation • Pinch analysis • Utility modeling Scenario planning • Steam system simulation • Steam system optimization Subject matter expertise • Best practices • Leading technologies Hydrocarbon Processing | OCTOBER 2016 23

Project Management TABLE 2. Examples of standard, leading and emerging technologies

Process level  Utility level

Standard technologies

Leading technologies

Emerging technologies

Gas turbines for cogeneration Hot water loops Absorption refrigeration

Online utility optimization Integrated fuel management Proprietary fluid and/or vapor cycles for improving thermal efficiency Renewables (wind, solar) Fuel cells

• Distributed control systems (DCS) • Simulation • Pinch analysis on all new designs

Monitoring and targeting Predictive modeling Pinch analysis for retrofit Advanced process control

Site-wide integration Trigeneration

• • • •

Variable-speed drives Liquid expanders Advanced heat exchangers Thermocompressors

Heat exchanger fouling monitoring Heat pumps Dividing-wall distillation Dual-effect distillation High-efficiency distillation column internals

• • • • • • •

Condensate returned to boiler plant Steam trap and leak repair programs Optimal insulation thicknesses Minimize pump redundancy Optimized furnace stack O2 and temperature 75%–80% efficient backpressure turbines Turbine/motor auto-start facilities

Distillation feed conditioning Distillation reflux optimization Equipment-level optimization Heat-integrated design

amount of effort put into energy management and overall energy performance. Results from recent energy studies indicate that a properly implemented EMS could lead to an energy performance improvement of up to 15%.

100

Effective efficiency, %

80

Process heating Condensing steam turbine Absorption chilling Organic proprietary water/vapor cycle

60

40

20 0 70

90

110 Approximate temperature range, °C

130

150

FIG. 1. Relative efficiency of low-grade heat recovery methods. TABLE 1 illustrates a number of elements of a refining/petrochemical industry best-practice EMS, grouped into three categories: personnel, technologies and methodologies. To be effective, an EMS must consider these three aspects together. The development of such a system is not to be approached casually. A best-practice EMS can be thought of as an iceberg, with most of the mass beneath the surface. The reporting system (dashboards, reports) is the visible tip of the iceberg, while other elements work beneath the surface: • Skills development • Organizational development and alignment with corporate strategy • Work processes • Data reconciliation/validation • Smart target setting • Optimizing assets • Targeting against technology standards. A common hurdle to implementing a well-planned and wellimplemented EMS is that it is difficult to estimate the benefits in advance. However, a strong correlation exists between the

24 OCTOBER 2016 | HydrocarbonProcessing.com

Return on investment. Many operating companies have restricted capital budgets and are focusing performance improvement on asset optimization. Analysis of results from studies has shown that, on average, energy consumption can be reduced by 4% through operational improvements alone. These energy savings double to 8% if investment projects with a simple payback of less than 1 yr are considered, and may increase further, to 13%, if investments with a simple payback of less than 3 yr are included. A good case for investing in energy performance improvement can be made. A common question in energy reviews concerns new technologies that are available for improving energy performance. In discussions of the topic, the following broad definitions have been used: • Standard technologies are widely implemented across the industry • Leading technologies are being implemented by leaders in energy performance • Emerging technologies have been limited, but are increasingly reaching new applications. These definitions are illustrated in TABLE 2. Note: The examples shown in TABLE 2 are by no means exhaustive. Globally, efforts have been focused primarily on “standard” technologies. Relying on mature technologies provides a high degree of confidence that a solution will work, and it usually means that the implementation is relatively inexpensive. Projects with a low cost and high return have always taken precedence. Additionally: • Real technological “game-changers” have been scarce in recent years, with most advances being evolutionary changes to “standard” pieces of equipment. • Many operating companies are reluctant to embrace new technologies until they have seen it elsewhere.

Project Management 60 50 Split of financial benefits, %

• New technologies must be accepted first by licensors before they will see widespread application within process designs. • Margin improvement through yield improvement is still seen as more attractive than energy savings; unfamiliar energy technologies will lose in competition for financial resources. The implications are that emerging technologies in industry practice are actually years or decades old. For example, lowgrade heat recovery has long been an area of research and development because the great majority of refineries and chemical plants have a large excess of heat between 30°C and 150°C that is rejected to air or cooling water. Finding a way to recover this energy would, therefore, be very lucrative. Potential technologies to convert heat into a more useful form of energy are absorption refrigeration and proprietary cycles using fluid and/or vapor to improve thermal power plant efficiency. In principle, these technologies could be utilized, but obstacles exist: • Distributed sources. In most chemical plants and refineries, heat sources that could be captured by these technologies are numerous and at a range of temperatures and heat contents. As a consequence, it is not always possible to exploit economies of scale that would justify investment. • Limited users. Absorption refrigeration is only worth considering if a suitable use for chilling exists. This may include replacing mechanical chillers or capturing additional light products. Whatever the case, the amount that can be utilized is limited. • Relative efficiency. As shown in FIG. 1, the most efficient use of low-grade heat is through process heat recovery— i.e., a standard technology. Assuming that an appropriate use can be identified, 100% of the waste heat can be recovered, making it the preferred option and one that will be maximized. The more heat that is recovered, the lower the temperature of the remaining low-grade heat and, consequently, the lower the efficiency of alternative means of capture. Ultimately, the low efficiency of the organic fluid/vapor cycle for energy recovery makes it unattractive, in most cases. Emerging technologies that manage to gain acceptance are generally a result of the combination of a number of factors. For example: • A particular industry sector discovers a unique application • Separate developments in other areas provide synergy • Evaluation techniques improve • Market economics become supportive (changing fuel and power prices). While a perception that new technologies are an untapped means of radically improving energy performance exists, the reality is that this is not the case. In fact, the authors’ experience shows that the greatest financial benefits are realized by applying standard technologies. Typically, energy performance improvement programs yield savings of 10%–15% of base energy costs. Of these savings, approximately 80% of benefits are realized through operational changes and standard technologies, as shown in FIG. 2. Emerging technologies do play a part, with approximately 10% of the

40 30 20 10 0 Zero cost

Standard

Leading

Emerging

FIG. 2. Split of financial benefits for energy performance improvement programs.

total benefit. It is worth noting that individual applications typically offer larger-than-average savings. Takeaway. The global trend of operating companies seeking

new solutions to improve energy performance continues to develop. However, in general, the better energy performers are those that are getting the basics right and looking to new solutions to gain an extra competitive edge. The trends observed in recent energy studies include: • Focus on no-/low-cost solutions o Despite increased industry familiarity with the tools and techniques of energy conservation, recent studies have revealed that financial benefits are mainly realized by applying standard technologies • Development of EMSs o Many companies are focusing on EMSs to capture no-/low-cost opportunities and sustain the benefits o A strong correlation exists between well-planned and well-implemented EMSs and energy performance • Increasing consideration of emerging technologies, particularly for low-grade heat recovery o Technological “game-changers” have been few in recent years o Low-grade heat recovery is of consistent interest, but application of emerging technologies is low. LITERATURE CITED KBC, “Innovative energy saving technology for the reduction of carbon dioxide emission from energy-intensive industry,” International trends of energy conservation in heavy chemical industry, 2014. 2 KBC, Energy Management System, KBC handbook. 1

DOUG HUTTON is a principal consultant at KBC Process Technology in the UK. He holds a first-class honors degree and a PhD in chemical engineering from the University of Edinburgh in the UK. Dr. Hutton has 20 years of industry experience, including leading projects to develop corporate energy strategies and energy management systems for refining and petrochemicals companies. JUAN GÓMEZ-PRADO is a senior consultant at KBC Process Technology in the UK. He holds a chemical engineering degree from Universidad Simón Bolívar in Venezuela, and an MS degree and PhD from the University of Manchester in the UK.

Hydrocarbon Processing | OCTOBER 2016 25

Codes and Standards

GOPAL MURTI, SENIOR CONSULTANT The Augustus Group, Montgomery, Texas

Overcome technical difficulties in field pumps ordered to API 610/ISO 13709 The continued objective of this column series is to increase awareness of code requirements, interpretations and limitations as they stand today, and the options available to engineers for “alternative engineering.” Another intention is to forewarn end users of the pitfalls of adhering to codes and standards indiscriminately, particularly those that have recently undergone major changes. The information provided here is based on actual field implementations and the resulting satisfactory experiences. If engineering or other professional services are required, then the assistance of a competent professional authority should be sought. The American Petroleum Institute’s API Standard 610, “Centrifugal pumps for petroleum, petrochemical and natural gas industries,” was first published in 1954, and it continued as an exclusive US standard until 2003. Beginning with the 9th edition issued that year, API considered amalgamation with the International Standards Organization (ISO), and this edition was declared as “…technically equivalent to the ISO final draft international standard 13709.” Beginning with the 10th edition in 2004, both documents were affirmed as “identical.” The connection continued with the release of the 11th edition (2011), but the proposed 12th edition may indicate a looming “divorce” with ISO 13709. The technical difficulties faced in the field for pumps ordered to API 610/ISO 13709 are discussed here. The problem arises on pump sets originating mainly from European suppliers. Does the 12th edition and its proposed delinking to ISO solve the problems, or will pump users continue to face compliance challenges? One of the visible differences with the proposed delinking would be that imperial units will take precedence over metric units, a reversal from past editions. Governing metric units in past editions would be placed into brackets, and the bracketed in-

formation would then become the governing information. This is a cosmetic change and is not considered a significant issue. However, it is feared that the proposed 12th edition may be more complicated than past editions. Pump application engineers must pay particular attention to these areas, and a careful approach during pump selection stages should help avoid potential problems and save operations and maintenance engineers the burden of carrying out expensive field modifications. In addition to specifying API 610, remedial measures could include a supplementary project specification with sketches.

DISCREPENCIES AND AREAS OF CONCERN Dimensional incompatibility. Per the draft of the 12th edition issued in May

2014, section 6.4.2.2 would continue to state, “All steel flanges shall, as a minimum requirement, conform to the dimensional requirements of ISO 7005-1 PN50.” A footnote would clarify, “For the purpose of these provisions, ISO 7005-1 PN50 and European standard (EN) 1759-1 Class 300 are equivalent to American National Standards Institute/American Society of Mechanical Engineers (ANSI/ ASME) B16.5 Class 300 and ANSI/ ASME B16.47 Class 300.” The statement that ISO 7005-1 is equivalent to ANSI/ASME B16.5 is factually incorrect, and is identified as a source of problems in the field. These two standards are not dimensionally identical. The intent of the footnote, “For the purpose of these provisions…,” may be that the pressure-temperature ratings of ASME and ISO flanges are similar. This state-

TABLE 1. History of ISO-7005-1 and ASME-B16.5 Standard

Edition/Yr Pages Remarks

ISO 7005-1

1st /1991

88

Includes pressure-temperature ratings and dimensional details. Includes PN50, referred as base rating by API-610.

ISO 7005-1

2nd /2011

22

Pressure-temperature ratings and dimensional tables removed, and EN-1092 referred.

EN 1092-1

2007

126

Includes pressure-temperature ratings and dimensional tables. PN50 not listed. Pressure rating designations substantially changed. Disassociation with ASME B16.5.

ASME B16.5

Issued in 1998, 2003, 2009 and 2013; older history not available.

FIG. 1. Bolt-hole drilling encroaching on the pressure-containing body. Hydrocarbon Processing | OCTOBER 2016 27

Codes and Standards ment may be sufficient for pump application engineers and manufacturers. Per an inference of the API footnote, engineers and manufacturers may proceed with pump flanges that meet either standard, but what happens when the pump reaches the field for installation? ASME B16.5 and ISO 7005-1 standards differ in flange-mating dimensions. The flange thicknesses, bolt-hole diameters and bolt diameters also vary in cerPiping side

Permitted by ISO 7005 Required by ASME B16.5 Pump side

FIG. 2. The difference in the ISO flange standard and the ASME flange standard.

tain sizes. The issue comes to light when a pump set manufactured to API 610/ISO 13709 reaches a facility that is essentially built to US standards. The pump casting that conforms to ISO standards arrives from foundries with the flanges undrilled. The pump machine shop reads the order data sheet, and an ASME drilling template is used. This invariably results in bolt-holes encroaching on the pump pressure-containing body, as shown in FIG. 1. This is deemed unacceptable and unsafe, and is clearly visible to safety engineers and visiting insurance teams, which will mandate the replacement of the pump set, a difficult task at site. Bolting incompatibility. ISO standards allow threaded bolt-holes in the pump body, while the US flange standard ASME B16.5 requires the use of a stud with two nuts. The schematic is illustrated in FIG. 2. Due to incompatibility issues, field engineers must undertake various modifications, such as enlarging bolt-holes to shave off threading (which, in turn, eats away metal from the pressure-containing

body), searching for longer bolts, and so on. While focusing on an expedited commissioning of the equipment, little attention is paid to the integrity and reliability of the planned machinery modifications. Additional nozzle sizes in ISO. ISO standards allow six additional sizes for the pump inlet and outlet: 1¼ in. (32 mm), 2½ in. (65 mm), 3½ in. (90 mm), 5 in. (125 mm), 7 in. (175 mm) and 9 in. (225 mm) nominal pipe size (NPS). These sizes are explicitly not permitted per section 6.4.1.1 in API 610. A pump’s hydraulic efficiency is affected by fluid velocities at the pump inlet and outlet. A wider choice in ISO means that pumps with higher efficiencies may be offered for certain combinations of flow and head. Many customers are, therefore, inclined to buy high-efficiency ISO/EN pump sets, despite the fact that API 610 prohibits these sizes. Which ISOs to follow? API 610 section 2, “Normative references,” affirms several ISO/EN standards as indispensable. It further states that for dated references,

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Codes and Standards only the edition cited applies; for undated references, the latest edition of the referenced document applies. Out of 34 listed ISO/EN standards, only three are listed with a dated reference. ISO 7005-1 is not dated, which implies that the latest edition is to be followed. This is the source of confusion and dimensional incompatibilities. Upon review of the draft of the 12th edition, it appears likely that the confusion will remain. An awareness of the brief history of ISO 7005-1 is prudent (TABLE 1). The standard was first issued in 1991, listing pressuretemperature ratings and dimensional details for flanges. ISO endeavored to make this document similar to the prevailing ASME B16.5 in terms of coverage. ISO 7005-1 included the following pressure classes: pressure nominal (PN) 2.5, PN6, PN10, PN16, PN20, PN25, PN40, PN50, PN110, PN150, PN260 and PN420. It included pressure class PN50, which API 650 picked up as the base pressure rating. In 2011, ISO 7005-1 was trimmed from 88 pages to 22 pages. Pressure-temperature ratings and dimensional tables were removed, and readers were advised to refer to EN 1092-1:2007. EN 10921:2007 comprises the following pressure classes: PN2.5, PN6, PN10, PN16, PN25, PN40, PN63, PN100, PN160, PN250, PN320 and PN400. A large variation and reclassification is seen. Only pressure classes PN 2.5, PN6, PN10, PN16, PN20, PN25 and PN40 remain, and those remaining classes have been redesignated or changed. A more detailed discussion is beyond the scope of this article. It is also seen that pressure class PN50 (equivalent to ASME 300) is not included in the latest EN 1092-1:2007, which is the backbone of API 610. Realizing that PN classes are no longer equivalent to ASME B16.5, ISO 7005-1 (2011 Ed.) section 4.1.2 declares that a flange series shall be specified as PN or class: • If a PN series is specified, flanges shall be in accordance with EN 1092-1 • If a class series is specified, flanges shall be in accordance with ANSI/ ASME B16.5 or ANSI/ASME B16.47, as applicable. This assertion requires that the purchaser be specific in ordering either ISO/EN PN series or ASME-class series. However, adhering to the footnote of API

610, it appears that API is still referring to the 1st edition of ISO 7005-1. Between revisions, ISO maintains a significant time interval that coincides with major change implementations. This time interval is understood and expected, as ISO standards have a wider global audience and require host-country approval, as opposed to ASME, which is able to issue comparatively minor updates at frequent intervals. For instance, the 2nd

edition of ISO 7005-1 was issued after a gap of 20 years. During this same period, ASME-B16.5 underwent a minimum of four editions. So, complete disassociation between API and ISO is desirable, as the standards simply do not keep pace with each other and have different audiences. For example, ISO 15649, “Petroleum and natural gas industries—Piping,” remains the same as when it was first issued in 2001. During the same period, ASME

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 29

Codes and Standards B31.3 has seen seven revisions—one every two years. Flange outside diameter mismatch.

A new section, 6.4.2.2.1, has been proposed to be added to the 12th edition. “Unless otherwise specified, the outside diameter of ANSI/ASME B16.5 flanges shall comply with the negative tolerances stated on EN 1092-1 Table 22.” ASME B16.5 flanges do not have any tolerances on the flange outside diameter. The B16.5 If reducer is installed on pump suction-side, it would be an eccentric type with flat side up.

Plant piping side

Pump side

ISO dimensions

Any mismatch in wall thickness to be tapered per API ASME RP 14 E, Fig. B1.1 Dimensions

FIG. 3. The vendor’s scope of supply for a companion flange set.

Interpretation 2-10 ( July 1990) confirms this understanding. This new requirement would introduce a visible difference in flange outer diameter for plant piping flanges (having no negative tolerance) and pump flanges. Although it may seem inconsequential, this raises another concern for noncompliance with plant piping. Discrepancy in large-size flanges.

An additional section, 6.4.2.2.2, is being proposed for the 12th edition. “ANSI/ ASME B16.47 flanges with nominal outside diameter (OD) > 0.125 in. (3.2 mm) shall be approved by the purchaser.” A discrepancy appears to exist here. ANSI/ ASME B16.47 covers large-diameter steel flanges (26 in.–60 in.), an unlikely size range for API 610 pumps. Suggested remedial measures. Com-

panion flange sets are a simple and effective measure to connect an ISO pump to ASME field piping in lieu of drilling holes, which penetrate the pump casing. It is recommended to request that the pump supplier provide a pair of companion flanges

with stud-bolts and gaskets, as shown in FIG. 3. A comparison chart of ASME B16.5 and ISO flange-drilling templates should be prepared by a pump application engineer, with assistance from piping engineers, to ensure if companion flange sets are required. Companion flanges would be necessary if a supplier chose to supply a pump set with API-prohibited sizes. Since they are familiar with site issues, most ISO pump suppliers readily supply companion flange sets. Adherence to ASME standards. It is suggested that API 610 adheres to ASME B16.5 for sizes up to 24 in., and B16.47 for sizes that range from 26 in.–60 in., if needed. Any reference to ISO/EN standards may be eliminated; in the field, piping codes would exclusively be either ISO/EN or ASME. Equipment and piping contractors are never given the choice of either standard. ASME standards are commonly used in the Americas, most of Asia and in the Commonwealth of Nations countries, while European standards prevail in Europe, CIS and North Africa. Separate requirements for heavyduty pumps. Until the 11th edition, API

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610 referred to ISO 9905 for heavy-duty pump applications in industries other than petroleum, petrochemical and gas processing. ISO 9905 provides absolutely no reference to any ASME or API documents. As API intends to break away from ISO, the 12th edition of API 610 will have a dedicated section to cover special-purpose, heavy-duty centrifugal pumps and delete reference to ISO 9905. Therefore, it is logical to dispose of ISO piping and flange standards, as well.

D. GOPALKRISHNA MURTI is a senior consultant affiliated to The Augustus Group in Texas. He has more than 40 years of experience in design and project engineering. He works within the onshore and offshore, LNG/NGL processing, refining, petrochemical and power industries, and has over 30 revisions/ additions/new standards to API, ASTM, ASME, NFPA, BSI, etc., to his credit, all based on field experience. He has authored several articles on design issues concerning the industry that have been published in industry publications. Mr. Murti obtained his BS degree in mechanical engineering from Jiwaji University in India. He is a registered engineer in Canada and India, with licensing in Texas in process.

Viewpoint

REBECCA LIEBERT, PRESIDENT AND CEO Honeywell UOP

New challenges demand innovative solutions

REBECCA LIEBERT is president and chief executive officer of Honeywell UOP, a strategic business unit of Honeywell’s Performance Materials and Technologies division. UOP is a leading international supplier of process technology, catalysts, engineered systems and technical and engineering services to the petroleum refining, petrochemical and gas processing industries. Ms. Liebert holds a BS degree in chemical engineering from the University of Kentucky, a PhD in chemical engineering from Carnegie Mellon University and an MBA from the Kellogg School of Management at Northwestern University.

As the industry works through the most volatile conditions in more than a generation, refiners are facing several challenges that will define the industry in the next decade. The future will belong to a new kind of entrepreneur—one driven by world-class research and development, innovative cost elimination, commercial excellence, and modern control and automation systems. Many state-owned companies are now emulating entrepreneurial upstart companies in other industries. Oil companies are crossing national boundaries, with Middle Eastern companies investing in Indonesia, and Chinese companies moving into Africa.

This trend is re-invigorating and changing trade flows. Evidence is seen in the thousands of miles of new pipelines, massive new ports, import facilities being re-engineered for exports and export facilities being reworked for imports. Every year, demand for refined products—i.e., fuels and petrochemicals— will continue to rise by approximately 1 MMbpd–1.5 MMbpd. We can predict this increase because demand is being driven by growth in the world population, which increases by around 75 MM people each year (FIG. 1). Even faster and equally unabated growth is seen in global standards of living, which are increasingly driven by emerging economies in Asia. Over the next 35 years, the number of vehicles will double to 2 B worldwide. Even with the growth of electric vehicles and renewable energy sources, demand for refined products will continue to rise in the foreseeable future. By the middle of the century, hydrocarbons will still provide three-quarters of the world’s energy. An increasingly prosperous population will consume more plastics, fertilizers, pharmaceuticals, fabrics and other products. All of this added demand will need to be met. However, as capacity grows—powered by exploration, production and more efficient conversion technology— what energy providers produce will change. Until two years ago, most of the world was consumed by the race to make more diesel. Fundamental changes in the speed and nature of economic growth, as well as government economic policies to stimulate consumption, have effectively ended that race for the moment. As lower crude oil prices stimulate consumption, especially among the growing middle class in emerging economies, the race now is about making gasoline. In China alone, a record 25 MM automobiles were sold last year, despite an economic slowdown.

Refiners participating in the gasoline race will reconfigure their plants, while others will simply add capacity. Either way, fuel producers will pursue the twin goals of adding production to meet demand growth and gaining the flexibility to capitalize on products that will generate the greatest profit. Beyond fuels, we are on the cusp of an age where more and more refineries will exist only to produce petrochemicals. Population and GDP growth will drive annual demand growth of more than 5% for olefins and aromatics. As a result, regions of the world that generate these molecules will show more urgent interest in moving downstream, and many refineries that are not yet integrated with petrochemical production will soon be. Another phenomenon fully under development is the tightening of clean fuels regulations. Whether Euro 6, China-5 or BS-6, these standards are becoming stricter as they converge toward a common set of specifications. Removing sulfur and volatile compounds, while enhancing gasoline pool octane, will be a challenge. With all the additional constraints, shifting refinery stream cut points is becoming less of an option. The new standards create opportunities, particularly as new trading patterns for these fuels emerge. In addition to meeting domestic demand, a world-scale refiner might plan to supply a dozen other countries bound to the same clean fuels standards—and do so quite profitably. Finally, while margins will change on every product over time, an ongoing imperative remains to constantly improve efficiency—i.e., the unending campaign to get more out of every drop of oil. A historical constant in our industry is that new challenges inevitably are met by new solutions. The industry is entering another revolutionary period where data can be collected more cost-effectively than ever before. Dramatic declines in the cost of sophisticated sensing devices are possible. We now Hydrocarbon Processing | OCTOBER 2016 31

Viewpoint can measure everything, and what we do with that data is changing how we operate refineries and petrochemical plants. By leveraging the Industrial Internet of Things, refiners can unlock additional

value by integrating smart-edge devices, secure cloud-based infrastructure and operational process knowledge. We can then run refineries with software-based systems that employ big data analytics.

The world population will grow by more than 2 billion people—to almost 10 B.

About $40 T will be spent on the energy infrastructure between now and 2050.

The middle class will grow from 2 B people to 5 B

World GDP will double to about $120 T

2050

World energy consumption will increase by 40%, and 70% of this energy will be provided by fossil fuels

The number of vehicles in the world will double to 2 B.

Fuels will burn more cleanly, and new engine designs will double the fuel efficiency of vehicles.

FIG. 1. With the growth in global population, the world will need a substantial amount of new energy by 2050.

These systems yield smarter plants that are capable of deep self-diagnosis and selflearning. We can sense operating anomalies days, weeks and even months before they become problems. The refinery of the future will solve another looming problem for the industry–the imminent retirement of skilled operators. In the phenomenon known as the “great crew change,” half of the industry’s skilled plant operators will retire in the next seven years. Here, software-based systems can help cover that knowledge gap. Until now, operators could only leverage experience at their own plants, supplemented by their general training. In a smart plant, every operator can benefit from the knowledge gained by every operational challenge experienced by anyone, anywhere, that uses the same or similar processes. The new, more entrepreneurial operator of the future will be diversified into petrochemicals; seize the opportunities provided by clean fuels standards; and count on technology to deliver smarter, higher-performing plants.

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| Special Focus PROCESS CONTROL AND INFORMATION SYSTEMS Advanced process control solutions provide hydrocarbon processing companies with inventive ways to manage difficult and complex operational and reporting applications. The right mix of simulation and automation, combined with talented and creative employees, can solve human resource issues, alleviate equipment malfunctions and failure, and bridge communication gaps. This month’s special report addresses concerns regarding human-machine interfaces, as well as innovations in control devices, hardware and software. Photo: An integrated approach for a plant’s automation and electrical systems brings related data together into a single interface. Photo courtesy of ABB Inc.

Special Focus

Process Control and Information Systems J. BIRD, D. SEILLIER and E. PIAZZA, Valero Energy Corp., San Antonio, Texas

Optimize a CDU using process simulation and statistical modeling methods A methodology was implemented to optimize the operation of a refinery crude distillation unit (CDU) using a combination of process simulation and statistical modeling methods. The primary objective was to estimate a set of operating targets for column pumparound and bottoms stripping steam flows. These targets were established to maximize the unit profitability over a typical range of crude rate and crude quality operating conditions. The crude unit has an advanced process control (APC) application that maximizes product draw rates, but does not optimize the variables above. Process simulation was used to evaluate the CDU performance over a feasible range of pumparound and bottoms stripping flows, as existing operating data did not provide sufficient data. Crude quality and crude feed rate were sampled randomly from actual operating data to account for their inherent process variability. To develop a robust set of operating targets that would perform well under varying market conditions, alternate market scenarios were considered—where gasoline margins exceeded diesel margins, and vice versa—when calculating the unit profit function. Several statistical modeling methods were used to build 3D profit response surfaces as a function of the

operating targets to determine the economic optimum. The estimated optimum operating targets for pumparound and bottoms stripping steam flows are being implemented. Study goals and parameters. A crude distillation unit

takes a crude stream and separates it into boiling point fractions, which include naphtha, kerosine, diesel and tower resid bottoms. A process diagram for a typical crude distillation unit, which has four tower pumparounds, is shown in FIG. 1. Pumparounds remove heat from the column to preheat the incoming crude prior to the crude entering the crude heaters, and to generate internal reflux for distillation.1 The optimum targets for pumparounds and bottoms stripping steam flows depend on the impact of these variables on both product yields and energy use. As the amount of heat removed from the column via pumparounds increases, the heater duty requirements are reduced at the expense of column fractionation efficiency.

Offgas

Naphtha Top P/A Kerosine P/A Steam Kerosine

LGO P/A Steam

HGO P/A

LGO

Crude from heater Steam

HGO

Steam Residue

FIG. 1. A crude distillation unit process flow schematic.

FIG. 2. The strong impact of the diesel pumparound ratio on product yields and on heater duty requirements can be seen in this scatter plot matrix. Hydrocarbon Processing | OCTOBER 2016 35

Process Control and Information Systems This study is based on the use of process simulation to evaluate the performance of the unit over a range of pumparounds and tower bottoms stripping steam flows. Process simulation was selected, as unit operating data did not provide a sufficiently wide range to allow the determination of the optimum targets.2 Pumparound flows were represented as the ratio of the pumparound to the crude flowrate. The bottoms stripping steam flow was represented as the ratio of the pounds of steam per gallon of tower resid bottoms. The process simulation results were used to construct response surfaces using multiple regression methods for product yields and heater duty requirements to validate the simulation results prior to building the profit response surfaces. Profit response surfaces were then built using the predicted product yields, heater duty requirements and product prices for different market scenarios with multiple regression methods. The profit response surface mapped out the

crude distillation unit profit as a function of the pumparound ratios and the tower bottoms stripping steam ratio. The following methodology was used in the study: • Develop a set of simulation cases that covers the range of pumparounds and stripping steam ratios considered • Randomly draw the crude feed composition, as well as the crude feed rate, for each simulation case • Run process simulations for the cases defined above • Use simulation results to build multiple regression models of product yields and crude heater duty requirements as a function of pumparound and stripping steam ratios • Produce 3D response surfaces based on the regression models to map out the product yields and the heater duty requirements as a function of pumparound and stripping steam ratios • Generate profit response surfaces for market conditions where gasoline margins exceed diesel margins, and vice versa • Validate results obtained with multiple linear regression models with those obtained with other statistical modeling methods that model non-linear behavior, including multivariate adaptive regressive splines (MARS) and classification and regression trees (CART). Detailed descriptions of the process simulation and statistical modeling of product yields and heater duty requirements are provided here, followed by the economic optimization analysis and key findings. TABLE 1. Model specifications common to all simulation cases Model specification

Value

Condenser temperature

140°F

Top tray temperature

261°F

Kerosine TBP, 90%

488°F

Diesel TBP, 90%

663°F

FIG. 3. Naphtha product yield contour map.

Heater outlet temperature

630°F

FIG. 4. Kerosine product yield contour map.

FIG. 5. Diesel product yield contour map.

36 OCTOBER 2016 | HydrocarbonProcessing.com

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Process Control and Information Systems

FIG. 8. Product margins and natural gas prices, diesel and gasoline mode. FIG. 6. Resid bottoms product yield contour map.

FIG. 7. Heater duty requirements contour map.

Process simulation. To conduct simulations, proprietary software was selected for a user-friendly spreadsheet interface that provides the capability to run multiple cases.a TABLE 1 summarizes the model specifications common to all of the simulation cases. The process variables that were modified for each simulation include crude feed composition, crude feed rate, pumparound flowrates and bottoms stripping steam mass rate. A total of 60 simulation cases were initially configured to define the range of operations with respect to pumparound flowrates and bottoms stripping steam mass rates. The 60 cases covered pumparound flowrates ranging from 20 Mbpd–50 Mbpd for kerosine, and diesel and bottoms stripping steam mass rates ranging from 6,500 lb/hr–12,500 lb/hr. To model the variability in crude feed composition, three different crude assays were used, corresponding with the re38 OCTOBER 2016 | HydrocarbonProcessing.com

finery’s three most typically run crudes. The percentages of two of the crudes were varied randomly, and the percentage of the third crude was calculated by difference so that the total crude volume percentage added to 100%. Crude feed rates were modeled using a normal distribution, with the mean and standard deviation estimated from operating data. A second set of 60 simulation cases was defined after preliminary analysis indicated that the direction of the optimum was at high diesel pumparound flowrates and high bottoms stripping steam mass rates. The impact of kerosine pumparound was not found to be as significant, so the second set of cases kept the same range of kerosine pumparound flowrates as the initial configuration. The second set of runs covered diesel pumparound flowrates of 40 Mbpd–50 Mbpd, and stripping steam mass rates of 10,500 lb/hr–12,500 lb/hr. Modeling of product yields, heater duty requirements. The impact of tower pumparound and bottoms stripping stream ratios on product yields and heater duty requirements was first assessed to validate the process simulation results prior to building the profit response surfaces. FIG. 2 is a scatter plot matrix illustrating the relationships between product yields and heater duty requirements against tower pumparound ratios, reflux ratio and bottoms stripping steam ratio. The strong impact of the diesel pumparound ratio on product yields and on heater duty requirements can be seen. The impact of the kerosine pumparound ratio was not found to be as significant. A strong correlation between product yields and duty requirements with reflux ratio can also be observed. Tower pumparounds were expressed as a ratio of pumparound flow to crude flow (P/A ratio). Heater duty requirements were expressed as MBtu/bbl of crude. Note that the diesel product yield was positively correlated, and the kerosine product yield negatively correlated with the diesel P/A ratio, as expected. The diesel P/A ratio was also found to be highly correlated with reflux ratio, as the top tray temperature was assumed to be constant and the simulator adjusted the reflux ratio to maintain this temperature. Since the reflux ratio was found to be highly correlated with the diesel P/A ratio, the reflux ratio was excluded as a regressor

Process Control and Information Systems

FIG. 9. Gasoline mode profit/bbl vs. diesel P/A ratio and steam ratio contour map.

to minimize the effects of multi-collinearity on the multiple linear regression models. The tower bottoms stripping steam ratio was found to be highly correlated with the resid bottoms yield, diesel yield and naphtha yield. As expected, the diesel P/A ratio and the heater duty requirements were found to be negatively correlated, as high P/A ratios translate to lower heating requirements. The scatter plot matrix illustrating these relationships was generated using a statistical graphics procedure.b To examine the relationships between product yields and heater duty requirements against the key factors, second-order linear regression models were constructed with both quadratic and interaction terms.3,4 These models were then used to build response surfaces to examine the unit performance over the operating range prior to proceeding with the economic optimization analysis. FIGS. 3, 4, 5 and 6 provide 3D surface contour maps of product yields as a function of bottoms stripping steam ratio and diesel P/A ratio. Note that naphtha yield is maximized at maximum diesel P/A ratio, kerosine yield at minimum diesel P/A ratio, diesel yield at maximum diesel P/A ratio, and resid bottoms yield at minimum diesel P/A ratio. In terms of the bottoms stripping steam ratio, diesel yields were maximized at maximum stripping steam ratio, and resid bottoms yield at minimum steam ratio. FIG. 7 shows that maximum heater duty requirements occur when the diesel P/A ratio is at a minimum, as expected. Economic optimization analysis. Once the process simulation results were validated based on the second-order linear regression model results, profit response surfaces were built to determine optimum targets. Profit response surfaces were constructed for scenarios where gasoline margins exceeded diesel margins, and vice versa, to develop a set of robust targets that would perform well under varying market conditions and minimize the need to adjust these targets. FIG. 8 illustrates the average product margins used to estimate product revenues, as well as the natural gas prices used to estimate the crude heaters fuel costs and stripping steam costs. This data was based on actual pricing data from November 2014 to October 2015. The resid

FIG. 10. Diesel mode profit/bbl vs. diesel P/A ratio and steam ratio contour map.

FIG. 11. MARS gasoline mode profit/bbl vs. diesel P/A ratio and steam ratio contour map.

bottoms product margin was estimated as 70% of the gasoline margin and 30% of the diesel margin. Profit response surfaces were first constructed based on second-order linear regression models. FIGS. 9 and 10 provide the profit per barrel of crude for both market scenarios considered. Note the higher density of points at the higher values of diesel P/A ratio and stripping steam ratio, which represent the second set of simulation runs configured. When gasoline margins exceeded diesel margins, the profit function was maximized at maximum diesel P/A ratio (FIG. 9). The profit response surface was found to be relatively Hydrocarbon Processing | OCTOBER 2016 39

Process Control and Information Systems Yes

Steam P/A ratio < 0.1111

Steam ratio < 0.08714 8.594 n = 14

8.617 n = 12

No

Kerosine P/A ratio < 0.2889 Kerosine P/A ratio < 0.2672

8.624 n = 12

Steam P/A ratio < 0.1446 8.63 n = 13

8.642 n = 16

Kerosine P/A ratio < 0.481 8.63 n=7

Diesel P/A ratio < 0.3145 8.632 n=7

Kerosine P/A ratio < 0.3595

Kerosine P/A ratio < 0.4039 8.64 n = 13

8.649 n = 11

8.646 n = 15

FIG. 14. CART regression tree diesel mode profit/bbl.

FIG. 12. MARS diesel mode profit/bbl vs. diesel P/A ratio and steam ratio contour map.

Yes

Diesel P/A ratio < 0.3706

No

Diesel P/A ratio < 0.1914

12.94 n=9

Diesel P/A ratio < 0.465

Diesel P/A ratio < 0.2758

12.97 n=9

12.98 n = 21

Diesel P/A ratio < 0.417

Steam P/A ratio < 0.1614

12.99 n = 24

13.02 n = 19

13.01 n = 31

13.01 n=7

FIG. 13. CART regression tree gasoline mode profit/bbl.

flat as a function of bottoms stripping steam ratio for this scenario. When diesel margins exceeded gasoline margins, the profit function was maximized at the highest diesel P/A ratio and at the highest bottoms stripping steam ratio (FIG. 10). To validate the results obtained with the multiple linear regression models, a model based on the MARS method, which uses piecewise linear basis functions to allow for the modeling of non-linear behavior, was also constructed. FIGS. 11 and 12 show profit response surfaces based on the MARS method for both market scenarios. Note that the behavior of both profit response surfaces was consistent with the results obtained with the multiple linear regression models. A proprietary procedurec was used to build the model, and two proceduresd were used to generate the profit response surfaces. As an additional verification of the analysis results discussed here, the CART methode was used to map out the unit profitabil40 OCTOBER 2016 | HydrocarbonProcessing.com

ity as a function of the key drivers. The CART method is based on binary recursive partitioning, which also models non-linear behavior. FIGS. 13 and 14 are regression trees predicting unit profitability for both market scenarios considered. Note that profit is maximized in either case at higher diesel P/A ratios. The CART regression tree results show that the range between the maximum and minimum terminal node values was $0.08/bbl when gasoline margins exceed diesel margins, and $0.06/bbl when diesel margins exceeded gasoline margins. Key findings. This work has determined optimum operating

targets for crude distillation unit pumparound flowrates and bottoms stripping steam mass rates using process simulation combined with statistical modeling. Diesel P/A ratio and the bottoms stripping steam ratio were found to be the key drivers impacting unit profitability. The analysis estimated maximum diesel P/A ratio and maximum bottoms stripping steam ratio as the optimum operating targets for the range of market scenarios considered. NOTES KBC’s Petro-SIM 4.1 process simulation software. b SAS PROC SGSCATTER procedure. c SAS PROC ADAPTIVEREG procedure. d SAS PROC TEMPLATE and PROC SGRENDER procedures. e R rpard a

REFERENCES Gary, J. H., G. E. Handwerk and M. J. Kaiser, Petroleum Refining: Technology and Economics, 5th Ed., CRC Press, Boca Raton, Florida, 2007. 2 Montgomery, D. C., E. A. Peck and G. G. Vining, Introduction to Linear Regression Analysis, 5th Ed., John Wiley & Sons Inc., Hoboken, New Jersey, 2012. 3 Montgomery, D. C. and R. H. Myers, Response Surface Methodology: Process and Product in Optimization Using Designed Experiments, 1st Ed., John Wiley & Sons Inc., New York, 1995. 4 Del Castillo, E., Process Optimization—A Statistical Approach, Springer, 2007. 1

JOSE BIRD is director of advanced analytics at Valero Energy Corp. He is responsible for implementing statistical solutions in the areas of process optimization, energy efficiency, process monitoring and ethanol manufacturing operations. DARRYL SEILLIER is a technology advisor at Valero Energy Corp., and is responsible for leading strategic projects and company-wide process improvement in the areas of energy efficiency and hydrogen systems. ERIC PIAZZA is a senior staff refinery models engineer at Valero Energy Corp. He is a subject matter expert in refinery process modeling.

Special Focus

Process Control and Information Systems A. KERN, Lin & Associates, Phoenix, Arizona

The history, and possible future, of model-less multivariable control

Model-less multivariable control. The slow pace of progress has stemmed from industry’s steadfast commitment to the original promise of multivariable control (to solve process control completely), coupled with at least two unanticipated structural limitations:1 • The process disturbances under control often alter the very models used to control them. • Operational precaution normally takes priority over error-minimization performance criteria. These two observations help illuminate the persistent challenges and present a path toward achievement. Where models change and error-minimization is not the main performance priority, detailed models become untenable and unnecessary. Developments such as robustness algorithms and move suppression techniques essentially serve to ignore model detail in favor of more reliable performance. Model details (precise steady-state gains) are also unnecessary to arrive at the correct optimization solution, which is usually wellknown by the operating team in the first place. Often, model

gains are “tuned” to get the desired optimizer result—not the other way around. Industry is aware of these limitations, but perhaps not of their deeper structural significance and the impact on aspects of multivariable control practice, performance and progress. The idea of model-less multivariable control can be unexpected, as industry is accustomed to the terms “model-based” and “multivariable” going together. This article will show that multivariable control can be, and in many ways has always been, model-less. Model-less multivariable control mimics (automates) prior proven manual operation methods, capturing the normal automation benefits of greater consistency and timeliness, as well as the traditional multivariable control benefits of increased capacity, efficiency and quality (FIG. 1). These benefits derive from reliably closing the constraint control and optimization loops, and not necessarily from using models to do so. Small-matrix design. Small-matrix design is another important methodology in which the multivariable controller matrix design primarily includes the variables and interactions (models) that are already being utilized in existing operation to manage process constraints and optimize operation. The purpose of a multivariable controller application is to automate the way the operating team manually manages and optimizes the process (in the absence of, or prior to the deployment of, an automated multivariable controller).

Automatic (closed-loop) multivariable control

Capacity, reliability

Terms such as “model-less,” “small-matrix” and “operational” do not, as yet, evoke the same excitement that accompanied the historic arrival of model-based predictive multivariable control three decades ago. That excitement is now long gone, and an overhaul of industry’s multivariable control paradigm is long overdue. When the overhaul comes, as it must, in response to lessons learned and to meet the needs of modern process plant operation, these terms may well emerge at the heart of the new lexicon. In those heady days when multivariable control was young, it seemed as though it might solve process control altogether, much like GPS technology, which emerged in the same era, would go on to solve navigation. All that remained was to improve the tools, which, in the case of GPS, has certainly happened. However, multivariable process control (MPC) has been slow to improve in the face of decades of relentless experience. The primary area of emphasis—better tools for step-testing, model identification and model-performance monitoring—has been unable to substantially alter the counter-intuitive experience of multivariable control: expected performance often remains elusive, rather than reliable. This historical interpretation will reveal why progress has always been slow, where the path forward now leads, and how it all fits together with many lessons from historical process control experience (both single-loop and multivariable).

Manual (open-loop) multivariable control

Process

constrain

t limits Efficiency, quality

FIG. 1. Model-less multivariable control mimics (automates) priorproven manual operation methods, capturing dual benefits of greater consistency and timeliness, and increased capacity, efficiency and quality. Hydrocarbon Processing | OCTOBER 2016 41

Process Control and Information Systems Small-matrix design results in a handful of variables and models vs. the hundreds that typically result from traditional big-matrix design (TABLE 1). It is assumed that the extra models netted by big-matrix practice contribute to a more complete solution. However, these models often lead to well-known (but still poorly understood) “degraded” MPC performance, and can also make the finished controller much larger and more difficult to own and operate. MPC engineers often instinctively prune matrices to increase controller operability, manageability and reliability. Few engineers have realized that the sensible conclusion to this trend is small-matrix design practice, which utilizes existing proven operation as the matrix design basis in the first place. History repeats itself, and necessity fosters invention. Observing the parallel experience of modern multivariable control modeling and historical single-loop tuning is compelling. The two activities (tuning and modeling) are fundamentally the same—measuring process response to derive controlTABLE 1. Summary of traditional big-matrix practice vs. small-matrix design practice Big-matrix design practice

Small-matrix design practice

Double-digit matrix dimensions, (e.g., 20 × 50)

Typically single-digit matrix dimensions (e.g., 6 × 8)

Hundreds of potential variable and models

One or two dozen potential models

Large and complex to own and operate

Intuitive to own and operate

Based on identifying all process interactions; in operation, many prove unwanted for a variety of reasons

Based on the variables and interactions used in existing operation to manage constraints and optimize the process

Directional move-solver

Preselected move rates

RBC rate-based control

DCS execution period ca. 5 sec.

ler settings. While both activities should be reliable one-time tasks, in practice they remain characterized by under-performance and chronic rework. An examination of common root causes points to MPC’s structural limitations and the realization that they would have (in hindsight, plainly have had) the same effect on single-loop practice. This makes the path forward clear: If industry is to achieve more reliable multivariable and single-loop control performance, it will need a control method that addresses these root causes, is more robust with regard to changing process gains, and delivers operational performance. This line of thought is only pursued reluctantly, as it appears to abandon the hope of solving process control altogether, and further complicates a technology that already suffers from unwieldiness. However, promising initial discoveries have been made, revealing a solution with the potential to overcome past structural limitations and simplify practice. The model-less multivariable control method. These experiences compelled the development of a model-less multivariable control method (XMC), shown in FIG. 2, which conceptually comprises three parts: • A logic-based directional move solver. Since it determines only move direction, the solver requires only gain direction and not detailed models. • Pre-engineered move rates. Like traveling by automobile, speed is based on reaching the destination safely (operational performance), rather than on the remaining distance or minimizing travel time (error-minimization). • Rate-based control (RBC). This technique tapers, reduces and halts moves predictively, so that when the controlled variable prediction equals the target, the direct control variable (DCV) moves are halted. This results in the indirect control variable (ICV) ultimately settling on the constraint limit or optimization target without overshoot or oscillation (operational performance), based on first-order process response mathematics and dynamics. FIG. 3 shows that RBC is inherently adaptive to changes in process gain. For example, if process gain increases, then actual

FIG. 2. The three components of a conceptualized model-less multivariable control method. 50 45

Target change

Change, %

RBC 40 35 30 25 -10

Target DCV, MV ICV, CV Prediction

Pre-selected move 0

10

20

Time, min.

30

40

FIG. 3. The RBC methodology tapers moves predictively.

42 OCTOBER 2016 | HydrocarbonProcessing.com

50

60

FIG. 4. A hydrotreater XMC application at a US refinery has been online for more than a year. Benefits include better sulfur control and increased product value from optimized hydrogen uptake.

process response will increase, and RBC will correspondingly taper the moves sooner. The same capability applies to changes in the predefined move rate, meaning that move rates can be adjusted to achieve desired operational performance without impacting control performance. This key component of the model-less method also provides industry with an inherently adaptive control algorithm. Interestingly, the “self-tuning controller” era came and went without discovering this inherently adaptive method. A prime strength of XMC is not its mathematical ingenuity, but is instead the creative combination of its three novel components and key historical learnings. History’s full circle. Successful operating teams have always relied upon a robust working knowledge of process interactions and an appreciation for the dynamic nature of process behavior. They rely on proven variables to manage constraints, optimize economics and avoid variables that have proven to be unreliable, unwanted or risky. Teams operate in safe steps to avoid overshoot or oscillation, which can cause or mask process instabilities. The XMC model-less multivariable control method mimics these traditional methods to capture the usual automation benefits of greater consistency, timeliness and reliability, as well as the traditional multivariable control benefits of increased capacity, efficiency and quality. Future prospects. A successful model-less multivariable con-

trol method has far-reaching implications by simplifying ownership at nearly every lifecycle stage and eliminating several stages and costs completely. For example, a prototype XMC controller has been developed and deployed natively on a standard distributed control system (DCS) platform. In place of a design project, the controller was fully designed in a single meeting. Deployment was guided by a routine management of change (MoC) checklist (FIG. 4), and ongoing support is provided by in-house DCS engineers. The model-less multivariable control technology is now in use at multiple industrial sites, and has operated continuously and successfully (in some cases) for more than 1 yr. Operator acceptance and uptime have been high. Benefits for a hydrocracker application include improved sulfur control, increased product value (through optimized hydrogen uptake), smoother crude switches, and the elimination of 90% of manual operator moves, freeing up time for other tasks and priorities. The prospect of automated multivariable control becoming a core competency for the process industries is real, as multivariable constraint control and optimization is an inherent aspect of essentially every process operation. 1

The Quiet Work-Horse!

Process Control

LITERATURE CITED Kern, A., “Take the path to model-less multivariable control,” Control Magazine, December, 2015.

ALLAN KERN is a control engineering consultant with Lin & Associates Inc., where he is responsible for advanced process control, including XMC. He has 35 years of industrial process control experience and has authored numerous papers on a wide range of practical process automation solutions. He earned a BS degree in chemical engineering from the University of Wyoming, and holds professional engineering licenses in chemical engineering and control systems engineering.

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Special Focus

Process Control and Information Systems S. HALL, Process Systems Enterprise, UK; and P. MAHONEY, Ambition Partner, UK

Use process knowledge management systems to accelerate innovation The amount of stored digital data is doubling every two years, and the rate of data acquisition is close to exceeding our ability to process it and sort out what is valuable. In this scramble for information control, organizations are turning to knowledge management (KM) systems to retain their know-how and capture knowledge capital that has commercial value. Companies hope to promote and accelerate the further innovation that is required to maintain their market positions. KM is adopted by organizations to identify, capture, evaluate, retrieve and share knowledge assets, which can be documents, policies, procedures, databases and working software, among others. KM also includes capturing and harnessing personnel expertise, knowledge and experience. In the process engineering community, this encompasses engineering work processes, design procedures, standards, guides and the knowledge of individual engineers, who represent valuable knowledge capital, often proprietary in nature, that may have been acquired over a working lifetime. To understand the subject and define the needs, we can consider knowledge in a number of different forms: • Explicit: Information or knowledge that is set out in tangible form • Implicit: Information or knowledge that is not set out in tangible form, but could be made explicit • Tacit: Information or knowledge that is extremely difficult operationally to set out in tangible form, such as experience and personal context. From these definitions, it can be seen that the challenge of a KM system is not just storage and easy retrieval of information. It also involves tangible knowledge

definition. This article discusses the role of KM systems within process engineering, design and development environments. Challenges of a dynamic industry. The process engineering sector is facing a number of critical challenges. The first is the large volumes of design and operating data being collected. Industry thrives on this data, and work practices have been adapted to manage and present the large volumes to users through analysis and visualization technologies. But how do we break through the mass of stored information to get at the actual knowledge within? Another challenge is the dynamic nature of our industry. The rapidly changing commercial environment has been demonstrated over recent years with the rapid growth in shale gas production and the fall in crude oil prices. To address the changing market focus, we see a significant movement of people between organizations, more outsourcing and frequent corporate restructuring, none of which promote the retention of expertise or the preservation of a company’s intellectual property and consequential market advantage. An organization’s knowledge capital, which resides within the workforce itself, may ebb and flow with staff changes. A senior engineer leaving or retiring from a company will likely take with them a lifetime of proprietary knowledge. It is vital to capture this knowledge before they leave, and equally important to secure internal knowledge as it is created. Information storage is another KM challenge. Knowledge embedded in tools, documents, working practices and databases becomes ineffective unless it can be easily accessed—either passively,

through enhanced search capabilities; or actively, by pushing the information to the right stakeholders. This informationpush capability is discussed here. Home, remote and virtual workplaces are becoming more common and efficient, but how do we replicate the transfer of knowledge that goes on, for example, around the morning coffee pot? How do we discern exactly who can most effectively answer our questions? KM systems have made significant inroads in addressing many of these issues, but the process engineering community continually places more requirements on such systems. Unlocking, harnessing and proactively delivering the knowledge that resides within an organization’s processes, systems and people is the new focus. These activities enable an organization to maximize its productivity and retain its competitive advantage. Selecting the correct approach. KM delivers value through data capture and sharing across the business, requiring the organization to work from top to bottom, as well as horizontally, in a structured way. This discipline is a factor in determining the rate at which KM can effectively be deployed. The steps for KM implementation are the same as for any business project, with particular attention paid to organizational readiness. Internal inconsistencies in how data is collected and described often exist. These differences will need to be resolved to achieve the full investment value with a plan for the ongoing development, management and support of the KM. As knowledge sharing and system management are achieved, opportunities to derive added value through enhanced collaboration will develop. Hydrocarbon Processing | OCTOBER 2016 45

Process Control and Information Systems System architecture. KM systems have

become fully embedded into work processes in other industries, such as research data management (RDM).1 Here, KM captures and extracts knowledge from laboratory data, harnessing individual insights and conclusions, proactively sharing the results and promoting a working collaboration. This model lends itself very well to the process engineering field. Essential features of a fit-for-purpose system (FIG. 1) include: • An enterprise database management system in which information is stored, not only as text or values, but with attributes so that the system knows whether a value is, for example, a temperature or pressure. This adds context and boosts the power of any search. • A user interface that reflects the needs of the user community— for example, the ability to upload engineering information, either

from process simulation tools or from computer-aided design packages, and succinctly present the information to the viewer. Much can be learned here from social media, and a good interface will provide the user with the “feel-good” experience necessary to keep their attention and continued use. • A search capability that understands both content and context. A typical enterprise search engine delivers the functionality found in popular web browsers, where underlying algorithms evaluate why particular fields are searched and report preferred results accordingly. Search results presented as simulation files and flow diagrams are possible. • An information-sharing platform, typically via an intranet, avoids the installation of software locally and allows direct integration to the web.

FIG. 1. Essential features of a process knowledge management system.

FIG. 2. A KM system showing a flow diagram and process simulation input summary.

46 OCTOBER 2016 | HydrocarbonProcessing.com

KM SYSTEMS FOR PROCESS ENGINEERS Common

language. From a global chemical and process engineering viewpoint, a common and consistent language is a primary KM system requirement. Expecting everyone within an organization to speak one language is unrealistic, so the system should have translation capabilities.

Information security. Preserving confidentiality and ensuring that information is accessible only to authorized personnel is crucial. Multi-layered authorization and access privileges are possible, not only to work areas and/or specific subjects, but also access levels for different file types. KM systems are good repositories for design standards, operating data from the field and details of licensed processes, as the information within such documentation can be enhanced with knowledge elements. Information integrity, where the accuracy and completeness of information and processing methods are preserved, is a key part of security. Some documents, such as process flow diagrams with heat and material balances, may need to be linked. Typically, each information item has a unique reference ID (beyond its saved title) that identifies it and its revision status, assigns a discrete position within the enterprise database, binds information items together and tracks what has been approved and what has not. The ability to recover inadvertently deleted or reassigned information is another aspect of any KM system. This recovery from earlier revisions is a typical role of administrator users. Information availability ensures that authorized users have access to information when required. Standard access will likely be granted from desktop computers, but if engineers are working in the field, access may need to be extended to mobile devices. This presents additional protection challenges from malicious cyber attacks that can cause widespread, irreparable financial loss and reputation damage. A structured information security risk assessment from which specific controls can be implemented is mandatory. International standard ISO/IEC 27002:2013 defines the range of controls in various areas and provides best-practice recommendations for initiating, implementing

Process Control and Information Systems or maintaining information security management systems (ISMS).2 Information storage. Storage of key

information applies not only to standard document types—PDF, Microsoft Word, Excel, picture and video files—but also to engineering files, drawings, process simulation and other design files. KM systems have drag-and-drop capabilities, and most store information in a database system, either in the cloud or on an internal server. Interrogator apps based on existing industry platforms have been developed that glean the information contained within process simulation input or results files and store the information within the KM system as structured data, allowing it to be searched. Process simulation tools are migrating to run on cloud servers, and KM systems must adapt so that users can retrieve their input files from the KM system and then run these files in the cloud, with results visible to all, as shown in FIG. 2. Avoiding duplication of data storage is important. A KM system should be a repository of knowledge that does not dupli-

cate data storage. Data historians are one example: a KM system should link to such data wherever possible, but not store such data itself. Workflows. Documents can be checked and approved by authorized users, and their sign-off statuses with date/time stamps are clearly seen (FIG. 3), accelerating how information is used. Approved documents appear immediately, and bulletins can be automatically issued to specific stakeholders. On fast-track projects, preliminary data that has been checked but is awaiting approval may alert engineers and allow follow-up work to start, expediting project development. Information supplied by third parties can also be uploaded, quickly reviewed and viewed by all approved users. User views. KM systems offer intuitive and informative “document view” capabilities so that users can dictate what viewers see. This could be a file icon or, more often, the title page or important content, allowing a user to scan the page for required information.

Search capabilities. Given that a KM

system will store hundreds of thousands of data items, search options are critical. Text searching is standard, but contextsensitive search is required for KM systems to achieve their potential. For example, the system should define what values are temperatures, pressures, etc., but it should also be able to search based on non-absolute information, such as “find all designs with a feed flow greater than 125 Mkg/h and a reactor temperature > 120°C,” (FIG. 4). Searches must interrogate engineering files, simulation files, process flow diagrams, and piping and instrumentation diagrams. Comments and tags. Applied to data

items to provide additional context, comments and tags may be trended to highlight popular topics to the user community and encourage involvement and debate.

Information push. Information and knowledge should be proactively distributed to key personnel. The system should watch for specific words or phrases as documents are uploaded, either in the main

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Contributing Analyst ARC Advisory Group

Unified Engineering—An Analyst Perspective on the Current and Future Direction of Simulation Technologies With the volatility of oil prices and much trepidation, many industrial process manufacturers are unwilling to invest in new capital intensive projects and facilities. The pendulum has swung toward optimization of existing facilities and “sweating the current assets”. New approaches, tools, expectations, and platforms are paving the way for the next-generation of process engineers. ARC Advisory Group will share current process simulation market dynamics, trends and drivers impacting workers, engineers, and operations.

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Process Control and Information Systems document or in tags and comments, and specific individuals should be notified, either by email or though the KM portal. This ensures that subject experts and approved users are aware of all internal activity in their area of responsibility. Working offline. KM systems are cloudbased, so any break in connection can render the system useless for field operatives. Working offline may be required, particularly for engineers in the field or remote locations, and it presents additional security and compliance issues. Working offline should be applied selectively to small segments of the information base.

The objective is to enable a user to synchronize part of the knowledge base with their own computer or tablet prior to the connection being severed. The user then accesses the site and continues to develop their own work within their offline version of the KM system. Once a connection is restored, the user’s computer automatically synchronizes, following typical rules of synchronization and checking each information item for changes. Extending the KM system. To take advantage of a single information environment and simplify the user experience, many KM systems have been extended

FIG. 3. Visualization of file content and sign-offs.

FIG. 4. The structured search capability recognizes context, and must interrogate engineering files, simulation files, process flow diagrams, and piping and instrumentation diagrams.

48 OCTOBER 2016 | HydrocarbonProcessing.com

to include a range of additional capabilities previously provided by other tools. Scientific papers and reference materials may be stored as a documents library. An expert locator allowing individual subject experts or regular staff to upload a simple resume can be included. This information becomes searchable by other staff, helping to bring together the right expertise for problem-solving and innovation. Discussion groups and “communities of practice” can be created where questions are asked and collaboration is encouraged. This is especially relevant as the increase in working from home reduces natural knowledge sharing. The intent is to virtually replicate the social spaces usually found at work. Lessons can be learned here from social media, where information-sharing and communication practices are becoming more advanced and integrated within our virtual lives. Externalization, outsourcing and collaboration are core business and development models for growing companies. A natural barrier to sharing information with third parties will always exist, and this can have a significant effect on the quality of information, the speed of a project and how information is shared. The standard vehicle for information sharing of PDF, Microsoft Word and Excel files is often SharePoint or Dropbox. However, such documents will contain results, not content or context. One way around this is to set up collaborative cloud environments.3 Tools provided by KM companies to map data and create a virtual process development environment recognize scientific content, workflows and the ability to publish information in the existing KM system. Application examples. A KM system adds real value for the process engineering community in process development, where projects require high degrees of innovation and may comprise teams based in different geographic locations, often with third-party suppliers involved, all working to tight deadlines. The amount of parallel work is significant, with a strong dependency of tasks between the various stakeholders. In this situation, using a KM system provides a number of advantages: • The ability to define a common work area to store all documents. User access should be defined so

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Process Control and Information Systems

FIG. 5. Embedding simulation tools within a KM system.

•

•

•

•

that confidentiality is preserved where needed. Sub-contractor areas can be set up to allow external access to all or just selected parts of the system. Any new information or documents can immediately be seen by all stakeholders, including the latest version of key files. This avoids the problem of multiple versions of the same file being distributed, and ensures that only one master version exists. The information should be stored with context added and should be tagged and commented, as needed, to provide a rich information landscape that can be searched and made available to other members of the process development team. On fast-track projects, a document that has been checked but not approved could be used to begin conceptualizing the next stage of development, thereby saving time. If details have been entered into the KM system, then subject experts can easily be located and informed of progress as the development project proceeds. Project managers can see individual activities and what is trending. All reports, designs, etc., are saved and can then be searched and referenced not only during the current project, but also in the future.

50 OCTOBER 2016 | HydrocarbonProcessing.com

Another application is the use of a KM system to manage simulation and optimization models for oil refinery steam and power systems. The working model is a steam/power system running in a simulation environment (FIG. 5). A site runs the optimization model daily to determine optimal operating conditions. However, other onsite staff needed access to the results of the model, both to approve the outcomes for plant implementation and to view the associated key performance indicators (KPIs). Central engineering, located in another part of the country, also needed to review the KPIs to ensure that the refinery was operating at optimal conditions. The model is stored and run from within the KM system itself, allowing all parties to view the same model and output via a web interface. The ongoing performance of the refinery is tracked over an extended period. Remaining challenges. Addressing company culture is one of the biggest KM system implementation requirements. The entire company must implement KM as a core activity in its mission statement and work practices. To realize the true potential of a KM system, key activities must focus on applying KM principles. Support to staff can be through training, ongoing workshops and competency management. Any deviations from the plan during implementation must be documented. Too often, project teams are disbanded and the team members reassigned before

any post-project debriefing. Organizations should counter this and develop a postproject action procedure that provides a forum to learn and improve, as necessary. Another challenge is converting tacit knowledge into internal knowledge—How can knowledge be extracted from a 40-year veteran who is about to retire? This necessitates proactive planning, and the approach will likely depend on both the level of documentation and the personalities involved. A structured approach to manage such knowledge transfers is required. Finally, KM must be fully integrated into the company’s everyday activities and culture. The value of knowledge sharing should be appreciated, recognizing that it does have sustainable bottom-line value for the future. The information target is constantly developing, as are business objectives, but a dynamic KM culture will ensure that the company and its people adapt quickly, while protecting their information and intellectual assets. ACKNOWLEDGMENTS This article was written while the authors worked at Nova Process Ltd., which has subsequently been absorbed into Process Systems Enterprise. The authors would like to thank ID Business Solutions for its support in the development of the KM system. LITERATURE CITED Complete literature cited available at HydrocarbonProcessing.com. STEVE HALL is the director of engineering solutions at Process Systems Enterprise (PSE). Previously, he served as CEO of Nova Process. He has 25 years of experience in process design, innovation and improvement across the oil and gas, petrochemicals and chemicals industries. Dr. Hall is involved in matching PSE technologies to industry needs and managing specific technology development initiatives. He holds a PhD from the University of Manchester Institute of Science and Technology, and is a chartered engineer and a fellow of the Institution of Chemical Engineers. Dr. Hall has presented papers in heat integration and utility system optimization, and coauthored patents in dehydrogenation and hydrocracking. PHIL MAHONEY is chairman of Ambition Partner Ltd. and Nova Process. He has 30 years of experience in the chemical and life sciences industries, and has been involved with process simulation, optimization and R&D data management. He began his career as a process and control engineer at ExxonMobil’s Fawley refinery in the UK, and later became managing director of Aspen Technology. Mr. Mahoney then joined ID Business Solutions in the area of life sciences. He earned a BS degree in chemical engineering from Loughborough University, and is a chartered engineer and a member of the Institution of Chemical Engineers.

Special Focus

Process Control and Information Systems D. WILLIAMS, VEGA Americas Inc., Cincinnati, Ohio

Measure naturally occurring radioactive material in polypropylene plants In the petrochemical industry, process plants often rely on radiation-based level and density detectors. Various vessels within polypropylene plants—reactors, product chambers, flash drums, and purge bins—utilize radiation-based detectors. However, naturally occurring radioactive material (NORM) compromises the accuracy of the level measurement in these vessels. NORM exists in oil and gas in the form of radon and its decay daughters. NORM amount varies depending on where the hydrocarbons are extracted and the length of time since the extraction. Some locations have a higher concentration of NORM in petrochemicals than others. During petrochemical processing, NORM can travel to different portions of the plant with various products, including propane. Feedstock for polypropylene comes from various locations, including fluid catalytic cracking units, ethane crackers and propane dehydrogenation plants (PDH). PDH units account for most NORM problems faced by plant operators. If a polypropylene plant obtains feedstock from PDH units, difficulties with NORM should be anticipated. As more PDH plants are built to supply on-demand propylene to polypropylene plants, the number of issues with NORM affecting nuclear measurement will increase. The effects of NORM. As propylene

enters a polypropylene plant, the NORM that follows can have a detrimental effect on radiation-based control systems. Level detectors will give low-level readings, potentially causing vessels to overfill. The drop-off in level output can be severe, dramatic and dangerous. See FIG. 1, a representation of a reading a customer shared with us, for an example of how quickly NORM can change a level reading.

FIG. 1. The primary level, represented by a red line, drops as the NORM level, represented by a black line, skyrockets.

Once NORM enters a polypropylene plant, it can accumulate in different vessels, especially in vessels with a gas-phase recirculation system. As fresh feed comes into the vessel and the gas is recycled back through the process, the concentration of NORM builds up inside the vessel. The amount of NORM can vary greatly from the bottom to the top of a vessel. This difference is caused by many factors, such as the amount of gas present and the size and shape of the catalyst or product. To understand the effects of NORM on radiation-based level, a basic understanding of how nuclear level measurement works is needed. Three main items are required to have a nuclear level measurement system (FIG. 2): 1. Vessel 2. Radioactive source 3. Nuclear detector. The nuclear level system works by measuring the amount of radiation reach-

FIG. 2. If NORM is present inside a vessel with a nuclear measurement system, the total amount of radiation to the detector increases.

ing the detector from the source. Gamma radiation is converted into a pulse of light using a scintillating crystal. Hydrocarbon Processing | OCTOBER 2016 51

Process Control and Information Systems

FIG. 3. The three main items required for a nuclear level measurement system are a vessel, a radioactive source and a nuclear detector.

The photomultiplier records the number of pulses of light as counts per second, commonly known as count rate. The lower the level, the more radiation reaches the detector, resulting in a higher count rate. The higher the level, the less radiation reaches the detector, resulting in a lower count rate. As the level in a vessel changes, so does the detected amount of radiation. The detector converts this change in radiation into a level signal. When NORM is present inside a vessel with a nuclear measurement system, the total amount of radiation to the detector increases (FIG. 3). As the product level increases, the amount of radiation from

FIG. 4. As NORM increases, detection of level decreases. In this representation, multiple level and NORM detectors are in use. The NORM detectors were not mounted on the vessel.

FIG. 5. NORM level can change rapidly, causing problems with level detection.

52 OCTOBER 2016 | HydrocarbonProcessing.com

the source decreases and the count rate drops. However, as the amount of NORM increases, the count rate does as well, and the detector infers a lower level than actually exists in the vessel. Limiting NORM effects. NORM is a

serious problem for plant operators, but ways exist to limit its influence on level measurement. Common ways to address the issue of NORM in a petrochemical plant include implementing an early warning system; increasing the size of the radioactive source; and measuring the NORM, either directly or indirectly, and subtracting it from the level measurement. A nuclear density meter can be mounted on a feed pipe to create an early warning system to alert operators to the existence of NORM. With no accompanying radioactive source, the density meter does not measure the propylene density, but monitors the concentration of NORM within the propylene. If placed in the proper position on the feed line, the unit will give operators time to switch the radiation-based levels to manual control to prevent an overfill. Early warning systems help operators compensate for NORM before it reaches vessels, but they must act quickly. FIG. 4 illustrates how level drops as NORM increases in an application, and FIG. 5 shows how quickly NORM level can change. In FIGS. 4 and 5, multiple level and NORM detectors are present. Increasing measurement source activity decreases the signal-to-noise ratio, so the detector treats the NORM as noise. This approach is effective in applications where the amount of NORM is low and tends to stay low. Most nuclear measurement applications are sized for a radiation field between 1.0µSv/hr and 0.3µSv/hr (0.1mR/hr and 0.03mR/hr) to give a reliable level output with 1% noise. If the level systems read a NORM level of 0.3µSv/ hr (0.03mR/hr) and the application was sized for 1.0µSv/hr, then the NORM will cause the level measurement to read 30% lower than the actual level. A frightening scenario occurs if an application is sized for 0.3µSv/hr. The system could read 0% when the vessel is actually 100% full. To minimize the effect of NORM on level measurement, operators must increase the size of the source so that the amount of NORM present has a minimal effect on the level output. Eduardo Scarnichia describes a case1 in which a poly-

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Process Control and Information Systems 30 25

Error, %

20

0.3μ Sv/hr NORM

15 10 5 0

1

2

4

6

8 Factor of increase

10

12

14

16

FIG. 6. Operators must increase the source size to minimize the NORM effect on level output.

FIG. 8. Mounting a detector (of equal length to the level measurement detector) 90° around the vessel from the level detector allows the operator to measure NORM directly.

having a means of effectively chopping the radiation beam (by using revolving rods, for example) to alter the radiation measured by the detector. Electronics inside the detector are set to run an algorithm that can effectively ignore the radiation from NORM. In applications with multiple sources, plants are required to install multiple modulators and an electronic synchronizer to keep the mechanical devices moving at the same time, and they must increase the size of the source due to some signal attenuation from the rotating rods. This adds a layer of complexity and effort to the installation of a NORM compensation system and forces plants to use resources to power the modulators. FIG. 7. The primary level, represented by a red line, remains relatively constant even as the NORM level, represented by a black line, experiences dramatic shifts.

propylene plant was forced to increase its source size by a factor of 16 to compensate for NORM. In general, the size of the source must be increased by a factor of 10 to reduce the error to 3%, and by a factor of 15 to reduce the error to 2%, in an application sized for 1.0µSv/hr (FIG. 6). Another way of addressing the issue of NORM when using nuclear level is to measure it directly. Since the presence and concentration of NORM is not constant and varies greatly, depending on the propylene origins and concentration of NORM inside the vessel, the nuclear level has to compensate for this change of radiation. By placing a detector of equal length to the level measurement detector and mounting it 90° around the vessel, 54 OCTOBER 2016 | HydrocarbonProcessing.com

operators can measure NORM directly (FIG. 8). The NORM-compensation detector measures the amount of NORM inside the vessel and sends a signal to the primary level detector. The primary detector runs an algorithm to subtract the effect of the material’s NORM radiation from the source holder’s radiation. Operators must use a detector of equal length to measure NORM, since the NORM level in the vessel can vary from top to bottom due to product distribution, shape and size of product or catalyst. FIG. 7 is a representation illustrating how direct NORM compensation stabilizes a level reading. Plants can also use an indirect method to address NORM, such as installing a mechanical device in front of the source and

Takeaway. Petrochemical plant operators

must have a plan for NORM compensation in their process vessels. To protect employees and the bottom line, plant operators should adopt a NORM compensation system that best fits their process. 1

LITERATURE CITED Scarnichia, E., A. Etchepareborda and M. Arribere, “Radon in propylene: Unexpected influence of NORM in a chemical plant,” Proceedings from IRPA 12: International Congress of the International Radiation Protection Association, Buenos Aires, Argentina, October 2008.

DAVID WILLIAMS is a senior applications engineer for VEGA Americas and has been with VEGA for 14 years. In the business development group, he works with process technology providers to improve level and density measurements in challenging applications, such as resid hydrocrackers, delayed cokers, solvent deasphalting and LDPE high-pressure separators. He holds a BS degree in nuclear engineering technology.

ADVANCES IN SULFUR MANAGEMENT

2016 Special Supplement to

Stricter standards in sulfur treating S–56

Low-sulfur projects dominate the downstream construction landscape S–64

CORPORATE PROFILES CB&I S–61

Enersul S–63 OHL Gutermuth S–67

Paqell S–69

COVER PHOTO A conveyor moves Enersul GX premium formed sulfur granules at the Kaybob sulfur processing and handling facility in Alberta, Canada. Photo courtesy of Enersul.

ADVANCES IN SULFUR MANAGEMENT

STRICTER STANDARDS IN SULFUR TREATING BOB ANDREW, Technical Editor

A number of technical innovations have advanced sulfur removal technology in refineries. Stringent vehicle fuel sulfur specifications are moving beyond North America and Europe and into other developed economies (TABLES 1 and 2, FIGS. 1 and 2). The displayed data shows national regulations only: several countries around the world, including Brazil, China and India, also have sub-national regulations requiring higher-quality fuels in key cities and regions.1 Around the world, legislation mandating decreased emissions and lower levels of airborne pollutants is coming into effect. In response, refiners are implementing operational and processing changes to reduce sulfur levels in transportation fuels. New technologies are moving the downstream hydrocarbon industry toward cleaner, lower-sulfur transportation fuels. A low-sulfur world does not come cheap, however. Refiners are investing billions of dollars in new units, upgrades/ retrofits and expansions to meet new sulfur and emissions regulations. These investments will help produce high-quality fuels that meet Euro 4, Euro 5 and Euro 6 specifications. These standards promote the reduction of carbon monoxide, nitrogen oxide (NOx ), hydrocarbons

and particulate matter in both diesel and gasoline passenger vehicles. Many nations around the world already produce transportation fuels that meet Euro 4 specifications. Other regions, such as the Middle East, are investing heavily to increase the production of Euro 4 and Euro 5 standard fuels. Major clean fuel initiatives include: US/Canada. Both countries’ governments will begin to enforce the new Tier 3 program starting in 2017. This program will set new vehicle emissions standards and lower the sulfur content in gasoline to 10 ppm. The program maintains the existing refinery gate per-gallon content of 80 ppm and the 95-ppm downstream distribution cap. The US EPA forecasts that the new rule will reduce NOx emissions by about 260 Mt in 2018 alone. Large US refineries (those producing more than 75 Mbpd) must comply with Tier 3 standards by 2017. Smaller refiners must meet Tier 3 standards by 2020. China. The country is implementing its National V fuel quality standard, which equates to Euro 5 standard transportation fuels. Euro 5 standard fuels will be required for the automotive industry by 2017. These new regulations are being implemented 1 yr ahead of schedule. Upgrading the nation’s fuel quality could cost

FIG. 1. Map of global national diesel sulfur limits, 2014. Source: International Council

on Clean Transportation.

S–56

ADVANCES IN SULFUR MANAGEMENT | OCTOBER 2016 | HydrocarbonProcessing.com

Chinese refiners more than $7 B. China is also writing new regulations for National VI. No time table has been announced for the implementation of National VI, but if adopted, it would take effect no earlier than 2020. India. The country will implement its Bharat Stage 6 (BS-6) standards by 2Q 2020, four years ahead of schedule. BS-6 fuels are equivalent to Euro 6 fuel specifications. These new regulations call for a 68% reduction in NOx emissions. India’s new regulations will bypass the BS-5 stage and move directly to BS-6. Middle East. Saudi Arabia and Kuwait are leading the charge in new clean fuels projects in the region. Saudi Arabia is investing billions of dollars to reduce sulfur content in diesel and gasoline to 10 ppm, and to lower benzene content in gasoline to 1%. Kuwait is investing more than $30 B on ambitious plans to overhaul its refining sector and become the region’s cleans fuels leader. This plan focuses on integrating the Mina Abdullah and Mina Al-Ahmadi refineries, as well as the construction of the 615-Mbpd Al-Zour refinery. Russia. The country has lacked the advanced facilities to produce highergrade transportation fuels. In response, Russia launched a $55-B program in 2011 to modernize its existing plants and encourage exports of high-quality products. The plan called for the installation of 130 new units by 2020. The modernization program will continue to focus on increasing its light products yields, with a key goal of meeting demand for gasoline and jet fuel, increasing fuel standards to Euro 5 specifications, and replacing old units to decrease residual product yields and maximize utilization. Adding to the international regulatory pressure to remove sulfur from fuels, the International Maritime Organization (IMO) has announced new regulations2 to reduce sulfur in shipping fuels from 3.5% to 0.5% by 2020. Commenting on the IMO’s new sulfur regulations, Adrian Tolson, senior partner at 20|20 Energy, said, “The transition to distillates is going to be messy…in 2020, or possibly 2025, IMO

ADVANCES IN SULFUR MANAGEMENT catalytic packing to concurrently carry out the reaction and fractionate the reaction mixture.” The principal application is in conjunction with fluid catalytic cracking, where it has been applied to the production of low-sulfur gasoline4 and low-sulfur diesel,5 with minimal olefin loss.

regulations will sideline intermediate fuel oil (IFO) as a bunker fuel, destroying a major outlet for refiners’ residual product. A switch by the world fleet to 0.5% distillate fuel would see global fuel oil demand collapse. Instead, ships will be expected to burn bunker with a sulfur content of no more than 0.5%. Only vessels equipped with emissions abatement technology will be able to use high-sulfur fuel oil.”3 Projects and technologies chosen for this supplement were often a “first-of-their kind.” The Hydrocarbon Processing editorial team chose these examples to showcase increased performance, versatility and reliability. These selections feature technologies that can handle sulfur from a wide variety of in-plant refinery process streams, or have selectivity to meet targets for key markets. The choice of a sulfur treating system is also dictated by the ultimate disposition of the sulfur. Production of bulk sulfur is dominated by product extracted from sour gas, but other options to create value include converting sulfur into sulfuric acid.

Research & Licensing Co., this process was described as “... a method for conducting chemical reactions and fractionation of the reaction mixture comprising feeding reactants to a distillation column reactor into a feed zone, and concurrently contracting the reactants with a fixed-bed

Catalytic distillation. Patented in 1980 by Lawrence A. Smith of Chemical

FIG. 2. Map of global national gasoline sulfur limits, 2014. Source: International Council

on Clean Transportation.

TABLE 1. Timeline of global diesel sulfur limits, ppm, 2014 Country

2005

Brazil

3,500 2,000

2006

China

2,000

EU-27

50

India

500

Japan

50

Russia

500

Thailand

150

US

500

South Africa

3,000 500

2007

2008

2009

2010

2011

2012

2013

1,800–500 transition

2014

2015

2016

2017

2018

2019

2020

2019

2020

500

350

50

10

10 350

50

10

350

50

10

10 50 15 10

Source: International Council on Clean Transportation

TABLE. 2 Timeline of global gasoline sulfur limits, ppm, 2014 Country

2005

Brazil

1,000

China

500

EU-27

50

India

500

Japan

50

Russia

500

2006

2007

2008

2009

2010

2011

2012

2013

2014

2015

2016

2017

2018

50 150

50

10

10 150 10 150

50

10

Thailand

150

US

30/90/300 30 (avg)/80 (cap)

50 10

South Africa

1,000

10

500

Source: International Council on Clean Transportation HYDROCARBON PROCESSING | OCTOBER 2016 | ADVANCES IN SULFUR MANAGEMENT

S–57

ADVANCES IN SULFUR MANAGEMENT Liquid-phase processing. This is a ma-

ture set of sulfur removal technologies, including those by Merichem, which in 1974 pioneered fiber-film contactor6 for caustic, amine and acid treating. Featured here is Merichem’s iron-based reductionoxidation7 (redox), which scales well and is capable of full turndown. Sulfur recovery via iron-based, liquid redox processing is a commercially proven approach to refinery sulfur management. The inherent characteristics of liquid redox processing, such as 100% turndown in respect to hydrogen sulfide (H2S) concentration, flowrate and sulfur loading, as well as single-stage removal efficiencies in excess of 99.9%, make the process attractive as a standalone device for applications of less than 15 tpd of sulfur, or in conjunction with a Claus unit at higher capacities, resulting in a system with an overall removal efficiency of 99.9+% and 100% turndown capability. Another attractive feature of iron-based liquid redox processing is its ability to process any type of gas stream, such as fuel gas, amine acid gas and sour water stripper gas. Consequently, for refineries with Claus plants, the liquid redox process can be employed as a tail gas treatment unit, with or without a hydrogenation/hydrolysis unit, while also directly processing the sour water stripper gas, relieving the Claus unit of this burden. All of this can be accomplished without recycling gas back to the Claus unit, thereby increasing the capacity of the Claus unit. Options in Claus technology. Interestingly, a Wikipedia entry does not exist for “sulfur treating.” When searching for “sulfur recovery,” the site redirects the user to the Claus8 process, which recovers elemental sulfur. Increasingly stringent sulfur and air quality specifications prompted development of tail gas treating units (TGTUs), such as the Shell Claus Offgas Treating9 (SCOT) unit with amine scrubbing. Alternatives to SCOT include lower-capitalcost LT-SCOT,10 SUPERCLAUS11 with wet caustic scrubbing, EUROCLAUS12 by Jacobs, and CANSOLV technology by Shell, which published a study on a facility after 10 years of operation.13 BAYQIK technology enables more efficient conversion of process gas with high sulfur dioxide (SO2) concentrations. Bayer developed, piloted and commercialized this technology over the past 10 S–58

years, which was acquired by Chemetics14 of Jacobs group.15 A full-scale installation based on the technology has operated successfully for seven years. Reliance Jamnager refinery.This export

project 16 was built in 2009, and was chosen because it illustrates how a wellplanned major expansion project can be executed in record time. The owner worked with the same engineering, procurement and construction (EPC)/ technology integrator to assimilate to changing, stricter regulatory standards.17 When Reliance’s Jamnagar Domestic Tariff Area refinery began commercial operation in 1999, it was the largest refinery complex ever built from the ground up. Black & Veatch designed the original sulfur recovery facility with three sulfur processing units. In 2005, Reliance initiated plans to double the refinery’s capacity with the Jamnagar Export Refinery Project. Reliance retained Black & Veatch to build a second sulfur facility that was a near-exact clone of the original facility. However, changes to environmental regulations required refineries to capture more sulfur from acid gases than was mandated in 1999. This new regulation made the idea of an exact copy of the original plant impossible. The scope of the expansion work included a new 3 × 675-tpd cold bed absorption (CBA) sulfur recovery unit, and a new, common, 2.025-Mtpd TGTU. The expanded complex has a total sulfur processing capacity of 6 × 675 tpd. This type of project commonly takes 30 months to complete, but Black & Veatch completed it in 24 months. The company provided maximum replication of the existing facility design to provide operational consistency between facilities. It also reduced the operations learning curve for refinery personnel, with respect to commissioning, startup and operations of the new facilities. Environmental restrictions presented another key challenge. India had increased its recovery efficiency requirement for new units to 99.9%. Black & Veatch combined the benefits of the previously installed CBA technology with the latest tail gas treating technology to meet the requirement. The project was the first to use this approach in a complex of this size. 1

LITERATURE CITED Miller, J., “Global Comparison: Fuels,” Transport Policy, June 12, 2014, http://transportpolicy.net/

ADVANCES IN SULFUR MANAGEMENT | OCTOBER 2016 | HydrocarbonProcessing.com

index.php?title=Global_Comparison:_Fuels International Maritime Organisation, “Sulphur Oxides (SOx) Regulation 14,” IMO, 2016, http:// w w w.imo.org/en/OurWork/Env ironment/ PollutionPrevention/AirPollution/Pages/Sulphuroxides-(SOx)-–-Regulation-14.aspx 3 Tolson, A., “The transition to distillates is going to be messy,” Ship and Bunker, 2016, http://shipandbunker. com/news/features/2020-vision/365052-2020vision-the-transition-to-distillates-is-going-to-be-messy 4 Korpelshoek, M., K., Rock and R. Samarth, “Keeping it clean without affecting quality,” CDTech, July 2010, http://www.cbi.com/getattachment/66f7507b-268d47c4-9496-328ecb032640/Keeping-it-Clean-withoutAffecting-Quality.aspx 5 Podrebarac, G., R. Samarth and K. Rock, “Doing wonders for diesel production,” CDTech, November 2009, http://www.cbi.com/getattachment/bb32ade0c25e-4541-a1db-107dedf32d64/Doing-Wonders-ForDiesel-Production.aspx 6 FIBER-FILM technology, http://www.merichem. com/FIBER-FILM 7 Merichem, “Flexibility of Liquid Redox Processing in Refinery Sulfur Management,” 2016, http://www. merichem.com/company/overview/technical-lit/ tech-papers/liquid-redox-flexibility 8 Wikipedia, “Claus Process,” September 2016, https:// en.wikipedia.org/wiki/Claus_process 9 National Energy Technology Laboratory, “SCOT Tail Gas Treating,” NETL, 2016, https://www.netl. doe.gov/research/coal/energy-systems/gasification/ gasifipedia/scot-tgtu 10 Jacobs, “Low-Temperature SCOT Process,” 2016, http://www.jacobs.com/uploadedFiles/ www jacobscom/20_Learn_About_Us/25_ Products/253_Comprimo_Sulfur_Solutions/ Technologies/Handout%20Jacobs%20CSS%20-%20 LT-SCOT%20Process.pdf 11 Jacobs, “SUPERCLAUS Scrubber Process,” 2016, http://www.jacobs.com/uploadedFiles/ www jacobscom/20_Learn_About_Us/25_ Products/253_Comprimo_Sulfur_Solutions/ Technologies/Handout%20Jacobs%20CSS%20-%20 SUPERCLAUS%20Scrubber%20Process.pdf 12 Jacobs, “EUROCLAUS Scrubber Process,” 2016, http://www.jacobs.com/uploadedFiles/ www jacobscom/20_Learn_About_Us/25_ Products/253_Comprimo_Sulfur_Solutions/ Technologies/Handout%20Jacobs%20CSS%20-%20 EUROCLAUS%20Process.pdf 13 Edkins, N. and N. Moreton, “CANSOLV SO 2 Scrubbing: 10 years of reliable operation,” Shell Cansolv, October 2012, http://www.shell.com/ business-customers/global-solutions/shell-cansolv-gas-absorption-solutions/cansolv-news-andmedia-releases/shell-cansolv-papers/_jcr_content/par/textimage.stream/1446472635746/ c62a02d3534304587e817ef39a98b96023757a5f889ac315c1ea8fa64b9a8a08/paper4.pdf 14 Jacobs, “Chemetics process-technology based engineering design,” 2016, http://www.jacobs. com/workwithus/products/chemetics/index. aspx#Overview 15 Jacobs, “Jacobs acquires BAYQIK sulfuric acid converter technology from BAYER AG,” August 31, 2016, http://invest.jacobs.com/investors/PressRelease-Details/2016/Jacobs-Acquires-BAYQIKSulfuric-Acid-Converter-Technology-from-Bayer-AG/ default.aspx 16 Black & Veatch, “Reliance Petroleum Jamnagar Export Refinery featured project,” 2016, http://bv.com/ Projects/reliance-petroleum-jamnagar-export-refinery 17 Black & Veatch, “Innovative sulfur complex completed at world’s largest refinery,” September 22, 2009, http:// www.yourpetrochemicalnews.com/innovative+sulphu 2

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31M082016H

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CB&I PROVIDES ADVANCED SULFUR SOLUTIONS TO TODAY’S REFINERS CB&I has designed and built more than 300 sulfur complexes, and provides a comprehensive range of sulfur processing equipment, including amine units, sour water strippers, sulfur recovery units (SRUs) and tail gas treating units (TGTUs), along with oxygen enhancement and liquid sulfur degassing. The company’s scope of services for sulfur complexes includes basic engineering services, sulfur removal and recovery technology design packages, FEED package development, modular and turnkey construction, proprietary equipment and technical services, including process studies and startup assistance. CB&I’s proprietar y technologies can be applied to both new installations and to revamp SRUs and TGTUs: • The Claus Combustor™ burner is a ring style acid gas burner that provides thorough mixing of air and acid gas for high combustion efficiency and ammonia destruction. It is also suitable for an oxygen enhanced mode of operation. • The patented SRU Waste Heat Boiler Tubesheet protection system uses ceramic fiber insulation board, high alumina tube ferrules and insulating refractory design to protect the tubes and tubesheet against high temperature during both air and oxygen enhanced modes of operation. • SulfSep™ is an entrainment separator designed to minimize sulfur carryover from the sulfur condenser plenum chambers. It has a removal efficiency of 99.9% for sulfur droplets 10 microns and larger. • The company’s proprietary degassing system reduces the H2S in the liquid sulfur to less than 10 ppm.

significant capacity increases by introducing oxygen directly into the thermal reaction zone via the Oxygen Injector™. It uses a patented cooling technique to eliminate thermal degradation from radiant heat. • OxyPAC™ Advanced Modular Control System. By regulating the transition in and out of an oxygen enhanced mode of operation, this proprietary control system helps maximize the air mode while limiting oxygen consumption and costs.

TAIL GAS TREATING CB&I offers a family of Resulf™ tail gas treating technology. When this is used along with the SRU, overall sulfur recovery of up to 99.99% (H2S < 10 ppm) is achievable.

CONTACT INFORMATION 2103 Research Forest Drive The Woodlands, TX 77380 USA Tel: +1 832 513 1000 Fax: +1 832 513 1005 Email: [email protected] Website: www.CBI.com

OXYGEN ENHANCED OPERATIONS When regulations or refinery operations drive the need to increase capacity while minimizing capital investments, CB&I offer the following technologies for increasing the sulfur recovery capacity of the SRU: • Oxygen Injection™ System. This proprietary technology achieves SPONSORED CONTENT

HYDROCARBON PROCESSING | OCTOBER 2016 | ADVANCES IN SULFUR MANAGEMENT

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ENERSUL

ENERSUL—LEADING THROUGH INNOVATION Enersul, a Marmon Group/Berkshire Hathaway Company, has thrusted the sulphur industry forward for decades by sticking to its core guiding principal; to provide the sulphur industry with leading technical and operational solutions catered to the changing needs of their clients. Enersul’s consistent and proven track record of continuous innovation, is supported by often developing proprietary technologies and processes that outpace the demands of the market. This is apparent in almost every process currently used in forming, handling, and transporting elemental sulphur.

GX™ M SERIES OF PREMIUM SULPHUR GRANULATORS Enersul’s industry leadership is perhaps best exemplified by proprietary sulphur granulation forming technologies, offering a wide array of technical and operational solutions for any requirement and any climate. The Enersul GX™ M Series of premium sulphur granulators was first introduced with the GX™ in the late 1970s. The GX™ granulation process is a size enlargement process. Small sulphur seeds (undersized granules) are repeatedly coated with a liquid sulphur spray and cooled. With repeated application of liquid sulphur, the seed increases in volume and weight. As the granule is enlarged, each coating of liquid sulphur is fully and structurally bonded to the layer beneath. This creates a spherical granule which is entirely dry and completely free of voids. Enersul’s GXM1™ (1250 tpd) was followed with the GXM2™ and more recently the GXM3™, the world’s first patented Single-Pass™, modular sulphur forming unit. Delivered 90% pre-assembled, the GXM3™ represents a giant leap in reducing construction, commissioning time and costs. In late 2016, Enersul will announce a new “M” series modular version of the GX™ unit to provide producers the same GX™ benefits at higher capacities with a smaller footprint.

WETPRILL™—SULPHUR FORMING PROCESS In the WetPrill™ process, molten sulphur is pumped onto perforated trays that direct the sulphur in narrow streams into an agitated water bath. Pellets form as the sulphur comes in contact with the water. The low thermal conductivity, high specific heat, and long transformation time of sulphur make it necessary to maintain the pellets in suspension for as long as possible, allowing them to harden. This process produces spherical, uniformly sized, low moisture pellets. Enersul’s WetPrill™ product is known for its low friability, low moisture content, and high bulk density. This is due to the shape of the prill, which is small, round, and uniform, with few entrained fines. Enersul offers both a Mini WetPrill™ (100tpd unit) and standard WetPrill™ (up to 1000 tpd) of production capacity.

WETPRILL2000™

New in 2016, the WetPrill2000™ is the latest design from Enersul using the long established WetPrill™ process. This latest model achieves output of 2000 TPD, the largest standard output of any Enersul WetPrill™ technologies, with improvements to both layout and plot space requirements.

SPONSORED CONTENT

HYSPEC™ SULPHUR DEGASSER Enersul developed the HySpecTM H2S degassing process to quickly, effectively and economically reduce the H2S content of liquid sulphur to low levels, to 10 ppm or less. The HySpec™ is a modular system composed of a series of reactors through which the process degasses the molten sulphur. The number of reactors determines the degassing capacity of the system, resulting in a customizable and modular system that is adaptable to various capacity requirements.

SAFEFOAM TRANSFER SYSTEM (STS™) New in 2016, the Enersul SafeFoam Transfer System—STS™ improves the handling of solid bulk sulphur by reducing fugitive fines at critical transfer points throughout any sulphur handling system. The result is a significantly safer and environmentally friendly sulphur handling system. The STS™ arrives 90% preassembled resulting in significantly lower assembly rates and commissioning times. The majority of the required system checks are completed prior to shipment. Ongoing operating costs are also reduced as the STS™ eliminates empty belt spraying and uses direct drive motor coupling. The unit is automated through PLC for operating at optimum levels which are set according to the specific conveying requirements.

REMELTING Sulphur Blocks are an excellent long term storage option and eventually, these blocks will be remediated and the sulphur moved. In the 1990’s Enersul offered innovative solutions for remediating sulphur blocks quickly with the High Efficiency Melter (HEM™) series along with pit and tank type remelt operations. In addition Enersul successfully tackled highly contaminated lower sections of the sulphur blocks at various sites. Through this extensive operational experience the High Contamination Sulphur Remelter (HCSR™) was developed. Both technologies were born from onsite, real-world experience, developed specifically for the needs of Enersul clients. New in 2016, Enersul will be introducing a innovative modular clean/contaminated sulphur remelter that can be applied for both block and formed sulphur remelting.

OPERATIONAL SOLUTIONS Enersul’s Operational Solutions has a depth of experience unmatched in the field. Long-standing relationships with sulphur producers have established Enersul’s reputation as a leader in reliability, safety and environmental consideration. The experience of having taken complete operational control of the sulphur requirements of a variety of projects enables Enersul to innovate their offerings to meet the needs of the real world. Lessons learned by Enersul’s domestic and international project teams are applied across every aspect of Enersul’s products and services. This is why Enersul is the industry leader For Everything Sulphur for over half a century. Enersul is committed to continuously improve, innovate and develop technologies for operations in the real world of sulphur handling. That is what innovation means to us, applying our decades of realworld experience to making the forming, handling, storing and transportation of sulphur cost competitive, safer and more environmentally friendly.

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ADVANCES IN SULFUR MANAGEMENT

LOW-SULFUR PROJECTS DOMINATE THE DOWNSTREAM CONSTRUCTION LANDSCAPE LEE NICHOLS, Editor/Associate Publisher

As the world continues to see more vehicles on the road, and as emerging economies invest in civil, industrial and energy projects, global fuels demand is forecast to increase through the end of the decade. The increased number of vehicles on the road equates to higher emissions rates and, in turn, more airborne pollutants. To combat these effects, legislation mandating decreased emissions and lower levels of airborne pollutants is coming into effect. In response, refiners are implementing operational and processing changes to reduce sulfur levels in transportation fuels. Nations are investing heavily in the construction of new processing units to produce higher-value products. These investments will help produce high-quality fuels that meet Euro 4, Euro 5 and Euro 6 specifications. Global refiners have adopted European standards for fuel quality, as Europe has been the frontrunner on regulations for low-sulfur, “clean” transportation fuels. European passenger vehicle emissions standards for Euro 4, Euro 5 and Euro 6 are detailed in TABLES 1 and 2. These standards promote the reduction of carbon monoxide (CO), nitrogen oxide (NOx ), hydrocarbons (HCs) and particulate matter (PM) in both diesel- and gasoline-fueled passenger vehicles. The refining industry has already made incredible strides in reducing sulfur

in transportation fuels. As shown in FIG. 1, sulfur levels in diesel fuel have decreased dramatically around the globe within the past decade. Refiners have invested, and continue to invest, billions of dollars in new units, upgrades/retrofits and expansions to meet new sulfur and emissions regulations. These investments will help produce higher-quality transportation fuels and continue to move the industry toward a low-sulfur world. US and Canada. Every country is doing its part to add the necessary processing units to meet stricter sulfur regulations. In the US and Canada, refiners are investing to meet new Tier 3 fuel regulations. Scheduled to begin in January 2017, sulfur content in gasoline will be limited to 10 parts per million (ppm). This limit is a reduction from Tier 2 standards, which reduced the sulfur content in gasoline to 30 ppm. The US Environmental Protection Agency (EPA) forecasts that the new rule will significantly reduce vehicle pollutants to the atmosphere. For example, the EPA forecasts that NOx emissions will be lowered by approximately 260 Mt (thousand tons) in 2018 alone. According to EPA documents, a number of flexibilities are being offered that will provide nearly six years of lead time for refineries that may need extra time to

TABLE 1. EU emissions standards for passenger vehicles (gasoline) CO, g/km

HC, g/km

NOx, g/km

PM, g/km

Euro 4

1.0

0.10

0.08

–

Euro 5

1.0

0.10

0.06

–

Euro 6

1.0

0.10

0.06

0.005

TABLE 2. EU emissions standards for passenger vehicles (diesel) CO, g/km

HC + NOx, g/km

NOx, g/km

PM, g/km

Euro 4

0.50

0.30

0.25

0.025

Euro 5

0.50

0.23

0.18

0.005

Euro 6

0.50

0.17

0.08

0.005

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ADVANCES IN SULFUR MANAGEMENT | OCTOBER 2016 | HydrocarbonProcessing.com

comply with the new regulations. These flexibilities are listed on the EPA website and include: • A credit averaging, banking and trading (ABT) program that will allow refiners to spread out their investments from 2014 through 2019, and provide a seamless transition from the Tier 2 ABT program • A delay in the start date to 2020 for approximately 30 small refiners and small-volume refineries • A 1-yr-deficit carry-forward provision that allows an individual refinery that does not meet the 10-ppm standard in a given year to carry a deficit forward for 1 yr, as long as the refinery makes up the deficit in the following year • Hardship provisions that allow refiners to petition for compliance assistance on the basis of extreme hardship or extreme unforeseen circumstances. Large US refineries (those producing more than 75 Mbpd) must comply with Tier 3 standards by 2017. Refiners producing less than 75 Mbpd must meet Tier 3 regulations by 2020. To comply with new regulations, US refiners must invest in additional units, such as hydrotreaters, to reduce the sulfur content of transportation fuels. In Canada, petroleum fuels constitute 95% of Canada’s transportation energy needs. The country has aligned itself closely with US fuel standards and is making strides to continually reduce sulfur levels in transportation fuels. This includes the introduction of Tier 3 fuel regulations for passenger vehicles and light-duty trucks. These fuel standards will begin in 2017, which coincides with the startup of US Tier 3 regulations. Canadian refiners have already invested more than $8 B over the past decade to reduce sulfur levels in gasoline and diesel fuels.

ADVANCES IN SULFUR MANAGEMENT Asia-Pacific. In the Asia-Pacific region, China and India are investing heavily in capital-intensive projects to decrease sulfur levels in transportation fuels. Both countries have announced stringent fuel regulations to curb air pollution, particularly in major cities. China has set aggressive fuel economy standards through 2020, and it is implementing its National V fuel quality standard, which equates to Euro 5-standard transportation fuels. Recent regulations required refiners to produce Euro 4-standard transportation fuels nationwide by the end of 2015.

Euro 5-standard transportation fuels will be required for the automotive industry by 2017. These new regulations are being implemented a year ahead of schedule. The implementation of National V fuel quality standards for non-automotive diesel has been pushed back a year to January 2018. This includes “general” diesel used in agriculture and industry. General diesel will need to meet Euro 5 standard requirements within this time frame. Upgrading the nation’s fuel quality could cost Chinese refiners over $7 B. However, the country is writing new reg-

ulations for the implementation of National VI. According to China’s Ministry of Environmental Protection, the plan for National VI will be finalized by the end of 2016. At the time of publication, no timetable has been announced for the adoption of National VI, but the standard will likely go into effect no later than 2020. In total, India’s refiners are investing more than $30 B in fuel quality projects through the early 2020s. Capital expenditures are expected to be even higher due to new regulations to curb air pollution and produce Euro 6-standard fuels by 1Q

15 and below* > 15–50 > 50–500 > 500–2,000 > 2,000–5,000 > 5,000 and above Conflicting/missing data

15 and below* > 15–50 > 50–500 > 500–2,000 > 2,000–5,000 > 5,000 and above Conflicting/missing data * Information in parts per million (ppm) FIG. 1. Sulfur levels in diesel fuel: Global status 2005 (top) vs. 2015 (bottom). Source: United Nations Environment Program, PCFV Secretariat. HYDROCARBON PROCESSING | OCTOBER 2016 | ADVANCES IN SULFUR MANAGEMENT

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ADVANCES IN SULFUR MANAGEMENT 2020. In January 2015, the Indian government announced that India will skip the implementation of Bharat Stage 5 (BS-5) standards and move directly to BS-6 standards. BS-6-standard fuels are equivalent to Euro 6 specifications. At the time of publication, BS-3 standards are mandatory across the country, with BS-4 standards mandatory in major cities such as New Delhi and Mumbai. BS-4-standard fuels will be mandatory nationwide by 2Q 2017, and BS-6-standard fuels will be required by 2Q 2020. The BS-6 regulation is being imposed four years ahead of schedule and calls for a 68% reduction in NOx emissions. These new standards were proposed in response to a World Health Organization (WHO) study that found that 13 of the world’s most polluted cities were in India. Air pollution has become such an issue that some cities, including New Delhi, are restricting drivers to using their vehicles every other day. The government is also investing in the construction of compressed natural gas (CNG) fueling stations in the hope that its citizens will switch to the cheaper, more fuel-efficient transportation option. India’s Road Transportation Minister, Nitin Gadkari, said that the new regulations will cost Indian refiners nearly $5 B. This spending includes the installation of secondary units to comply with the BS-6 fuel standard. Indonesia is the world leader in the production of palm oil, and is promoting its use as a biofuel. The country boosted the mandated amount of blending in diesel in 2014 from 7.5% to 10%, and subsequently to 15% in 2015. Indonesia raised the blending requirements to 20% (referred to as B20) in 2016, and plans to increase the requirement to 30% in 2020. According to the Indonesian Biofuel Producers Association, Indonesia’s biodiesel consumption will climb from 1.1 kiloliters in 2015 to 7.9 kiloliters by the end of 2016. The additional biofuels usage is expected to lower vehicle emissions substantially. Throughout 2016, however, the Indonesian government has been able to meet only a fraction of its intended blending targets for B20. According to a Biofuels Digest report, just 50,000 kiloliters of biodiesel were blended with non-subsidized fossil diesel during the first half of 2016, even though required mandates called for more than 650,000 kiloliters to be blended. It S–66

remains to be seen if the B20 fuel implementation will be successful. South Korea is also mandating initiatives to improve air quality. The country aims to match European air quality by 2025. To accomplish this goal, South Korea will take the following measures: • Institute real-world emissions tests of diesel vehicles, beginning in 2017 • Raise the number of eco-friendly vehicles to 30% of new car sales (from 2.6% at present) by 2020 • Increase the amount of vehicle charging stations tenfold • By 2019, dispose of diesel vehicles built prior to 2005 • Phase out diesel-powered buses and replace them with CNG-fueled models • Shut down old thermal plants. The country is Asia’s second-largest diesel car market, and South Korea hopes these steps will help improve its air quality. The Middle East. Some of the most capital-intensive projects to lower sulfur in transportation fuels will take place in the Middle East. Through grassroots facilities, expansions and debottlenecking projects, the Middle East plans to add nearly 1.5 MMbpd of refining capacity by 2019. In total, the region is expected to add 1.9 MMbpd between 2015 and 2020. This forecast represents an investment of more than $50 B. Refining capacity will center on domestic demand and export opportunities to markets in Asia-Pacific and Europe, as well as the production of lowsulfur fuels that meet Euro 5 specifications. Traditionally, Middle East refineries have had simple configurations and high fuel oil yields, partly due to strong power generation requirements. This condition is changing. A new generation of highly complex plants, combined with upgrades and expansions at existing plants, is radically altering the product mix. New unit configurations include hydrocracking, catalytic cracking and hydrotreating capacities designed to minimize fuel oil output and maximize low-sulfur middle distillate, diesel and gasoline production. Saudi Arabia and Kuwait are leading the charge in new clean fuels projects in the region. To comply with mandatory sulfur specifications for gasoline and diesel, Saudi Arabia is spending billions of

ADVANCES IN SULFUR MANAGEMENT | OCTOBER 2016 | HydrocarbonProcessing.com

dollars to construct multiple clean fuels projects. The country is seeking to reduce sulfur content in diesel and gasoline to 10 ppm and to lower benzene content in gasoline to 1%. This represents a dramatic shift in sulfur levels from 2012, when Saudi Arabia’s maximum sulfur level for diesel was greater than 500 ppm. The country plans to commission its 400-Mbpd Jazan refinery by 2018. The refinery will produce higher-grade transportation fuels, including ultra-low-sulfur diesel (ULSD). Along with its JV partners, Saudi Aramco will upgrade all of its domestic refineries to produce lower-sulfur transportation fuels. Several projects—the Ras Tanura clean fuels and aromatics project (which had been placed on hold, but was reinstated in mid-2015), the Riyadh clean transportation fuel project, the Saudi Aramco Mobil Refinery Co. clean fuels project (completed in 2014) and the PetroRabigh clean fuels project—are designed to accomplish Saudi Arabia’s goal of producing near-zero-sulfur fuels. Kuwait is investing more than $30 B on ambitious plans to overhaul its refining sector and become the region’s clean fuels leader. The plan focuses on modernizing and integrating the country’s Mina Abdullah and Mina Al-Ahmadi refineries, as well as building the region’s largest refinery, the Al-Zour plant. Once completed, the reconfigured and integrated Mina Abdullah and Mina Al-Ahmadi refineries will decrease the sulfur in gasoline production from 500 ppm to less than 10 ppm. Benzene and aromatics concentrations will also decrease. Bunker fuel oil sulfur content will be reduced from 4.5 ppm to 1 ppm, and maximum sulfur content of full-range naphtha will drop from 700 ppm to 500 ppm. With the construction of the Al-Zour plant and the upgrading and integration of its domestic refineries, Kuwait is set to become the largest producer of clean fuels in the Middle East by 2019. Other countries in the region are making sizable investments to produce higher-quality transportation fuels. Efforts include the Ruwais refinery expansion (completed in 2015) in the UAE; the Jebel Ali and Fujairah projects in the UAE; the Sohar refinery upgrade and Duqm refinery projects in Oman; the Sitra refinery modernization project in Bahrain; and the SOCAR Turkey Aegean refinery project in Turkey.

OHL GUTERMUTH

TAILOR-MADE VALVES “MADE IN GERMANY” OHL GUTERMUTH MAKES SPECIAL VALVES FOR THE WORLD’S LARGEST PROJECTS Valve manufacturer OHL Gutermuth offers a broad range of butterfly and linear valves for both shut-off and control duties, plus special types, custom designs and accessories. Nominal diameters are up to DN 4000, with pressure ratings of 200 bar and more, for temperatures from –196°C to 1,450°C. The company is producing triple offset butterfly valves for 25 years to be used in LNG, chemical and petrochemical applications as well as solar power and shipbuilding industries. The enterprise traces its origins back to 1867. “Providing individual advice and consulting to our customers right from the start is at the centre of our efforts,“ explains Managing Director Wolfgang Röhrig. The Altenstadt-based company invests heavily in quality assurance, with certification to ISO 9001:2008, ISO 14001:2009 and Module H of the EU’s Pressure Equipment Directive. All products are also certified under the Russian GOST and RTN standards and licensed for use by Gazprom and most of the other players in the Oil & Gas industry. Since 2007, the company has operated a sales office in Beijing, and opened a company in Moscow in 2013. “In the past 40 years we have supplied customized valves for more than 150 gas purification plants, among them the largest in the world, in Europe, Russia, Kazakhstan, Turkmenistan, India, China, the Middle East and America“, says Röhrig.

Recent projects include a list of LNG terminals also valves for FLNG plants have been supplied. OHL Gutermuth supplies valves for the world’s largest thermosolar power plants, Methanol and coal gasification plants. We also supplied valves for the German and French navies, and for a 170 m-long mega-yacht belonging to a Russian oil billionaire, this shows our broad range of products and customers.

A control and shut off technique you can rely on.

OHL Gutermuth

Others simply sell you a product – we offer a solution.

OHL Gutermuth Industrial Valves GmbH

BEST VALVES

MADE IN

Y

GERMAN

SINCE 1867

Customized Valve Design

Helmershäuser Str. 9+12 63674 Altenstadt/Germany Phone +49 6047.8006-0 Fax +49 6047.8006-29 www.ohl-gutermuth.de [email protected]

WE PROVIDE CUSTOMIZED SYSTEM SOLUTIONS FOR GAS DESULPHURIZATION, WITH THE EXPERTISE OF AN INTERNATIONAL LEADER IN INNOVATION. RESULT: BETTER RELIABILITY, BETTER ECONOMY.

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THIOPAQ O&G stable by nature

The proven gas desulphurisation technology.

How to reach highest value when treating natural gas streams for sulphur? THIOPAQ O&G puts you in control of sulphur removal and sulphur recovery. Perform well on safety, sustainability, reliability, cost and operability. Oil & Gas companies worldwide rely on THIOPAQ O&G. See why on paqell.com/thiopaq. Paqell’s THIOPAQ O&G - exceptional achievements in H2S removal.

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PAQELL

PAQELL IS A JOINT VENTURE OF SHELL AND PAQUES AT PAQELL WE BELIEVE IN THE POWER OF NATURE.

Since 2011 the joint venture of Paques and Shell Global Solutions sells a safe, stable and environmentally friendly desulphurisation process for removing H2S from sour gas streams. The unique aspect of the main product, THIOPAQ O&G, is that it utilises naturally occurring bacteria to oxidise H2S to recover elemental sulphur. Our solution offers the Oil & Gas Industry a new, sustainable solution to choose from. In an environment where health and safety are of the utmost importance, Paqell can provide you with the competitive advantage your company is looking for. Especially since the THIOPAQ O&G unit is a relative low investment and can be operable within weeks. The process was originally marketed by Paques BV for the treatment of biogas, which is produced by the anaerobic digestion of waste water. Co-operation with Shell Global Solutions led to further development of the process for application at high pressure in oil and gas environments. It can be economically applied to projects recovering up to 100 t/d of sulphur.

PROCESS DESCRIPTION The process integrates gas purification with sulphur recovery in one unit. The feed gas is first scrubbed with a mildly alkaline sodium hydroxide solution. This solution absorbs the H2S to form sodium bi-sulphide, and sweet gas exits the contactor. Depending on the sour gas pressure, the bi-sulphide-rich solution is routed to a flash vessel or directly to the bioreactor, which operates at atmospheric pressure and ambient temperature, where a controlled amount of air is introduced. Naturally occurring bacteria consume the bi-sulphide ions and excrete elemental sulphur, which is separated from the circulating solution. The process produces hydroxide ions that effectively regenerate the caustic solution used in the absorption step, which reduces the consumption of chemicals. The process can replace a complete train of H2S removal and sulphur recovery installations, see FIG. 1. Another option is to retain the amine unit (for example, when carbon dioxide removal is also required) and replace only the sulphur recovery unit, the tail-gas treating unit, the degasser and, possibly, the incinerator by a single THIOPAQ O&G unit. Simplicity translates in lower cost. In excess of 99.9% of the H2S can be removed, treating gas to meet pipeline specification. Very low maintenance requirements results in >99% availability.

ADDED VALUE FOR YOUR BUSINESS THIOPAQ O&G technology offers a series of benefits: • reduced operating costs. The expensive chemicals required for liquid redox processes are not necessary; only sodium hydroxide and nutrients are required. • reduced capital expenditure. The process operates at ambient temperature and does not require fired equipment such as burners and reboilers. The regeneration and sulphur recovery sections operate at ambient pressure and -temperature. • ease of operation. The biologically produced sulphur is hydrophilic. This feature eliminates plugging problems. THIOPAQ O&G requires minimal operator attendance. SPONSORED CONTENT

FIG. 1. THIOPAQ O&G process

• safety. An additional feature of the process is that there is no free H2S (no acid gas) after the bio-reactor. No fired equipment or high pressure is required.

REFERENCES Over 200 THIOPAQ plants haven been installed since 1993 in a variety of industries and clients all around the world have chosen THIOPAQ O&G as their preferred solution for sulphur removal and recovery. In the Oil and Gas industry units have been installed in e.g. USA, Canada, Mexico, China, The Netherlands, and Indonesia. Currently plants are under design and construction in Indonesia, Germany, Australia and Belgium. It is clear that the technology has been accepted by the international oil & gas community as a welcome alternative to Claus/SCOT and redox processes.

WEBSITE Our website Paqell.com offers more detailed process information as well as a handy online ‘quick scan’ to assess suitability to your project. An animated movie of the THIOPAQ O&G process helps to get a quick understanding of the technology. Please contact us via our website if you are interested in the THIOPAQ O&G technology. We are happy to discuss with you the possibilities for your application. More information can also be obtained through our authorised licensors Cameron (based in USA), Frames (based in The Netherlands) and Paques Environmental Technology Shanghai (based in China). At Paqell we are convinced that the power of nature can provide a simple, safe & cost effective solution to traditional (chemical) sulphur removal at lower total cost of ownership.

HYDROCARBON PROCESSING | OCTOBER 2016 | ADVANCES IN SULFUR MANAGEMENT

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Available October 14th

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HPI Market Data 2017 features: • Global spending in the refining, petrochemical and gas processing sectors • Capital, maintenance and operating spending broken out by region • Short-term and long-term implications of today’s low crude oil prices • An exploration of changing markets and demand within the global HPI, with discussion of emerging markets • More than 40 tables and 80 figures, including information and data collected from governments and private organizations • Editorial analysis of worldwide economic, social and political trends driving HPI activity across all sectors

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Process Engineering and Optimization G. G. PIPENGER, Amalgamated Inc., Fort Wayne, Indiana

Build a diesel fuel performance additive, the right way—Part 1 Many diesel fuel additive companies in North America supply different additive products formulated for diesel-powered equipment. Each of these vendors claim that their specific products are the best available to upgrade diesel fuel for the fuel producer, fuel supplier and final fuel consumer. However, actual testing has often proved to the contrary. Nearly all diesel fuel additive suppliers make generic product claims that their additives will “Yield better fuel economy,” “Increase engine power,” “Reduce smoke,” “Increase fuel lubricity,” “Reduce system deposits,” “Lower maintenance costs,” “Save money,” etc. Unfortunately for the additive purchaser, there is no global “watchdog group” or oversight mechanism to monitor and verify diesel fuel additive product claims. The burden of proof for any diesel fuel additive’s performance has become the sole responsibility of the additive product purchaser. While some fuel additive claims can be verified with laboratory testing, proving other claims can only be conducted in the field by treating the diesel fuel and then actually running the “additized” fuel in the buyer’s daily fleet operations. Unless diligent steps are taken to laboratory test the prospective product before the purchase is made, there is no guarantee that the chosen product will achieve the benefits claimed. Diesel fuel additive testing begins with choosing a reputable American Society for Testing and Materials (ASTM) qualified laboratory with established experience in testing diesel fuel additives. This can be challenging for the diesel fuel ad-

ditive buyer, but a good laboratory can be found if the right questions are asked. The laboratory search should begin with a review of ASTM D-975, “Standard specification for diesel fuel oils.”1 This document lists the various physical parameters that must be met for all diesel fuels, whether in an on-highway or offroad application. Unfortunately, ASTM D-975 provides only the basic recommended physical fuel property boundaries, and does not specifically address performance upgrades achievable with diesel fuel additives. However, many of the same laboratory tests described in ASTM D-975 are appropriate to precisely determine the enhancements achievable with a prospective diesel fuel additive. This work is Part 1 of a series, and it will review the details of the diesel fuel physical properties (and additive enhancements) required to optimize summer diesel fuel performance in today’s common-rail, fuel-injection diesel engines. Part 2, to appear in the November issue, will address relevant information regarding diesel fuel additives for modern common-rail, fuel-injection engines during cold-weather operations. Cetane improver additive. This additive, which is predominantly 2-ethylhexyl nitrate, is the main chemical additive component used to decrease the ignition delay time (raising the fuel’s engine cetane number [ECN]) and improve the combustibility of a diesel fuel. Improving combustibility by raising the ECN is important, as there is a finite amount of time (microseconds) between fuel injec-

tion and exhaust during the combustion cycle in each cylinder. A diesel fuel ignited at the optimum time, before piston top-dead-center (TDC) of the compression stroke, will burn correctly and release the maximum Btu from each fuel droplet during the power stroke. The fuel’s heat energy is converted into usable power that pushes the piston down, resulting in the production of maximum horsepower (hp). This directly correlates into lower unburned hydrocarbon emissions and improved fuel economy. Enhanced fuel combustion manifests as less-visible smoke emissions, decreased combustion chamber fuel-related deposits, fewer post-combustion exhaust gas recirculation (EGR) valve and diesel particulate filter (DPF) deposit cleanings, reduced driver low-power complaints and less vehicle downtime. In North America, ASTM D-975 requires a minimum ECN of only 40. While some diesel fuel refiners produce diesel fuels well above this minimum, others supply diesel fuels only two or three numbers above 40. The average ECN ranges from 44 to 46, despite the fact that diesel equipment operators know that engines run more efficiently on fuels with a much higher ECN. Enhancing diesel fuels with a cetane improver additive treatment is often left to the purchaser. In theory, any increase in a diesel fuel’s ECN should provide some improvement in engine operation. However, only an increase of 4 to 5 ECNs will actually be noticed by the driver. This is especially relevant with the increase of European diesel vehicles entering North American markets. These diesel engines are designed to Hydrocarbon Processing | OCTOBER 2016 71

Process Engineering and Optimization operate best on the European-mandated 51-plus ECN diesel fuels, and indications are that the European mandate will increase to 55-plus in 2019 or 2020. It is important to note that an ECN increase of 4 to 5 numbers is needed to effectively measure and document a significant fuel economy improvement. It is vital to test a prospective diesel fuel additive with the buyer’s diesel fuel in an engine test cell under ASTM D-613 test2 procedures, “Standard test method for cetane number of diesel fuel oil,” to verify the additive’s potential ECN increase (FIG. 1). Detergent additive. This predominantly amine-based chemistry is the major additive component used to clean existing fuel-related deposits and prevent them from reoccurring in the fuel delivery system. Detergent additives also play an important role in diesel fuel combustion and power production, as any deposits in fuel pumps or fuel injectors will negatively affect the spray pattern produced in the engine cylinders during each injection cycle. If the fuel spray pattern droplets are not uniform, or the amount of injected fuel is impeded or limited due to internal or external deposits on even one fuel injector, then optimum combustion and maximum power production are impossible (FIG. 2). This condition will produce increased smoke, more unburned hydrocarbon emis-

FIG. 1. ASTM D-613 engine cetane test cell.

FIG. 2. Peugeot XUD9 fuel injector test apparatus and test injector tip photos.

72 OCTOBER 2016 | HydrocarbonProcessing.com

sions and reduced engine hp for all throttle settings, leading to increased downtime. Unburned diesel fuel related to poor fuel detergency will increase deposits in the engine combustion chamber, escalate deposits in post-combustion areas (exhaust valves, EGR valves, DPF, etc.), increase fuel dilution in the crankcase oil and shorten the normal engine maintenance overhaul period. These situations increase costs and decrease the useful life of diesel-powered equipment. An ASTM test procedure or easy rating method to determine the detergent content in diesel fuel or diesel fuel additive does not exist, nor does an easy methodology to determine a detergent’s effectiveness in keeping fuel injectors free of deposits and operating properly. The only recognized means of testing diesel fuel detergency is the costly and time-consuming Coordinating European Council (CEC) F-98-08 (S), “Direct injection, common-rail diesel engine nozzle coking test.”3 Diesel fuel additive suppliers should be required to “certify” that their particular additive product treated at the recommended treat rate will achieve a DW-10 pass rating. This diesel engine injector test is conducted over a 72-hour cycle period, alternating high-speed/load and low-speed/ load to determine the power loss resulting from fuel injector deposits. The test uses a European Peugeot 4-cylinder, 2.0 liter, direct-injection turbocharged light-duty, common-rail engine with a maximum injector pressure of 1,600 bars. The CEC F-98-08 test runs for a specified time using a base diesel fuel treated with 1 ppm of zinc (to increase injector deposits). As the fuel injector deposits increase, fuel flow through the injector and power production decrease. The same diesel fuel treated with detergent additive is tested under the same conditions and compared with the non-additized fuel. A power loss of less than 2% during the test with detergent-treated fuel is considered a DW-10 pass. Lubricity additive. This diesel fuel additive component (non-acid synthetic type) is important, as the diesel engine fuel delivery system is lubricated by the diesel fuel itself. If the diesel fuel’s lubrication value is inadequate, then the fuel pumps and injectors will not operate properly, leading to increased wear.

Diesel engine fuel injectors are designed to operate with extremely high injection pressures (35,000 psi and more) to better atomize each fuel droplet. Fuel is injected multiple times through extremely fine injector tip holes during each injection cycle in the excessively high temperature environment of each engine cylinder. If the fuel does not provide proper lubrication to the system, then the fuel injectors will “stick,” causing a chatter-like noise, and the required fuel will not be injected. The net effect will be incomplete combustion, poor power production at all engine power levels and reduced fuel efficiency. Ultimately, without adequate fuel lubrication, the engine will seize. Diesel fuel additive buyers should undertake their own laboratory testing for diesel fuel lubricity value enhancement using ASTM D-6079-11, “Standard test method for evaluating lubricity of diesel fuels by the high-frequency reciprocating rig (HFRR).”4 This procedure requires only 90 minutes of laboratory time, and can be accomplished with a small amount of diesel fuel and additive to verify the lubricity enhancement claim. The HFRR test method (FIG. 3) measures the wear scar produced on a small metal ball reciprocated against a polished disc, which is immersed in the subject diesel fuel. The simulation provides the expected internal wear in fuel injectors and pumps in the diesel fuel delivery system related to the fuel lubrication value. Improved fuel lubricity results in a decreased HFRR wear-scar measurement, and a correlation to the improved lubrication of the fuel delivery system. While ASTM D-975 shows a maximum HFRR wear scar of 520 µm, the EU Engine Manufacturers Association recommends a 460 µm maximum rating for fuel lubricity. European engine manufacturers have lowered their recommended HFRR fuel lubricity for initial engine break-in to less than 400 µm of wear scar. The US is experiencing a significant influx of European dieselpowered equipment, all manufactured with common-rail fuel injection systems, and new diesel engines being built in the US are incorporating common-rail fuel injection. Stability additive. This chemical com-

ponent protects diesel fuel and ensures optimum engine performance. As an organic product, diesel fuel degrades, oxidizes and breaks down from the time it is

Process Engineering and Optimization refined until it is consumed. Determining diesel fuel’s rate of oxidation (instability) through laboratory testing defines the extent of degradation that will occur between its manufacture and its use. This rating is important because, as a diesel fuel oxidizes (degrades), it generates fine, free carbon particulates that are abrasive and often collect internally in fuel injectors. Unstable fuels also manifest in the fuel system as varnishes that coat and cause scoring of the moving parts. Varnishes cause sticking of the injectors, preventing the delivery of the proper amount of fuel to the engine and potentially stopping it entirely. Unstable diesel fuels will also create sludge materials that collect and build up in low areas of the fuel storage and delivery system (i.e., tanks and lines). The longer an unstable fuel is used, the more degradation byproducts will be formed. A measurement of diesel fuel thermal stability was determined in a petroleum laboratory using the ASTM D-6468-08 test, “Standard test method for high-temperature stability of middle distillate fuels.”5 FIG. 4 shows that test pads can be visually rated to define the free carbon created during the test, or analyzed using a laboratory light reflectometer. The more carbon that is created, the less light will be reflected (lower reflectometer result). Diesel fuels treated with a “good stability additive” will have test pad ratings of 2 or less, and a light reflectometer result greater than 90%. A quality stabilizing additive can be used to dramatically slow diesel fuel degradation, but it cannot completely stop the natural degradation process. Diesel fuel users must have the base (incoming) diesel fuel tested for stability to determine

FIG. 4. Dupont F-21 stability chart, pad readings.

the state of degradation before treating with a stabilizing additive product. The prospective stability additive should be tested at the same time and in the same base (incoming) diesel fuel to determine whether the stability additive will, in fact, stabilize the diesel fuel. It should be noted that other performance additive products can negatively affect the fuel’s oxidation rate (stability). Therefore, the purchaser should have fuel treated with any other additives and then stability tested. If necessary, extra stability additive should be added to counteract the negative effects of the other additives. Corrosion inhibitor additive. This compound should be added to diesel fuels because: • All diesel fuels naturally contain water, which is corrosive

• All water contains dissolved salt, which will create deposit buildup in fuel injectors. While the diesel fuel water content may be small (typically 40 ppm–100 ppm), that amount of moisture is more than enough to cause rust and corrosion in the fuel delivery system. Although the salt content in the water may seem extremely small (typically a few ppb), the dissolved salt particles continually circulate throughout the fuel delivery system during operation. Since most diesel engines return 75%– 80% of the fuel to the vehicle fuel tank as “return fuel,” the same salt particles can flow through the injector more than 500 times during the consumption of one tank of diesel fuel. This provides ample opportunity for the salt particles to form a deposit buildup inside the fuel injector.

FIG. 3. Two-place HFRR lubricity test apparatus.

FIG. 5. Diesel fuel storage tank pump/meter assembly and island filter with fuel corrosion. Hydrocarbon Processing | OCTOBER 2016 73

Process Engineering and Optimization TABLE 1. Component list for performance diesel fuel additives Additive component

"Incorrect" additive formulation

Beneficial effect in performance

"Correct" additive formulation

Beneficial effect in performance

Cetane improver

Up to 1 ECN increase

None

4 to 5 ECN increase

Much improved

Up to 2 ECN increase

Little to none

> 6 ECN increase

Optimum

2 to 3 ECN increase

Slight to some

Detergent

Maintains clean injectors

Inconclusive

DW-10 certified

Optimum

Lubricity

Lubricates injectors

Inconclusive

≥ 400µm HFRR

Optimum

Stabilizer

Stabilizes fuel

Inconclusive

Pad rating of ≤ 6 and > 90 reflectance

Optimum

Corrosion inhibitor

Prevents corrosion

Inconclusive

Provides ≥ NACE 1-A

Optimum

Deposit modifier

Prevents deposits

Inconclusive

Certified

Optimum

age rate. FIG. 6A shows a diesel engine fuel injector pintle after operation without a corrosion inhibitor. A similar pintle without corrosion resulting from fuel treated with an effective corrosion inhibitor additive is shown in FIG. 6B. There can be no doubt that the first pintle will not operate properly in the injector, nor supply the correct amount of diesel fuel into the cylinder. Deposit modifier additive. The final

FIG. 6. A diesel engine fuel injector pintle after operation without a corrosion inhibitor (A, left); and a similar pintle (B) without corrosion resulting from fuel treated with an effective corrosion inhibitor additive.

illustrates the corrosive effects on the fuel delivery system in a diesel fueling facility (after six months of continual use). Similar corrosive effects can occur in fuel injectors and pumps. A NACE spindle corrosion test (ASTM D-655)6 should be conducted to determine the effectiveness of the corrosion inhibitor additive. A NACE 1-A or better result will indicate the prevention of fuel-related corrosion and rust when the additive is treated at the correct dosFIG. 5

74 OCTOBER 2016 | HydrocarbonProcessing.com

chemical component is particularly appropriate for diesel fuels in engines with high-performance, common-rail fuel injection systems. The proper amount of deposit modifier additive treatment will: • Maintain cleaner intake valves, exhaust valves and piston tops to reduce hot spots • Keep the EGR valves free of excessive carbonaceous buildup and reduce the need for replacements • Reduce regeneration frequency through a better-maintained diesel particulate filter unit • Extend the useful life of the entire exhaust system before replacement. Although nearly all diesel fuel additive suppliers claim to “reduce fuel-related deposits,” no laboratory test method exists to determine the amount of deposit modifier additive included in a diesel fuel additive product. The only means of verifying the benefits claimed is to actually run the additive-treated diesel fuel in an engine for an extended period of time (TABLE 1). Prospective fuel additive buyers should request a written certified statement from the supplier that the product includes adequate deposit modifier additive to reduce or minimize fuel-related deposits. This certification may eliminate exaggerations from unscrupulous sup-

pliers and help ensure that the additive buyer receives a valuable product. Takeaways. No organization oversees

the performance claims and marketing of diesel fuel additives. This article (and the subsequent Part 2) will assist the additive buyer in sourcing the best performance additives for diesel-powered equipment.

Next month. Part 2 of this article will appear in November. LITERATURE CITED ASTM International, ASTM D-975-15c, “Standard specification for diesel fuel oils,” Vol. 05.01. 2 ASTM International, ASTM D-613, “Standard test method for cetane number of diesel fuel oil,” Vol. 05.05. 3 Coordinating European Council, CEC F-98-08 (S), “Direct injection, common rail diesel engine nozzle coking test,” Iss. 8, August 2015. 4 ASTM International, ASTM D6079-11, “Standard test method for evaluating lubricity of diesel fuels by the high-frequency reciprocating rig (HFRR),” Vol. 05.02, 2016. 5 ASTM International, ASTM D-6468-08, “Standard test method for high temperature stability of middle distillate fuels,” Vol. 05.03, 2013. 6 National Association of Corrosion Engineers International, NACE TM0172 (ASTM D665/ D7548), “Diesel fuel spindle corrosion test method,” 2013. 1

GARY G. PIPENGER has worked in the additive manufacturing, chemical formulating, and marketing of fuel additives industries since 1973. He also has expertise in international sales and products marketing in the petroleum industry. Mr. Pipenger engineered and owns a chemical blending plant, and previously owned a petroleum testing laboratory. He developed a proprietary testing method to evaluate and quantify microbiological growth in distillate fuels, and assisted in the US military’s development of MIL-SPEC-S53021 for microbiological growth and storage stability prevention in distillate fuels. He is also a certified aircraft and engine mechanic. Mr. Pipenger earned a BS degree in industrial engineering and technology from Purdue University.

Process Engineering and Optimization K. E. LITZ, K. EDISON and J. RANKIN, Auterra Inc., Schenectady, New York

Remove sulfur and nitrogen from liquid hydrocarbons with absorption process Adsorption is a well-known and simple separation technique finding application in water treatment, specialty chemicals production, gas separation and removal of trace impurities.1–4 Although adsorbents have yet to be broadly adopted in the treatment of refinery liquid hydrocarbon streams, a cursory review of literature reveals that the field is very active. Engineers and scientists have gained remarkable levels of control of the properties of surfaces and interfaces leading to the molecular-level design of new materials. As capacity, adsorption kinetics and selectivity continue to improve, widespread adoption may soon be realized. In the case where a refiner wants to preserve the aromatic or olefin character of a stream, one promising application of adsorption is the removal or concentration of organo-sulfur or organo-nitrogen compounds. Traditional hydrotreating would saturate olefins and perhaps some aromatic compounds. In the case of zeolites, activated carbons and microporous coordination polymers, these aromatic compounds and olefins tend to compete for adsorption sites, typically leading to a decrease in sulfur (S) compound adsorption capacity. Little information exists on the efficacy of these aromatics and olefins with nitrogen (N2) compounds.5 One company has been working to exploit a molecular feature of its proprietary catalysta that has shown promising results for S and N2 compound separation in a way that does not saturate aromatics and olefins. The catalyst removes S and N2 compounds without the need for hydrogen (H2). Furthermore, with the growing importance of biofuels to the global oil supply, the need for S and N2 removal processes that accommodate a variety of functional groups in oil is clear. Here, the early results of an effort to develop a new separation processb are discussed. Separation process overview. The proprietary separation process uses two columns for the continuous adsorption processing of contaminated feeds. One column runs in adsorption mode, which removes heteroatom compounds from the liquid hydrocarbon stream. The other operates in regeneration mode, removing the adsorbed heteroatom compounds and regenerating the bed (FIG. 1). While in adsorption mode, the feed oil flows into the column over the catalyst adsorbent. Heteroatom compounds bind to the adsorbent, and the contaminant-free oil flows out as purified product.

In regeneration mode, oil feed is switched to the fresh column. The spent column is flushed with a solvent to remove residue treated oil, the solvent is recovered and recycled, and the residue oil is recycled back to the operating column. The adsorbed heteroatom compounds are then removed from the spent column by flushing it with a small amount of an inexpensive organic hydroperoxide. The heteroatom-rich stream flushed from the column is then vacuum distilled, and the concentrate is removed from the bottom of the recovery column. The byproducts of the process are an organic alcohol collected overhead and the heteroatom concentrate as column bottoms. Depending on the feed heteroatom content, the enriched concentrate may be burned as fuel, sent to a coker or sent to a fluid catalytic cracking unit (FCCU). As an alternative, the hydrocarbons can be reclaimed using a proprietary desulfurization and upgrading process.c Desulfurization and upgrading performance. To demonstrate the utility of the proprietary process to remove S from oils without the use of H2 , a variety of petroleum oils and intermediates were tested in a plug-flow column. Breakthrough curves (FIG. 2) for four different feeds demonstrate the S-removal capaOxidant/sulfones

Solvent/oil FlexULS product

Adsorption column (online)

Oil feed pump

Solvent/oil recovery step— blue

Sulfone removal step—red Alcohol byproduct

Adsorption column (regeneration)

Recovered oil

Sulfonerich stream

Solvent pump Oxidant

Makeup

Regeneration pump

FIG. 1. Separation process flow scheme. Hydrocarbon Processing | OCTOBER 2016 75

Process Engineering and Optimization bility of the process. TABLE 1 shows the results for the tested feeds. As shown, S compounds can be adsorbed from many different stream types, and the extent of S removal is feed-dependent. Analysis of the N2 content before and after treatment is shown in TABLE 2. The level of N2 removal in these feeds is unimpressive, with the only exception being the nearly 40% reduction of N2 from dicyclopentadiene (DCPD). The lack of significant N2 removal results in these feeds was surprising because experience with whole crudes and bitumen using the proprietary process produced 40%–70% N2 removal. 1.0

Cv/Co

0.8 0.6 Heavy white oil High-sulfur diesel Kerosine DCPD

0.4 0.2 0.0

0

10

20

30

40 Bed volumes

50

60

70

80

FIG. 2. Breakthrough curves for sulfur adsorption of various feeds using the proprietary desulfurization and upgrading process. 0.08 0.07 OPEX, $/gal

0.05 0.04 0.03 0.02 0.01 0.00 0

100

200

300

400

500 600 700 Sulfur removal, ppm

800

900

1,000

1,100

FIG. 3. OPEX vs. sulfur removal for various feeds (different sorbent capacities).

DCPD High-sulfur diesel Kerosine Heavy white oil

CAPEX, $MM

8

Petroleum feed

Feed S, ppm

Product S, ppm

Bed volumes

Capacity, mg S/g sorbent

Heavy white oil

1,050

62

6

9

High-sulfur diesel

504

146

10

3.8

Kerosine

104

4

55

5.5

Dicyclopentadiene (DCPD)

95

22

9

1.8

TABLE 2. Performance for N2 removal with proprietary process

12 10

Separation process economics. Process economic estimates were generated for an 18-Mbpd process treating a 100-ppm feed down to ultra-low-sulfur standards with 40 bed volumes, prior to the required regeneration step. The process equipment includes: • Two adsorption columns • Standard fractionation column equipment with a feed heat exchanger • Pumps. Capital expenditure (CAPEX) is estimated at $3,000,000 (inside battery limits, installed cost), and operational expenditure (OPEX) is estimated at $0.0058/gal (cost of makeups and utilities). The type of feed being treated influences the economics of the process. Lower S removal and higher sorbent capacity reduce costs, while the opposite conditions raise them. OPEX TABLE 1. Proprietary desulfurization and upgrading performance with different feeds

DCPD High-sulfur diesel Kerosine Heavy white oil

0.06

To further investigate this apparent discrepancy, model feeds containing indole, acridine and quinoline were prepared in hexadecane and similarly tested in a plug-flow column at ambient temperature. The model compounds were measured by high-performance liquid chromatography, so detection limits on the basis of elemental N2 were less than 1 ppm. Results are shown in TABLE 3. All of the N2 compounds were completely removed from the model feed through greater than 58 bed volumes. The efficacy for aromatic cyclic N2 compound adsorption appears to be high. Future work will focus on clarifying the generality of N2 compound adsorption by the proprietary separation process.

6 4

Petroleum feed

Feed N2, ppm

Product N2, ppm

Heavy white oil

36

33

High-sulfur diesel

29

26

Kerosine

36

33

DCPD

41

26

TABLE 3. Adsorption of aromatic heterocyclic N2 compounds

2 0 0

100

200

300

400

500 600 700 Sulfur removal, ppm

800

900

1,000

FIG. 4. Relationship between CAPEX and sulfur removal for various feeds.

76 OCTOBER 2016 | HydrocarbonProcessing.com

1,100

Initial N2, ppm

Product N2, ppm

Bed volumes

Capacity, mg N/g sorbent

Indole

25

<1

58

2.3

Acridine

22

<1

85

2.1

Quinoline

21

<1

> 92

>2

Compound

Process Engineering and Optimization is more sensitive to S removal than CAPEX and scales almost linearly with it. It is difficult to discern significant differences in OPEX for feeds tested, mainly due to the cost of makeup materials needed to regenerate the sorbent bed. Conversely, CAPEX shows more sensitivity to changes in sorbent capacity. The sorbent capacity directly influences the size of the fractionation equipment, which is the largest factor for CAPEX. FIGS. 3 and 4 highlight the influence of S removal and sorbent capacity on process economics. Takeaway. The proprietary separation process provides an

alternative for heteroatom removal to meet market needs for feeds where preservation of olefin or aromatic character may be important, or where H2 constraints exist. As the technology continues to mature and additional feeds are treated, improved clarity on which feeds offer the best economic advantages will emerge. NOTE The proprietary catalyst is Auterra’s FlexOX. b The proprietary separation process is Auterra’s FlexULS. c The proprietary desulfurization and upgrading process Auterra’s FlexUP. a

LITERATURE CITED Worch, E., Adsorption Technology in Water Treatment, Walter de Gruyter, Berlin, Germany, 2012. 2 Yang, R. T., Adsorbents: Fundamentals and Applications, John Wiley & Sons, Hoboken, New Jersey, 2003. 1

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Wu, L., J. Xiao, Y. Wu, S. Xian, G. Miao, H. Wang and Z. Li, Langmuir, pp. 1080– 1088, 2014. 4 Jia, S. Y., Y. F. Zhang, Y. Liu, F. X. Qin, H. T. Ren and S. H. Wu, Journal of Hazardous Materials, Vol. 262, pp. 589–597, 2013. 5 Cychosz, K. A., A. G. Wong-Foy and A. J. Matzger, Journal of American Chemical Society, Vol. 131, Iss. 40, 2009. 3

KYLE LITZ is the chief technology officer of Auterra Inc. He has over 22 years of experience in process chemistry and new materials development. Dr. Litz is the lead co-inventor of the FlexOX catalyst and the FlexUP process, and is responsible for Auterra’s technology development toward commercial introduction. He has more than 30 issued US patents co-filed in more than 30 countries, and he has been published in 12 peer-reviewed publications. He earned his BS degree in chemistry from the University of Texas at Dallas and his PhD in chemistry at the University of Michigan. In 2016, he was admitted as a fellow of the Royal Society of Chemistry. KEITH EDISON is the lead process engineer at Auterra Inc. He has been with Auterra for five years and is responsible for the evaluation, scale-up and optimization of various process designs surrounding Auterra’s proprietary technology. He has also designed, built and operated bench and pilot units needed for process development work. He earned a BS degree in chemical engineering from Rensselaer Polytechnic Institute. JONATHAN RANKIN is the research and development team leader at Auterra Inc., and has been with the company for seven years. He oversees day-to-day research and development activities at Auterra. To date, he has nine granted US patents in the areas of catalyst development and petroleum upgrading. Mr. Rankin obtained a BS degree in chemistry from Rensselaer Polytechnic Institute in 2009.

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Process Engineering and Optimization N. KASIRI and P. JOUYBANPOUR, Iran University of Science and Technology, Tehran, Iran; and M. REZA EHSANI, Isfahan University of Technology, Isfahan, Iran

The effect of various parameters on tray point efficiency Distillation as a volatilities-based process is used to separate components of a mixture to a liquid phase richer in less volatile components, and a vapor phase richer in more volatile components. Component separation is achieved through differences in boiling points between the species. In the chemical and petroleum industries, distillation is an important process for separating fluid mixtures, so a considerable amount of attention has been given to understanding and improving the performance of tray distillation columns.1 Tray efficiency is the most important parameter in the design and analysis of tray-type columns. Although tray efficiency has been studied extensively, little progress has been made in predicting it since the publication of the American Institute of Chemicals Engineers (AIChE) Bubble Tray Design Manual.2 Most existing tray efficiency models were developed based on experimental data of systems in which mass-transfer resistance is confined entirely to either the liquid phase or the gas phase. These models tend to predict a large value for the number of liquid-phase transfer units for distillation systems and, therefore, give a small value to the fraction of liquid-phase resistance over the total mass-transfer resistance. According to the AIChE model, researchers have assumed that distillation is a gas-phase-controlled operation. However, experimental data indicates that the liquid-phase resistance in distillation is also significant.1,2 The newly developed and presented model takes into account important parameters in tray point efficiency by using a genetic programming technique. Genetic programming is one of the computer algorithms in the family of evolutionary-computational methods that have provided reliable solutions to complex optimization problems. Genetic programming is a biologically inspired, domain-independent method that automatically creates a computer program from a high-level statement of a problem’s requirements.3,4 Study background. Many parameters influence distillation

column and tray point efficiency. To develop a correlation that included all influential and effective parameters, a comprehensive review of published works was conducted. The effective pa-

rameters on tray point efficiency were identified and combined into meaningful dimensionless groups, such as Reynolds number (R e ), the ratio of molar densities and molecular diffusivities in both phases, the ratio of the liquid inventory (hL ) to the sieve tray perforation diameter (DH ), and the effective froth density ϕe.3,5,6 Therefore, the slope of the equilibrium curve (m), Reynolds number (Re = ρVVHhfe/µ V), the ratio of vapor mole density to liquid mole density (ρMV/ρML), the ratio of vapor molecular diffusivity to liquid molecular diffusivity (DV/DL), the ratio of the liquid inventory to the sieve tray perforation diameter (hL/DH), and (1-ϕe ) are chosen as independent and dimensionless variables, with tray point efficiency (EOG) as the dependent variable. Genetic process automatically eliminates the superfluous independent variables from the model. The correlation is based on a generalized model of mass transfer in a five-sieve tray distillation column developed to account experimental measurements on a 0.6-m-diameter sieve tray with 0.0185-m-diameter perforations, indicating the presence of vapor entrainment into the downcomer for the three systems. The database consists of experimentation results under total reflux conditions for two wide-boiling systems (methanol/ water and isopropanol/water) and one close-boiling system (methylcyclohexane/toluene), all of binary nature. The entire database consisted of 52 data sets. On the basis of measured and calculated data, an attempt was made to determine an equation for the point efficiency of distillation columns using genetic TABLE 1. Dimensionless grouping validity ranges Parameters Slope of equilibrium line, m

Minimum

Maximum

Max./Min.

0.224

2.8

12.5

ρMV /ρML

0.0007

0.0047

6.714

DV /DL

1.9407 × 1.0e + 004

5.1902 × 1.0e + 004

2.674

hL /DH

3.6305

6.0872

1.677

1-ϕe

0.4533

0.6977

1.539

5.3755 × 1.0e + 004

2.2822 × 1.0e + 005

4.245

Re = ρVVHhfe /µV

Hydrocarbon Processing | OCTOBER 2016 79

Process Engineering and Optimization programming. In this work, 47 out of the total 52 data sets were randomly used to develop the desired correlation, and the remainder were used for error evaluations. The experiments were carried out under atmospheric pressure.7 Results and discussions. To develop the new correlation,

part of the available experimental data sets of three binary mixtures were used for the development phase, while the remainder were used for the validation and evaluation phases.7 The new developed correlation8 is shown in Eq. 1: ⎡ ⎤ ⎢ ⎥ ⎛ ρ MV ⎞ ⎢ ⎥ ⎜ ⎟+R e ⎢ ⎥ ⎝ ρ ML ⎠ ⎥ Eog =1−exp⎢−0.867789− (1) ⎢ (1−φe )(exp(m)+10.6125) ⎥ ⎢ ⎥ ⎛ DV ⎞ ⎢ ⎥ − (1−φe ) ⎟ ⎜2 ⎢ ⎥ ⎝ DL ⎠ ⎣ ⎦

The superfluous independent variable (hL/DH) was automatically eliminated during the genetic process and did not appear in the model. The average relative error associated with the model for predicting tray point efficiency was evaluated using Eq. 2: 0.70 1-Qe = 0.4533 1-Qe = 0.5133 1-Qe = 0.5733 1-Qe = 0.6333 1-Qe = 0.6977

0.69 0.68 Tray point efficiency

0.67 0.66

Error =

0.65

0.6448132

0.64

0.644813195

0.63

0.64481319 0.64098737

0.61

0.640987365

0.60 0.0

0.5

1.0

1.5 m

2.0

2.5

3.0

Tray point efficiency

0.80 m = 0.224 m = 0.824 m = 1.424 m = 2.124 m = 2.68

0.78 0.76 0.74

0.64098736 0.635578266

FIG. 1. The influence of the slope of the equilibrium curve (m) on tray point efficiency.

0.6635578264 0.6635578262 0.62848862 0.628488616 0.628488612 0.620065792

0.72

0.662006579

0.70

0.620065788

0.68

0.609822602

0.66

0.6098226

0.64

0.609822598 0.000

0.62 0.60 0.4

(2)

According to this equation, a relative error of 1.85% was measured when the output model was checked against the five data sets not used in the optimization procedure, and an error of 1.75% was measured when the output model was checked against the complete set of available data. TABLE 1 shows the coverage range of the dimensionless groups, according to the composite database. In this work, the Peng-Robinson equation of state was used to evaluate the densities of gases and liquids for methanol/water and isopropanol/water systems; Nationally Recognized Testing Laboratories (NRTL) was used solely for the methylcyclohexane/toluene system; and the Brokaw model for gases and the Hayduk-Laudie model for liquids were used to evaluate viscosity.9 The influences of the dimensionless groups on tray point efficiency were studied within the valid range of the model, varying one as variable and one as parameter, while the rest of the variables were constant. FIG. 1 shows the influence of the slope of the equilibrium curve (m) on tray point efficiency, with (1-ϕe ) as parameter (at DV /DL = 1.9407 × 1.0e + 004, ρMV/ρML = 0.0007, Re = 5.3755 × 1.0e + 004). When the slope of the equilibrium curve is increased, the mass-transfer resistance increases, but the molar rate of diffusion decreases, causing tray point efficiency to decrease. Also, by increasing (1-ϕe ), efficiency is decreased. When the effective froth density is decreased, it decreases the mass transfer and causes the separation to be more difficult. As observed in FIG. 1 at smaller values of m, (1-ϕe ) has a larger influence.

0.62

Tray point efficiency

n

1 ∑ E(OG , calculated)i − E(OG , experimental)i n i=1

0.6

0.8

1.0

1.2

Re

1.4

1.6

1.8

2.0

2.2 × 105

FIG. 2. The influence of Re on tray point efficiency, with m as parameter.

80 OCTOBER 2016 | HydrocarbonProcessing.com

0.001 m = 0.224 m = 0.724

0.002

0.003 m = 1.224 m = 1.724

0.004

0.005

m = 2.224 m = 2.8

FIG. 3. The influence of the ratio of vapor mole density to liquid mole density on tray point efficiency, with m as parameter.

Process Engineering and Optimization fusivities, the mass-transfer resistance is increased and efficiency is decreased. Also, the increase in Reynolds number causes a decrease in the mass-transfer resistance, so efficiency is increased. FIG. 5 details the influence of (1-ϕe ) on tray point efficiency with Re as parameter (at m = 0.224, DV/Dl = 1.9407 × 1.0e + 004, ρMV/ρML = 0.0007). As shown here, increasing the (1-ϕe ) factor decreases the efficiency. When the effective froth density is decreased, it decreases the mass transfer, meaning that distillation will be more 0.80 Re = 5.3755* 1.0e + 004 Re = 7.2755* 1.0e + 004 Re = 9.1755* 1.0e + 004 Re = 1.0755* 1.0e + 005 Re = 2.2822* 1.0e + 005

0.78 0.76 0.74 Tray point efficiency

The influence of Re on tray point efficiency with m as parameter (at DV/DL = 1.9407 × 1.0e + 004, ρMV/ρML = 0.0007, (1-ϕe ) = 0.6977) is illustrated in FIG. 2. The Reynolds number has the most influence on tray point efficiency, as its variations make larger changes on it. Increasing the Reynolds number causes a decrease in the mass-transfer resistance, thereby increasing tray point efficiency. Increasing the slope of the equilibrium curve increases the mass-transfer resistance, but the molar rate of diffusion is decreased, causing efficiency to decrease. It can also be concluded from FIG. 2 that at higher Re, m has a great influence on tray point efficiency. FIG 3. shows the influence of the ratio of vapor mole density to liquid mole density on tray point efficiency with m as parameter (at DV/DL = 1.9407 × 1.0e + 004, (1-ϕe ) = 0.6977, Re = 5.3755 × 1.0e + 004). As a consequence of increasing the ratio of the molar densities of the phases, the mass-transfer resistance is decreased, while the molar rate of diffusion and efficiency are increased. As observed in FIG. 3, this influence is infinitesimal and, therefore, negligible. The influence of the ratio of vapor molecular diffusivity to liquid molecular diffusivity on tray point efficiency with Re as parameter (at ρMV/ρML = 0.0007, m = 0.224, (1-ϕe ) = 0.6977) is shown in FIG. 4. The ratio of the diffusivities has an inverse influence on efficiency. These systems are limited by the liquid phase since the mass transfer is more efficiently accomplished there, so the liquid diffusivity has more influence. By increasing the ratio of the dif-

0.72 0.70 0.68 0.66 0.64 0.62 0.60 1.5

2.0

2.5

3.0

DV DL

3.5

4.0

4.5

5.0 × 104

FIG. 4. The influence of the ratio of vapor molecular diffusivity to liquid molecular diffusivity on tray point efficiency, with Re as parameter.

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Process Control Engineer Chevron

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Mike Rhodes

Managing Editor Hydrocarbon Processing Hydrocarbon Processing | OCTOBER 2016 81

Process Engineering and Optimization difficult. Also, increasing the Reynolds number causes a decrease in the mass transfer resistance, so efficiency is increased. Key findings. In this study, a genetic programming-based model was used to study the influence of different parameters on point efficiency in a distillation column. The model demonstrated 1.85% relative error within its validity range, which is adequate accuracy for distillation column analysis. To study the effect of any of the parameters, one was changed in 0.90

D DH EOG hf hfe hL m Re VH VS

Re = 5.3755* 1.0e + 004 Re = 7.2755* 1.0e + 004 Re = 9.1755* 1.0e + 004 Re = 1.0755* 1.0e + 005 Re = 2.2822* 1.0e + 005

0.85

Tray point efficiency

turn while the rest were kept constant. At the modeling stage, (hL /DH ) had minimal influence on point efficiency, which was eliminated automatically by the genetic process. The parameters that had high effect were (DV/DL ), (m), (1-ϕe ), Re and (ρMV/ρML ). As expected, the first three parameters had an inverse effect on point efficiency, while the last two had direct influence. Among these parameters, R e has the most influence on efficiency, while (ρMV/ρML ) has the least effect.

0.80 0.75

Subscripts e Effective f Froth V Vapor H Hole L Liquid M Molar.

0.70 0.65 0.60 0.40

0.45

0.05

0.55

1-e

0.60

0.65

0.70

NOMENCLATURE Molecular diffusivity, m2/s Sieve tray perforation diameter, m Tray point efficiency Froth height, m Effective froth height, m Liquid inventory on tray, expressed in height of clear liquid, m The slope of the equilibrium curve Reynolds number Vapor velocity through perforation, m/s Vapor velocity over bubbling surface of the tray, m/s.

0.75

Greek symbols: ρ Density, kg/m3 µ Viscosity, N.s/m2 ϕe Effective froth density.

FIG. 5. The influence of 1-ϕe on tray point efficiency with Re as parameter.

LITERATURE CITED Bjorn, I. N., U. Gren and F. Svensson, “Simulation and experimental study of intermediate heat exchange in a sieve tray distillation column,” Computers and Chemical Engineering, Vol. 26, 2002. 2 Chen, G. X. and K. T. Chuang, “Determining the number of gas-phase and liquidphase transfer units from point efficiencies in distillation,” Industrial & Engineering Chemistry Research, Vol. 33, 1994. 3 Koza, J. R., Genetic Programming: On the Programming of Computers by Means of Natural Selection, MIT Press (Massachusetts Institute of Technology), Cambridge, Massachusetts, 1992. 4 Koza, J. R., F. H. Bennett, D. Andre and M. A. Keane, “Synthesis of topology and sizing of analog electrical circuits by means of genetic programming,” Computer Methods in Applied Mechanics and Engineering, Vol. 186, 2000. 5 Bennett, D. L., D. Watson and M. A. Wiescinski, “New correlation for sieve-tray point efficiency, entrainment and section efficiency,” AIChE Journal, Vol. 43, 1997. 6 Bennett, D. L. and K. W. Kovak, “Optimize distillation column,” Chemical Engineering and Processing, May 2000. Complete literature cited available at HydrocarbonProcessing.com 1

DR. NOROLLAH KASIRI graduated with his BSc degree from Glamorgan University before pursuing an MSc degree and PhD at Swansea University in Wales, UK. Dr. Kasiri joined the School of Chemical Engineering at Iran University of Science and Technology (IUST) as an assistant professor, where he established the CAPE center. Over the past 20 years of CAPE activity, he has managed professional chemical, process and reservoir engineers, resulting in the presentation and publication of over 200 papers, the conclusion of 70 research projects and the development of 14 software packages. He works with IUST as an associate professor. PARVIN JOUYBANPOUR began her career in chemical engineering at the Science & Research Campus of Islamic Azad University, where she graduated with a BSc degree, followed by post-graduate studies at the CAPE center at Iran University of Science and Technology (IUST), where she earned her MSc degree. She earned a second MS degree in industrial management in Sweden. DR. MOHAMMAD REZA EHSANI serves as a professor in the chemical engineering department at Isfahan University of Technology (IUT). He began his career in chemical engineering at Sharif University of Technology, where he graduated with a BSc degree. Dr. Ehsani earned his MSc degree and PhD at UMIST University, Manchester, UK.

82 

Select 157 at www.HydrocarbonProcessing.com/RS

Maintenance and Reliability J. MURRAY, Emerson Process Management, Boston, Massachusetts

Monitor medium-voltage switchgear in refineries When medium-voltage switchgear (MVS) in a refinery fails, it can force a process unit or even an entire refinery to shut down. Costs can climb into the millions of dollars as a result of lost production, environmental issues and litigation arising from injuries or fatalities, and the cost of repairing and/ or replacing the damaged equipment. Fortunately, modern sensing technology makes it possible for a refinery to continuously monitor the health of MVS (FIG. 1) and related equipment, and inform the maintenance department when, or even before, problems arise. This article describes the potential causes and problems of MVS failures, as well as the various types of sensing systems that can be used to detect and predict problems. Continuous monitoring of temperature, partial discharge and humidity in switchgear can uncover small problems before they become large issues that can shut down an entire refinery. Scope of the problem. Refinery power outages have been in

the news over the last few years. In January, a sitewide power outage at ExxonMobil’s refinery in Beaumont, Texas required operators to empty systems by burning product, wasting 365 Mbpd of crude oil. The burning created large flames and plumes of black smoke that drifted over nearby neighborhoods.1 According to the Billings Gazette, the city of Billings, Montana has been pulled into an ongoing lawsuit between ExxonMobil and NorthWestern Energy that was filed by the oil corporation after two power outages disrupted its Billings refinery.2 In January, ExxonMobil sued NorthWestern for refinery outages in 2014 and 2016 that led to excess flaring. ExxonMobil claims the outages cost the company millions of dollars, and it is seeking undisclosed damages in the lawsuit. In February, a California federal jury found Pacific Gas and Electric (PG&E) negligent and partially to blame for a power outage at Tesoro’s Martinez, California refinery, and ordered PG&E to pay $3.5 MM in damages.3 Power outages at refineries are a serious problem, resulting in millions of dollars in damages, environmental lawsuits, flaring, smoke, fires, lost production, bad public relations and threats to employee safety. While not all of these outages are caused by MVS failures, the impacts of these types of failures can be similar. Refinery power concerns. Studies by ARC Advisory Group,

Hydrocarbon Publishing Co., and Harris and Williams assert

that 82% of power outages in refineries can be attributed to random failures. The majority of refinery power systems are more than 25 years old, with many running beyond design life. Electrical problems accounted for 20% of all refinery disruptions between 2009 and 2013, according to the US Department of Energy. While power failures can be caused by utility outages, snowstorms, hurricanes and other unforeseen events, these studies claim that nearly 20% of refinery power disruptions are the result of electrical power equipment failures. In many cases, these failures are caused by MVS. A typical refinery power system has generators, generator circuit breakers, transformers, MVS, bus ducts, low-voltage switchgear, motor control centers and other equipment needed to distribute power throughout the refinery. Power substations are located at the sources of incoming power and at various smaller distribution stations throughout the refinery. Diagnosing and detecting all potential problems in this myriad of equipment (such as generator vibration or transformer oil deterioration) is beyond the scope of this work. Instead, it will examine how to monitor three of the main sources of electrical failures in MVS: overheating of conductors, insulation breakdown and high-moisture environments. Note: While this article concentrates on MVS, much of the discussion also applies to the other electrical assets.

FIG. 1. Typical installation of MVS in a refinery. Hydrocarbon Processing | OCTOBER 2016 83

Maintenance and Reliability MVS issues. MVS is subject to overheating due to excessive loads, normal wear and tear, and challenging environmental conditions. Left unattended, these conditions can lead to failures that result in costly damage to switchgear and surrounding equipment, power production loss and, in extreme cases, severe injury or death. Common failure modes include excessive temperature, partial discharge and high humidity. Excessive temperature. Circuit breaker, bus bar and cable connections tend to loosen and/or corrode over time, resulting in thermal failure of the connection and nearby cable insulation. Partial discharge. As insulators age, weak spots and defects evolve. Under certain load conditions, a dielectric breakdown will initiate across the defect, causing a partial arc between conductors at different potentials. This effect is known as a partial discharge. The breakdown causes a small but sudden rise in current accompanied by a current pulse, as well as electromagnetic (radio or light), acoustic and ozone emissions. Left unattended, this condition can cause the switchgear to explode (FIG. 2). High humidity. Moisture in switchgear can create shorts or be absorbed by the insulators, leading to insulation breakdown. Humidity also causes metallic corrosion, which can lead to elevated heating, partial discharge, surface tracking and the potential for shorts and flashover. Manual inspections. Switchgear monitoring has been in

practice for some time, and is often carried out through periodic manual inspections while the switchgear is powered down. Such inspections look for obvious problems, such as physical damage, frayed connectors, degraded insulation and evidence of overheated components. Inspectors confirm proper alignment of primary and secondary interlocks, tightness of bolted connections and correct phasing of bus bars. Electrical measurements can also be conducted while the power is off. Applying voltage with calibrated AC and DC highpotential test sets checks insulation resistance in the panel enclosure, bus bars, circuit breakers and other components. It also assesses contact resistance to confirm bus bar joints are connected properly. Manual inspections with infrared (IR) equipment can be conducted while the power is on. Periodic IR monitoring techniques

FIG. 2. Partial discharge and high humidity caused voltage stresses in this switchgear, leading to an explosion.

84 OCTOBER 2016 | HydrocarbonProcessing.com

require a glass window to be installed in the switchgear, a relatively expensive IR camera and a trained technician. One significant limitation of this type of inspection is that personnel cannot perform monitoring procedures behind bus insulators and cable shrouds because line-of-sight is required with IR technology. All manual inspections require trained technicians and specialized test equipment, often provided by an outside service company. Depending on the testing service provided (e.g., thermal and/or partial discharge), typical costs for an onsite visit by a service provider could total up to $30,000. Any electrical problems occurring after the inspection can go undetected until the next inspection, which could be a year or more. During that time, small problems can become large ones, potentially leading to a complete failure of the asset(s) and a power shutdown. A better solution is to employ continuous monitoring of switchgear, which affords refineries the ability to collect data generated during the switchgear’s normal operating conditions, thereby providing awareness to problems in real time. Real-time trending during full load of electrical stresses, vibration, insulation breakdown and environmental influences provides insight into the health of the switchgear. When performing continuous monitoring, it is not always critical to identify the exact location of degradation, but rather to understand the trend of the defect over time. Monitoring and trending the most common failure modes allows for planned and proactive maintenance, instead of running the MVS to catastrophic failure. Thermal monitoring. Temperature monitoring is a primary

method for detecting switchgear problems, although some facilities employ only partial discharge monitoring. One challenge in implementing continuous temperature monitoring of critical connection points (circuit breaker, bus bar and cable connections) inside MVS is that the sensors must maintain the impulse-withstand voltage, also known as basic impulse level (BIL). Consequently, conductors at different potentials must have a minimal distance between one another to prevent breakdown and ensure the impulse rating. Another key challenge is powering the sensors so as to avoid the requirement for regular maintenance of these devices. Specific to the BIL concern, Section 5.2 of the Institute of Electrical and Electronics Engineers Standards Association’s C37.20.3 standard, “Metal-enclosed interrupter switchgear,” states that switchgear rated with a maximum voltage of 15 kV must have an impulse voltage of 95 kV, relating to a distance (in air) of approximately 160 mm. This requirement eliminates the most common types of direct contact temperature-monitoring systems, such as thermocouples and RTDs, because they are susceptible to electromagnetic interference/radio-frequency interference in the high-voltage environment. This leaves only non-invasive systems, such as fiber optics, continuous IR sensing and wireless direct contact sensors, as viable options for switchgear thermal monitoring. Fiber optic systems. Fiber optic temperature sensors can be routed directly to critical switchgear-monitoring points. The sensors are rigidly attached to hot-spot locations, and are completely immune to electromagnetic interference and noise bursts caused by high-voltage switching. These systems require direct surface contact. They are considered non-invasive, but when op-

Maintenance and Reliability gear out of service to replace batteries is untenable from an operational perspective, and also elevates the actual cost of system ownership to an unacceptable level. Wireless passive direct-contact sensors. Wireless passive sensor systems provide real-time continuous monitoring via direct connection to critical measurement points. These systems are easy to install, require no maintenance and have a life expectancy comparable to the switchgear itself. These sensors employ surface acoustic wave (SAW) technology (FIG. 3). SAW technology utilizes a piezoelectric substrate (a material that changes electrical charges due to mechanical stresses), an interdigitated transducer (IDT) resonator and an antenna. One of the most common piezoelectric materials used is quartz crystal, onto which a metalized IDT is fabricated. SAW transducers, when activated by a radio frequency (RF) wave transmitted by a system control device, reflect back a surface RF wave that changes frequency linearly with temperature. Compared to other sensors, SAW temperature sensors have longer reading distances. They also do not require line-of-sight. These characteristics maximize basic impulse levels, while using lower transmission power to minimize control system electromagnetic compatibility issues.

erating under the dusty and humid environments often found in refineries, fiber optic cables can create a conductive ground path and compromise the aforementioned withstand voltages. Continuous IR sensing. IR sensors can detect “hot spots” in switchgear, cables and components. Monitoring switchgear with handheld IR equipment during manual inspections is common, but IR sensors can also be permanently installed inside switchgear for continuous monitoring. The sensors capture a thermal image and then send data back to a control or SCADA system, where specialized thermal mapping software is required to detect problems. These systems are often expensive and difficult to install, and have several measurement limitations, such as adjacent surfaces having different emissivity or reflections causing false readings. These sensors also cannot measure the bus connections behind bus insulators. Wireless battery-powered direct-contact sensors. Wireless temperature sensors measure the temperature of strategically important points on metal or insulating surfaces of current-carrying parts. The sensors are attached to components with a high-temperature adhesive. They measure temperatures from 0°C to 150°C and send data to a receiver, which then transmits the data to a control system. At a first glance, these systems seem to provide an adequate solution. Unfortunately, the batteries have a limited life span that is usually reduced in the high-temperature environments found inside MVS. The associated cost of taking the switch-

Partial discharge monitoring. The most commonly used

partial discharge detection instruments directly measure the current and voltage spikes in high-frequency current transformers or high-voltage capacitive couplers, as outlined in the Inter-

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„„

 

™ 

Maintenance and Reliability national Electrotechnical Commission’s IEC 60270 standard, “Partial discharge measurements.” This method has several strengths, including the ability to analyze pulse shapes and assemble a graph of the discharge events relative to the phase of the power line waveform. These systems are very expensive and require trained technicians to

FIG. 3. Three SAW temperature sensors (orange sensors in middle) mounted in MVS.

analyze the data, and do not lend themselves to permanent, continuous monitoring installations in switchgear. Utilities are evaluating additional partial discharge detection methods with IEC 62478, a prospective standard for acoustic and electromagnetic PD measurements. These methods use instruments to make indirect analytical measurements and obtain a relative signature of partial discharge pulses that testers can use for system trending. A number of testing methods exist, as outlined in the following sections. Transient earth voltage (TEV). Refinery or outside service personnel use TEV test equipment to check switchgear externally by measuring the electromagnetic emissions conducted to ground. Although testers can use TEVs for continuous monitoring, these voltages cannot always monitor faults between phases. High frequency (HF) and very-high frequency (VHF). HF/VHF systems operate between 3 MHz and 300 MHz and use large antennas, high-frequency current transformers or coupled sensors to detect partial discharge spikes. Large antennas are not compatible with measurements inside switchgear, and direct-coupled sensors impair safety. Acoustics. Testers use a microphone or similar acoustic sensor to monitor frequencies between 10 Hz and 300 kHz. Refinery personnel can use this wireless method for continuous monitoring, but it has a limited detection range due to sound damping within dielectric objects. Ultra-high frequency (UHF). This broadband (300 MHz– 3 GHz) detection method monitors transient electromagnetic waves through an antenna. Traditional UHF methods are susceptible to noise from cell phones, radios and other transmitters. However, newer instruments use selective, banded UHF monitoring to detect partial discharges while rejecting noise sources. UHF provides the safest, most non-intrusive continuous partial discharge monitoring system (FIG. 4). UHF partial discharge detection. Effective UHF partial

discharge detection for continuous monitoring requires the distillation of an overwhelming amount of complex data down to a concise piece of information, all without the intervention of a highly trained operator. Band-pass filtered UHF partial discharge detection methods are capable of avoiding strong interfering signals at close proximity—while advanced digital filtering methods are capable of

26

100

22

80

18

60 14

86 OCTOBER 2016 | HydrocarbonProcessing.com

40

10 0

FIG. 4. A monitoring device installed in a safe position (black box on inside panel) within an MVS compartment monitors partial discharge.

RH, %

nC cumulative/cycle

30

Ongoing partial discharge Periodic partial discharge Humidity

FIG. 5. Continuous trending of partial discharge events captures periodic spikes due to condensing humidity that would not have been recognized with non-continuous detection methods.

20

Maintenance and Reliability detecting the presence of partial discharge—even with residual, unfiltered noise. To provide an autonomous approach for PD monitoring, advanced system algorithms must be implemented to present data that can be easily processed for health assessment and long-term system trending. UHF emissions signals can be broken into three major categories: noise, and asymmetric and symmetric discharges. Noise. Noise denotes UHF energy in the selected frequency band(s) not in close correlation with the power line frequency. External radio interference is a reliable classification of noise; however, weak and erratic partial discharge, which occurs early in the evolution of a defect, is also a noise classification. Asymmetric discharge. Events occur primarily on the negative half cycle of the power waveform, where electrons emitted from the metal ionize the air. Asymmetric discharge (also called corona or surface discharge) presents significant results at both odd and even harmonics of the power line frequency. Symmetric discharge. Discharge events that occur in the bulk of the material, often referred to as internal or symmetric partial discharge, happen at the positive and negative polarity portions of the power cycle and are often represented as even harmonics of the power line frequency. One of the causes of partial discharge is excessive humidity, which condenses onto cables and connectors. Continuous UHF monitoring can detect these voltage spikes (FIG. 5). Continuous monitoring. Complete switchgear continuousmonitoring systems are available with temperature, humidity and partial discharge sensing capabilities. In addition to the wireless temperature and partial discharge sensors outlined above, wired sensors are often installed on the chassis of the switchgear cabinet to provide ambient temperature and humidity readings. Ambient temperature readings are important because the critical issue is temperature rise of hot spots above ambient, as opposed to absolute temperature. A typical continuous-monitoring system for MVS would include a monitoring unit connecting temperature, humidity and partial discharge sensors, as shown in FIG. 6. Each temperature and partial discharge air interface device shown in FIG. 6 uses banded UHF technology to sense partial discharge directly. Each air interface device can also wirelessly link to three or more SAW temperature sensors. Up to four air interfaces can be wired to the monitoring unit via low-loss coaxial cables. The monitoring unit can also accept up to eight conventional wired humidity or ambient temperature sensors, which are well suited for taking measurements of these variables in bus ducts. The monitoring unit can be a full-featured human-machine interface (HMI) with monitoring capabilities, or a reader that provides remote monitoring. The reader provides all the necessary wireless interrogation signals for the SAW sensors through the air interface device, internally implements the partial discharge detection algorithms and communicates directly with the humidity and ambient temperature sensors. All reader data is accessible through industry-standard Modbus remote terminal unit (RTU) (RS485) digital communication protocols, affording ease of integration into a plant’s existing supervisory control and data acquisition system, distributed control system or historian.

Monitoring unit (CAM-4 or reader)

Humidity 1

Temperature and partial discharge air interface

Temperature sensor

Temperature sensor

Temperature sensor

Temperature and partial discharge air interface

Temperature sensor

Temperature sensor

Temperature sensor

Temperature and partial discharge air interface

Temperature sensor

Temperature sensor

Temperature sensor

Temperature and partial discharge air interface

Temperature sensor

Temperature sensor

Temperature sensor

Humidity 2

Humidity 8

FIG. 6. Switchgear monitoring system with temperature, humidity and partial discharge sensors.

The HMI unit provides the same monitoring capabilities as a reader, but adds functionalities including a local touchscreen for real-time display, data storage, alarming and multi-unit functionality. Multi-unit functionality provides an input device port for connecting up to seven readers via the Modbus RTU protocol, as well as extending the HMI display, alarm and communication capabilities to all connected readers. This option allows for easy integration into switchgear lineups. The HMI unit can be connected to an existing plant SCADA system or historian via Modbus transmission control protocol (TCP), distributed network protocol 3 (DNP3) or IEC-61850 digital communications. Takeaway. Continuous condition-based monitoring systems provide clear, early detection of evolving issues with MVS, thereby helping avoid unplanned downtime and outages. Data provided by a condition-based monitoring system facilitates historical trending capabilities, as well as the ability to detect failures before they occur, giving refinery operators the information they need to perform proactive maintenance. A continuous thermal and partial discharge monitoring system that focuses on main switchgear failure modes will help lower maintenance costs, reduce downtime, provide necessary early warning information and ensure the highest level of operational safety. LITERATURE CITED Kang, Y. P., “Jury dings PG&E $3.5M for Tesoro refinery outage,” Law360, February 2, 2016, online: http://www.law360.com/articles/754290/jury-dingspg-e-3-5m-for-tesoro-refinery-outage/ 2 Besson, E. and B. Scott, “Smoke at Exxon refinery in Beaumont due to power outage,” FuelFix, January 21, 2016, online: http://fuelfix.com/blog/2016/01/21/ flames-smoke-seen-at-exxon-refinery-in-beaumont/ 3 Hudson, M., “NorthWestern: Fault began with city employees in Exxon refinery power outages in Billings,” Billings Gazette, April 5, 2016, online: http://billingsgazette.com/news/local/northwestern-fault-began-with-city-employees-inexxon-refinery-power/article_c30dd4ca-0dc4-5ba9-b809-2f9a6fd83bd0.html 1

JONATHAN MURRAY is director of IntelliSAW products at Emerson, focusing on new business strategies, product development and marketing of the IntelliSAW Critical Asset Monitoring (CAM) Platform. Prior to joining IntelliSAW, Mr. Murray focused on automation and control systems for smart grid, energy storage and electric vehicles, holding positions such as business unit manager, business development manager, project manager and senior engineer. He holds a BS degree in electrical engineering from the University of Massachusetts. Hydrocarbon Processing | OCTOBER 2016 87

Process Control and Instrumentation B. JAZAYERI, Reacxion, Orange County, California

Lessons learned in commercial scale-up of new chemical processes Commercializing a new chemical process can be as simple as installing one or more homogenous batch reactor(s), or as complex as designing a fully integrated chemical complex requiring one or more heterogeneous reaction steps processing gas, liquid and/or solids, with other units required to prepare feeds, recover products/byproducts and recycle streams. The latter is focused on here, with further clarification that there are always exceptions to the rules. Typically, process scale-up evolves from lab scale to pilot, demo and commercial (FIG. 1). The lab scale is usually limited to studying the reactor and catalyst performance. The pilot plant should be a scaleddown version of the commercial process configuration to the greatest extent feasible. The pilot is used to confirm/expand reactor and catalyst performance data and to test the balance of plant concepts. The demo stage is usually used when large quantities of product are needed for performance testing by end users. Know the reaction chemistry. The reaction chemistry must be well-advanced at the lab scale. However, having a developed idea of the range of selectivity, yield and potential byproducts may be sufficient to proceed to the next step, such as piloting, where these items can be further solidified. Often, the right group of people with varied backgrounds can brainstorm and produce engineering solutions to mitigate this risk. Every effort should be made to obtain data under conditions anticipated for the commercial unit—e.g., pressure, temperature and gas/solid residence time. 88 OCTOBER 2016 | HydrocarbonProcessing.com

A common mistake is to use lowpressure units to study the chemistry of a high-pressure process. This decision may transpire either because the reactor cost is lower, or because the low-pressure unit already exists. Note that tests in a lab require personnel and a host of equipment and instruments, of which the reactor is only one cost component. Therefore, the lifecycle savings from using a low-pressure reactor is a small fraction of the total cost of the program. What is the drawback of using a lowpressure unit? In one example, a partial oxidation reaction converts a hydrocarbon to an oxygenated main product containing carbon, oxygen and hydrogen, and byproducts consisting of water and carbon oxides. Extensive tests were conducted in a readyto-use low-pressure unit. Concerns about catalyst deactivation in the presence of high partial pressure of the hydrocarbon feed were addressed by raising the concentration of the hydrocarbon in the feed, effectively reducing the partial pressure of other species. No deactivation was noticed, and very high yields were obtained. A larger unit was built and operated at actual conditions, with significantly lower yields and measurable catalyst deactivation. These results were due to several factors: • Yield of this partial oxidation reaction decreased as pressure was raised. • Reducing the partial pressure of the other reaction species resulted in easier desorption of these species in the lab unit, making catalyst sites more readily available for the main reaction path. Under actual pressure, it was more difficult for these

other species to desorb, resulting in both reduced yield and more rapid catalyst deactivation. Such oversights are far more common than published, occurring even within R&D at major corporations. Knowing the catalyst. By the time the process is commercialized, the catalyst formulation, shape, size distribution, porosity, attrition properties, deactivation rate, and (if required) reactivation method and rate, must be known. But what about the mid-step involving pilot or demo units? Does everything about the catalyst need to be known to design these units? The answer is: sometimes, no. Pilot and demo plants can be designed to test catalyst formulations that are vastly different with respect to activity and size. For example, a new process used catalyst circulation between a reaction zone and a regeneration zone, similar to that practiced in refinery fluid catalytic cracking units (FCCUs). Several formulations of catalyst were being developed with a wide range of reactivity and attrition properties. It was estimated that the time span needed to develop a final catalyst formulation was approximately the same as that needed to design and build a larger unit. Rather than waiting for the catalyst formulation to be finalized, the project to design a larger unit was undertaken. The challenge was posed to design a flexible unit to handle a range of catalyst formulations with low to high activities, resulting in the approximate concurrent operating ranges: • 2:1 turndown in operating pressure • 3:1 turndown in reactor gas residence

Process Control and Instrumentation • 4:1 turndown in solid circulation • 2:1 turndown in auxiliary vessel solid holdup • 2:1 turndown in regenerator solid holdup. Readers familiar with FCCU design and operation know that these are challenging requirements; however, they were successfully met. The solutions increased the pilot plant’s cost somewhat, but they also measurably reduced time to market. A different pilot plant project was designed to operate with both fine and coarse catalysts, while research proceeded in parallel to determine which option to follow. Readers familiar with fluidization know that this is not an easy design. One of the key mistakes made in the early stages of development is that the rapid replacement of catalyst can mask potential long-term catalyst deactivation. Another is limited run length (i.e., not operating continuously). Run length is especially critical when chain reactions can result in molecular growth that causes vaporizable products to turn into non-vaporizing highboiling-point compounds, or when material deposition can result in size growth. Stepwise vs. scale-up/scale-down commercialization. Two extreme ap-

proaches exist in taking a process from the laboratory to the commercial stage. Both are practiced, and both can lead to success. In the stepwise process, only the scale at hand is considered. For example, imagine that a 1-tpd plant is built and operated, followed by a 10-tpd unit and finally a 100-tpd commercial unit. No effort is made to think about the next step at each stage. Startup companies that rely heavily on government funding often use this model due to funding limitations imposed by these organizations. In a simplistic way, with the scale-up/ scale-down approaches, the commercial scale is always being examined. A concept plant for the commercial unit is designed to evaluate the kinds of challenges that scale may impose on engineering and design. The commercial unit is then scaled down, and the lab unit is scaled up. This approach, when performed by experienced personnel, will quickly identify ways in which scale will impact design, what elements can and cannot be piloted in a practical manner, and what elements must be addressed using other methods, such as cold-model testing.

The author’s experience suggests that this approach can reduce time to market by months or even years. A number of actual examples of potential drawbacks to the stepwise approach exist: • A 12-in.-inside-diameter (ID) moving-bed waste-conversion reactor using oxygen or enriched air was piloted successfully. The same concept on a 10-ft-ID commercial-scale (roughly 100:1 capacity scale) may pose a serious heat removal challenge from the central section of the reactor, possibly requiring a completely different design and raising questions about the applicability of data collected on the smaller scale to date to the larger scale (yield, selectivity, etc.) (TABLE 1). The technology developer can opt to use multiple trains to keep reactor size small in the commercial plant, but this carries a negative economic impact that will not be faced until the commercial design stage is reached. • Gasification, combustion and many other processes produce solid byproducts (slag, clinkers, spent solids, etc.) that require gravity removal from the reaction zone. In a small unit, a few lb are removed daily and often dropped into a 50-gal container that is sealed on top and purged with nitrogen. The container is then emptied under safe procedures on a regular basis. Pressure letdown and heat dissipation of material removed occurs in the oversized container. On a commercial scale, the solids removal may be thousands of lb/hr, requiring a large train of equipment

Start

to cool, depressurize and inert the material at a significant cost. Often, this added equipment also raises the elevation of the reactor, adding more cost due to a taller structure, increased pipe run length and other elements. These modifications pose a significant negative impact on process economics that will not be identified until the commercial design is started, possibly putting the project in jeopardy. • Startup and shutdown operations make up another variable with potentially significant negative economic impacts, and must be studied at the pilot scale. In the lab, an inert gas is passed through a temperature-controlled heater to either heat or cool the reactor. This once-through approach is infeasible on commercial scale, often requiring the addition of a dedicated recycle compressor loop and associated equipment, with negative economic impacts on the process. Sequential vs. parallel engineering. Ideally, all of the information needed to commit to a pilot plant, demo plant or commercial plant is known when the decision is made. Sometimes, this is not the case. Some part(s) of the process are not as well defined. However, pressures caused by the market’s window of opportunity, fund shortages or other factors make it necessary to commit to the start of design. The risk associated with this situation can be mitigated when prior art can be relied upon to develop a design. More than one solution is often feasible, and each solution must be developed, cost-estimated, technically assessed and risk-ranked. The most common approach is to rank all options based on a team assessment after

R&D stage

Pilot plant stage

Demo plant stage

3 to 120 months

9 to 36 months

24 to 48 months

Commercial plant

FIG. 1. Typical step sequence in scale-up. Note: Pilot and/or demo stages are optional.

TABLE 1. Typical range of reactor parameters as a function of scale Scale

Lab

Pilot

Demo

Commercial

Inside diameter, in.

<3

3 to 14

6 to 36

> 24

Weight rate, lb/hr

<30

10 to 1,000

400 to 2,000

> 1,500

Hydrocarbon Processing | OCTOBER 2016 89

Process Control and Instrumentation one or more brainstorming sessions, and then develop and cost out the first (or best team-assessed) option. The evaluation is ended if the first option is found to be technically and economically feasible. If it is not, then the next option is assessed, and so on. This process can be termed “sequential engineering.” The risk with the sequential approach is extended time. In “parallel engineering,” more than one option is examined concurrently, cost-estimated and technically assessed. The risk with the parallel approach is waste of labor if the first option turns out to be acceptable. It is important to remember that these activities are being performed while a task force of tens to hundreds of engineers and designers are spending money developing design for other parts of the process. The author favors the parallel approach, as it limits the time to solution at a limited increase in engineering cost. Process engineers with front-end engineering and design (FEED) experience are comfortable with either approach. Unfortunately, other engineering disciplines and project management often favor the sequential approach, as they are asked to focus on one task at a time. Backup option. Engineers are often

faced with choosing between multiple options, none of which can be ruled out. As an example, multiple vendors may have equipment components that appear to be technically suitable for a certain step in the process. Unfortunately, very large quantities of material are needed to test each unit on a continuous basis under actual conditions, which requires commitment to a large-capacity, expensive demo plant. This cost may not be justifiable if the only reason is to produce the required quantity of material for such a test. In this situation, it is feasible to select the top two options, design the unit based on the first, and allow flexibility in the design to replace the system with the second option. Obviously, this process will require a shutdown and added costs. A team of experts and engineers are used to select the top two options.

Total system design. New chemical

processes often involve new reactor concepts. Unfortunately, the reactor often becomes the sole focus of research, with

90 OCTOBER 2016 | HydrocarbonProcessing.com

little thought given to the balance of the plant. This can be detrimental to the design. Consider the common scenario of the startup company with limited cash struggling to fix the back end of its inadequately designed pilot unit so that it can operate the new reactor concept, prove the concept and raise more funds. Meanwhile, a crew of operators are sitting idle, burning cash. The lesson here is that a successful chemical facility requires a robust design across all sections. The author recommends using the total system design approach, selecting contractor(s)/ consultant(s) with proven applicable experience to at least provide the process design package, and not using “low cost” as the sole criteria for the selection of consulting/engineering firms for design. Cost growth. It is common to see an-

nouncements by startups on how a new route will reduce the production cost of an existing process, usually by a good margin, and often from a reduction in installed plant cost. Variable costs associated with chemical processes such as raw materials, utilities, operations, maintenance, waste disposal and product sales are generally easy to estimate and with a high level of accuracy, based on process yield and selectivity. Capital costs are much more difficult to assess. Most startups do not have the funding resources to hire an engineering firm for this purpose, and so they must rely on either public domain information or cost-estimating software. Public domain information issued by government entities should be used with caution. This information is usually acceptable if used to rank different schemes. However, it should not be used to estimate production and/or plant costs, nor for comparing new process data with non-governmental published data for existing commercial plants run by competitors. The result often presents an unrealistically favorable comparison. Cost-estimating software programs can be useful and are generally reasonably accurate; however, they can produce misleading data. Often, it is unclear what pieces of information may require manual adjustment to obtain a realistic output. Engineering contracting firms rely heavily on these tools for initial cost estimates, but they also invest considerable effort to continuously benchmark their estimates

against actual purchases to identify where manual adjustments may be needed. These two factors can often lead to cost growth for the commercial plant between early-stage estimates and actual estimates obtained during the design phase by contractors, eliminating most, if not all, of the perceived economic benefit of the new process. The other cost growth comes from incorrect early-stage design. Omissions may result in missed equipment, as discussed in some of the aforementioned examples. Continuous engineering. Engineering

should be part of the commercialization effort; this is the essence of the scale-up/ scale-down approach. The engineer(s) leading this effort must have the necessary experience and skills to conceptualize the entire plant with a high level of realism, starting from day one, with occasional reassessments to follow. Good engineers can identify gaps in research knowledge early on and ask R&D to obtain the solutions before the gaps become a critical path to pilot or commercial design efforts. If early efforts show viable economics, then it is highly unlikely that the final results at the commercialization stage will be different. The converse is also true. An early negative result from engineering may not be insurmountable if viewed as a challenge for improvement. For example, a negative result can be turned positive with a small increase in reactor yield, which can then guide further research. In another example, a major producer invented a process to produce a certain compound in high global demand. The front and back ends of this new process were fairly standard and easy to estimate. The initial concept for the “reactor” made the process economically unattractive. The consultant started with the published market price of the product, deducted required margins and variable costs, and came up with a contribution to the selling price due to depreciation of capital. The maximum permissible total plant cost was estimated from these calculations. Deducting the known cost of the balance of the plant yielded the required “imputed” cost for the reactor. The client’s team then worked with alternate concepts to see if the imputed cost was achievable. The technology is now in the advanced stages of demonstration on a large scale,

Process Control with a reactor concept quite different from the initial one. Fast-tracking new technology. A final

example shows how experience coupled with a capable and focused team effort can yield fantastic results with new technology development. A large producer discovered that a catalyst used in another application could be used to address new global regulations on a high-demand product, with good results. Many producers had already committed to using alternates in meeting these new regulations. This producer had a short window of opportunity to demonstrate its process for use in its own facilities and for possible license to others. A small amount of data was obtained in the lab, providing a feasible range of operation, including pressure range and weight hourly space velocity (WHSV), or lb/hr of feed per lb of catalyst. The producer assembled a large research group, selected an engineering firm and asked for a plant to be started within what may have been considered an unrealistically short time frame. Weekly meetings were held between research and engineering. Research updated engineering on the new discoveries made that required tweaks to the initial design concepts. Engineering advised research on how it planned to design sections of the plant lacking data, and asked research to confirm the acceptability (but not necessarily the optimality) of the design by testing, with target dates set for their responses. Certain design aspects were challenging to address in the lab, and were resolved using cold model testing. The plant was ultimately built within the allocated time frame and proved successful. It also yielded many global licensees.

ACKNOWLEDGMENT The author wishes to thank R. Sandel and Steve Fusselman for their valuable inputs. BEHZAD JAZAYERI has more than 20 years of experience in process engineering, with a focus on the engineering of new chemical technologies to commercial success. He has designed many pilot plants and six first commercial plants, and has performed numerous techno-economic assessments of existing and emerging process technologies. Mr. Jazayeri is an expert in the design of fluidized bed and fluid-solid systems and has extensive background in gasification, combustion, oxycombustion, bioconversion, alternative energy, gas-to-liquids, chemicals, petrochemicals and solar-grade polysilicon. He has authored four publications and coauthored one patent.

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Process Control and Instrumentation M. A. ALÓS, Inprocess Technology and Consulting Group, Barcelona, Spain

Use dynamic simulation to maximize plant operating performance The oil and gas industry has witnessed seismic downward shifts in barrel prices and uplifts in market competitiveness. Owner-operators are challenged to squeeze more from their operations, and engineering, procurement and construction companies (EPCs) are under increased pressure to reduce project risk and achieve better alignment with owneroperators to ensure effective management of their clients’ capital. Companies must maximize operating performance and produce quality products faster. The use of rigorous dynamic simulation helps achieve these objectives. Analyzing operating scenarios using dynamic process modeling gives owner-operators confidence that new plants will start up safely, meet budgets and perform to plan. Dynamic simulation is vital to successful plant design and operation. Designs can be improved by incorporating a system’s dynamic response to changes within the design model. To accommodate and study scenarios like shutdowns and emergency situations, dynamic process simulation of the “real world” delivers effective results.

to confirm a safe and efficient design. The net result to a project is significant time and cost savings. The pillars of an effective use of steadystate and dynamic simulation to support engineering design best practices are based on good communication, efficient knowledge transfer and comprehensive support training. Tailoring training programs to maximize process simulation knowledge helps impart a better understanding of the plant lifecycle. It also maximizes project return on investment (ROI). Benefits of dynamic simulation. Unlike steady-state simulation, dynamic models factor time, which helps model complex transient behavior (i.e., changes of temperature, pressure, etc.). This type of simulation reproduces real behaviors of a process plant, providing values that correspond to process variables at a given time. These simulations possess an interactive interface to help the engineer understand the demands of working inside an actual plant.

The main benefit of dynamic simulation is the deeper knowledge of the process it provides as a result of improvements in system control design, plant operations and staff training. It enables the verification of the appropriate size of equipment used to determine design constraints covering the plant’s normal operation. As the plant undergoes many modifications during its lifetime, dynamic simulation provides a means of continuous assessment of the operability of the proposed design solution. Better design decisions through detailed analysis enable engineers to make necessary trade-offs and optimize the design. Dynamic simulation makes it possible to simulate a realistic understanding and behavior of the plant, and helps improve decision support and safety. In summary, adopting dynamic simulation ensures: • Reduced project risk, with increased cost accuracy • Improved plant operability

Gaining reputation. In both the refin-

ing and upstream oil and gas sectors, lost production overshadows installation costs, so it is vital for plant designs to be robust to minimize downtime. Dynamic simulation can show transient responses that are not determined during traditional steady-state design methods, providing engineers with a deeper understanding of plant operational behavior, which can be crucial in safety-related equipment sizing. Process simulation knowledge providers can conduct dynamic simulation studies

FIG. 1. Dynamic simulation flowsheet. Hydrocarbon Processing | OCTOBER 2016 93

Process Control and Instrumentation Best practice. In a recent case study,

models helped the client across different phases of the facility’s lifecycles. In phase one, during which an advanced simulation platform1 was used, firstprinciple models (steady-state Effective use of steady-state and dynamic simulation supports and dynamic) provided a sigengineering design best practice, good communication, efficient nificantly better understanding of the process dynamics and knowledge transfer and comprehensive support training. interactions. These models also enabled the process simu• Troubleshooting for process upsets based in Central America. The engineers lation engineers to evaluate and fine-tune • Insight for startup and first plant were asked to verify and perform a feasi- strategies before implementation. Using operation before actual plant bility study on engineering design mod- dynamic simulation early in the design commissioning els supplied by an international EPC. phase allowed the engineers to identify • Safer designs and identification Verification of the EPC’s design formed important operability and control issues, of potentially undersized the project’s first phase. The second along with equipment sizing adjustments, equipment (pumps, compressors, phase involved verification of the multi- which led to design improvements. In the second phase, the simulation valves, etc.) before commissioning phase pipeline behavior, assessing plant and first startup startup and support of plant operability platform integrated all relevant aspects of plant operation into one simulation envi• Assessment of existing facilities to to reach production targets. check if they can accommodate new The project involved analysis of all ronment, reducing the time required for production volumes operating variables, including safety “what-if ” studies (FIG. 1). During startup • Comprehensive learning resources analysis, process dynamics and unit in- operations, a holistic simulation model for improved workforce expertise. teractions. The application of simulation was used that incorporated both the pipelines and the structure of the platform (topside). The model was able to determine the correct sequence of steps for starting up the initial wells and reaching the early production target rate. The design of pipeline systems requires complex considerations, such as the pipeline flow path, terrain profiles, expected volumes to be received and delivered, and the physical properties of products. The dynamic simulation of pipelines with updated compositions from the wells conveyed crucial information about the development of liquid holdup profiles, as well as pressure and temperature data. This information was used to determine predicted hydrate formation risk and slug catcher sizes. The FIG. 2. Process trainer for operators. dynamic modeling results also helped prevent interruptions to the plant and avoided reduced profits. Substantial savings were achieved on the project through improved and faster startup procedures. The over-design of relief systems was avoided, which yielded capital cost savings. The topside model was tuned and calibrated by reconciling discrepancies with plant data. This was achieved by adjusting certain values and parameters to ensure that the model corresponded to actual unit operation before proceeding with further development. The availability of the upgraded model helped the client troubleshoot operFIG. 3. Dynamic simulation lifecycle. ability issues and reach targets earlier. • Reduced number of plant shutdowns with an improved control narrative

94 OCTOBER 2016 | HydrocarbonProcessing.com

process simulation engineers provided consultation to a large owner-operator

Process Control and Instrumentation In addition, using a common platform facilitated effective communication between all stakeholders, which was essential for knowledge transfer. Collaboration and information-sharing became the lifeblood for determining accurate outcomes and analyzing plant behavior. The process simulation engineering company also delivered a process trainer for the plant (FIG. 2). This updated dynamic model contains real plant data. It also features a user-friendly interface that is identical to the plant human-machine interface. The trainer facilitates the execution of preprogramed scenarios of different operational conditions to troubleshoot detected problems. In the early field production phase, this operational support yielded a database of operational cases that were cumulative during the plant lifecycle. Consequently, newly hired engineers and operators are able to learn how the plant operates by following the simulation cases. This support also helps minimize the impact of personnel turnover on plant operability and safety (FIG. 3).

Confidence in operational outcomes. An effective design and control

strategy requires a comprehensive understanding of the process to successfully avoid unplanned downtime. The use of cutting-edge software is essential to support engineers through key stages of a project and assist with knowledge, collaboration and learning. Process simulation experts can now quickly create robust models to validate EPC design and increase confidence that projects will run on time and to standard, and be easily maintained, resulting in optimized operations. Dynamic plant models allow companies to achieve faster and safer plant startups while maximizing productivity. Dynamic simulation avoids disruptions and minimizes the impacts of unnecessary costs on real plant operations. For example, simulation models help owner-operators ensure that production goals are achieved according to already signed contracts, thereby avoiding penalties for not meeting agreed production or product quality.

Similarly, owner-operators can plan for suboptimal plant operating points, thereby preventing full plant shutdowns and keeping the plant in operation. Two examples of suboptimal operation are a partial power failure in a remote location and unstable power supply. Using dynamic simulation, owneroperators and EPCs can enjoy tangible benefits that deliver a realistic understanding of plant behavior and, ultimately, yield significant savings. 1

NOTE AspenTech’s Aspen HYSYS simulation platform.

MIQUEL ANGEL ALÓS is the services manager at Inprocess. He has more than 20 years of experience in the modeling of oil and gas and chemical processes, and in the delivering of educational training programs for process simulation. Before joining Inprocess in 2007, Dr. Alós worked at AspenTech, where he led simulation projects for the support and teaching of process simulation users. Dr. Alós holds a PhD in chemical engineering from the Ramon Llull University of Barcelona. He also lectured on reaction engineering and process simulation courses at the university from 1997 to 2001.

INDIA 17–19 April 2017, Taj Palace Hotel, New Delhi

IRPC India Call for Abstracts is Now Open Join us for IRPC India hosted by Indian Oil and supported by Petrofed.

For eight years, IRPC has been the leading downstream technology event. In 2017, IRPC Europe and IRPC Americas join IRPC India to provide refining and petrochemical professionals from around the globe, events that highlight technical innovation, share regional market insights and provide exceptional opportunities to network with the industry’s finest. Gulf Publishing Company and Hydrocarbon Processing invite you to take part in this market-leading event by submitting an abstract for consideration. Suggested topics and areas of interest for IRPC India 2017 include: • Alternative Feed Stock Fuels • Olefin and Aromatic Developments • Clean Fuels/Renewables • Water treatment/Management • Catalyst Developments • FCC Strategies • Heavy-Oil Conversion • Best Practices • Emerging Technologies • Optimization and Profitability • Bottom of the Barrel Upgrading

Hosted by:

Supported by:

The call for abstracts closes 1 December.

For a complete list of abstract topics, please visit HPIRPC.com/India

Hydrocarbon Processing | OCTOBER 2016 95

BOB ANDREW, TECHNICAL EDITOR [email protected]

Innovations Condition-based prognostics for reliability

availability, and it enables maintenance planners to make better decisions.

Where large machines are in use and the overall processes depend on a smoothly functioning plant, companies often rely on fixed maintenance schedules and diagnostic tools. However, the field of predictive analytics is gaining in relevance, as it sends early warning signals and can anticipate malfunctions before they occur. Predictive analytics provide diagnostic information about the condition of the plant and clues for the root cause analysis. However, predictive analytic approaches remain vague about the future. Cassantec AG eliminates this vagueness by calculating explicit forecasts and probabilities when malfunctions will occur. The Cassantec Prognostics algorithm (FIG. 1) helps businesses adjust maintenance management to the needs of the plant. It does so by reinterpreting the collected data so that the time windows for possible malfunctions are known well in advance. A change in the maintenance strategy is reflected in lower maintenance costs and higher system

Refinery example. The average maintenance costs for a 200-Mbpd oil refinery are approximately $37 MM. By using Cassantec Prognostics, a cost reduction of up to $2 MM is possible. The operator can save up to $1 MM more during regular plant turnarounds. A reduction in downtime of around 30% can be achieved, and plant availability can be improved. In these ways, companies reduce additional operating costs. Expectations get closer to reality. The results show that use of the prognostic approach makes it possible to reduce costs by an average of up to 30%. The methodology is particularly suitable for plants and machinery with components that are constantly in operation and show signs of wear and tear over time. Both increased plant availability and better maintenance measures provide for long-term cost savings in overall operations management. Bundling of maintenance interventions and avoiding unnecessary maintenance work provide for further reductions

Condition data

Current and historical

Pro his tor ica l

Onsite experience

Alarm function

Proprietary computational model

Select 1 at www.HydrocarbonProcessing.com/RS

Total water reuse in refining GE announced that Federated CoOperatives Ltd.’s Co-op Refinery Complex (CRC) in Regina, Saskatchewan, Canada is installing GE’s advanced water recycling technology for a wastewater improvement project. The project will enable the refinery to clean 100% of its wastewater onsite. Once fully operationally, CRC will be the only refinery in North America to recycle all wastewater for steam production. Steam is used for heating, hydrogen production, equipment power and cooling towers. Several years ago, the refinery expanded its operations to produce an additional 30 Mbpd, taking the facility’s total capacity to 130 Mbpd, which increased its water usage. CRC’s water source is a blend of well water and city water, and restrictions on water use required CRC to find a new source of water. GE offered a solution combining ZeeWeed (FIG. 2) membrane bioreactor (MBR) technology and a high-efficiency reverse osmosis (HERO) system to recycle and reuse 2 MMgal/d of wastewater. In addition to the water reuse solution, GE provides the refinery with wastewater

Prognostic reports Generation Aggregation

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in operating costs. In addition, companies acquire a tool that protects them from making critical decisions based on guesswork or “gut feeling.”

Availability forecasts

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Commercial

FIG. 1. Cassantec Prognostics helps businesses adjust maintenance management to plant needs.

96 OCTOBER 2016 | HydrocarbonProcessing.com

FIG. 2. The wastewater recycling solution combines a membrane bioreactor and a high-efficiency reverse osmosis system.

Innovations specialty chemicals and monitoring solutions to provide system optimization. After commissioning, the refinery will reduce its use of freshwater by 28% on an annual basis, which is the equivalent of approximately 3,100 households in Regina. By recycling 100% of its wastewater onsite, CRC will significantly decrease volatile organic compound emissions from wastewater ponds and reduce associated nuisance odors. The wastewater improvement project is expected to be fully operational by autumn 2016. Select 2 at www.HydrocarbonProcessing.com/RS

Passively cooled walk-in shelter Intertec has developed an innovative approach to housing remote instrumentation and communications equipment in harsh environments, in the form of a passively cooled walk-in shelter (FIG. 3). The shelter can reduce the problems of installing equipment in remote locations where reliable power is unavailable, and where dust and sand in the atmosphere can make it difficult to cool electronics equipment using conventional air conditioning systems. Another major element of the shelter’s performance is its construction from glass-reinforced polyester (GRP) panels employing a composite “sandwich” construction to provide a high degree of insulation, plus surface protection that can survive the extreme challenges of the Middle East climate and environment. These challenges include high levels of ultraviolet rays, dust and sand abrasion. GRP is an inherently inert material that is virtually immune to corrosion and atmospheric pollutants. It is also resistant to a wide range of petrochemical media. Intertec’s shelter employs an efficient passive cooling system that exploits the energy storage capacity of water, which circulates by natural convection. This passive, unpowered system can be boosted by a small, externally mounted electrical cooler driven by solar panels that optimize performance on hot sunny days. The high levels of insulation of Intertec’s shelters can substantially reduce the total cooling power required compared with insulated steel shelters, and provide stable operating environments for sensitive equipment, such as analyzers. Intertec’s composite GRP sandwich panels

include thick polyurethane insulation layers, which are bonded inside GRP sheets. This style of fabrication and assembly eliminates the “thermal shortcuts” between the shelter interior and the exterior that can result from the fixings that are often used with traditional insulated metal constructions. Multi-function composite material ensures that internal walls are smooth and stable, making it simple to mount equipment.

ing process. Using Stanhope-Seta’s latest technology, a test typically takes less than 2 min. to perform, requiring just 2 ml of test sample. For easy record-keeping, the instrument has a 1-GB memory that stores up to 100,000 test results. The new Setaflash instruments are designed to provide users with a cost-effective and simple way to perform flash point tests in-house, avoiding the expense and delay of outsourcing this service.

Small-sample volume flash point testing

Expanded hydrotreating catalysts portfolio

Flash point is an important parameter in fuel specifications. Testing supports check for quality and specification compliance, safety and transport regulations, batch consistency and contamination. Typically, tests are made at refineries, terminals, storage and distribution facilities, and test labs. Storage, transport and disposal charges are based on the flammability of the product; hazard classification is a legal requirement to warn of potential risk and to ensure that correct handling precautions are taken. Changes in flash point can also indicate that a sample may have been contaminated or adulterated. The new range of Setaflash Series 3 instruments (FIG. 4) make flash point testing possible for even those with minimum operator experience to quickly and reliably perform. Compact and ruggedly designed, the instruments are suitable for laboratory or portable tests and ideal for use where space is limited. Only a small amount of sample is required to perform the test, reducing the cost and waste per test. The new instrument features a simple user interface with color digital display and touchscreen icons that guide the operator through a straightforward test-

Honeywell UOP has introduced an expanded portfolio of new hydrotreating catalysts used to remove impurities and contaminants from petroleum and other refining feedstocks to produce cleanerburning gasoline and diesel that meets new global emissions regulations. The addition of a range of hydrotreating catalysts expands Honeywell UOP’s line of catalysts, which are used to produce transportation fuels and petrochemicals. Hydrotreating is a critical step in the refining process, where hydrogen and proprietary catalysts are used to pretreat petroleum and other products to remove sulfur, nitrogen, metals and other contaminants before conversion into transportation fuels. Hydrotreating helps produce cleaner-burning gasoline and diesel to meet increasingly stringent global fuel regulations, including Euro 5, China 5, and Bharat Stage 6—all of which specify sulfur content of less than 10 ppm in transportation fuels. Honeywell UOP’s new offerings include more than two dozen hydrotreating

Select 3 at www.HydrocarbonProcessing.com/RS

FIG. 3. The passively cooled walk-in shelter reduces the problems of installing equipment in remote locations with harsh climates and environments.

Select 4 at www.HydrocarbonProcessing.com/RS

FIG. 4. The new instrument line makes flash point testing quick and reliable. Hydrocarbon Processing | OCTOBER 2016 97

Innovations Dissolved gas analysis for transformers

FIG. 5. Advanced dissolved gas analysis methods are used for assessing large power transformers.

FIG. 6. The GARO AB4000R liquid ring compressor is designed for high pressure and performance requirements.

catalysts for applications, including hydrocracking and fluid catalytic cracking (FCC) pretreat, diesel and kerosine hydrotreating and coker naphtha hydrotreating. The new catalysts will be produced at Honeywell UOP’s production facility in Shreveport, Louisiana, which in June inaugurated new and upgraded production facilities to produce the catalysts. With the introduction of the new catalysts, Honeywell UOP is ending an alliance with Albemarle that began in 2006 when the two companies partnered to provide hydroprocessing technologies. While the alliance was a success for both companies, Honeywell UOP will now apply its expertise in catalytic chemistry to compete across a wide range of hydroprocessing technologies, while completing the work started with Albemarle on projects initiated under the alliance. Select 5 at www.HydrocarbonProcessing.com/RS

98 OCTOBER 2016 | HydrocarbonProcessing.com

Advanced dissolved gas analysis (DGA) methods are commonly used for assessing large power transformers (FIG. 5) in the international power generation and transmission sector. However, there is still room for improvement to their reliability and cost-effectiveness. Vaisala, a company working in environmental and industrial measurement, has highlighted that false alarms from onsite DGA systems and errors during routine oil inspections still pose challenges for asset owners. Utilities can tackle these issues with solid real-time condition monitoring, which leads to reduced maintenance costs and improved grid reliability and financial performance. Reducing the probability of error. The

financial severity of transformer failure is well-known to the industry. An extensive service costs more than $100 M, while a replacement transformer can cost up to $4 MM. Associated loss of production can further increase these figures. As a result, online DGA condition monitoring is becoming standard practice for aging transformer fleets. It enables proactive diagnosis and mitigation of developing faults before they lead to costly downtime. However, streamlining monitoring procedures and reducing the probability of errors is highly important. False alarms from onsite DGA monitors disrupt utilities’ maintenance schedules and generate unforeseen costs. In addition, the need to validate onsite data via regular laboratory sampling increases the complexity and duration of the testing process. Combined, these factors highlight a growing requirement for a dependable DGA monitor that not only provides 24-hr, real-time, online access to critical performance data, but also eliminates false alarms and cuts potential error. Vaisala’s Optimus DGA monitor for transformers has been designed to fulfill this requirement.

Dependable real-time monitoring. The robust plug-and-play Optimus DGA monitor can be installed in less than 2 hr, and uses partial vacuum gas extraction to provide a fully representative sample of all dissolved gases in the transformer oil. Readings taken by the built-in infrared

sensor are unaffected by oil temperature, pressure or type, and the system regularly auto-calibrates to provide consistent and dependable data. This data is available via an online interface, allowing continuous, reliable, real-time analysis of transformer performance without false alarms. Select 6 at www.HydrocarbonProcessing.com/RS

Improved efficiency liquid ring compressor GARO has expanded its two-stage compressor series with the debut of the AB4000R liquid ring compressor, which features a new frame. The GARO AB4000R offers up to 7% higher efficiency over competing liquid ring compressors in this range, and delivers proven reliability and performance. The GARO AB4000R liquid ring compressor operates at a pressure of up to 12.5 bar abs (160 psig) and is specifically designed for the higher pressures and performance requirements of oil, gas and chemical process applications. These applications include flare gas recovery, vapor recovery units, corrosive gas handling (e.g., vinyl chloride monomer and chlorine) and H2S gas sweetening. Designed for severe service, the rugged construction and easy-to-service design of the GARO AB4000R liquid ring compressor is said to minimize maintenance and reduce downtime for service and repair. Performance and features of the AB4000R liquid ring compressor include: • Speed: 740 rpm–900 rpm • Pressure: 4.5 bar abs–12.5 bar abs (50 psig–160 psig) • Capacity: 2,700 m3/hr–4,000 m3/ hr (1,500 cfm to 2,300 cfm) • High-efficiency performance • Robust construction designed for severe service • Simple design allowing for low maintenance and reduced downtime • Available in carbon steel, low-temperature carbon steel, austenitic stainless steel 316 (CF3M) and duplex stainless steel; or, on request, alloy 825 titanium and other materials • Suitable for oil splash, pure oil mist, purge oil mist, oil circulation and air-oil forced lubrication. Select 7 at www.HydrocarbonProcessing.com/RS

MARKETPLACE / [email protected] / +1 (972) 816-3534

Why Should You Filter Your Water?

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Scale  formation  reduces  the  heat  transfer  rate  and increases  the  water  pressure  drop  through  the  heat exchanger  and  pipes.  In  fact,  one  study  has  shownthat .002" fouling will increase pumping needs by 20%.    

The Best Engineered Water Filteration Solution Always Costs Less 2672 S La Cienega Blvd, Los Angeles, CA 90034 USA (800) 226-1942  (310) 839-2828   Fax (310) 839-6878 www.tekleen.com    [email protected]

Select 202 at www.HydrocarbonProcessing.com/RS

Select 201 at www.HydrocarbonProcessing.com/RS

SURPLUS GAS PROCESSING/REFINING EQUIPMENT 25 MMCFD x 1100 PSIG PROPAK REFRIGERATION PLANT 28 TPD SELECTOX SULFUR RECOVERY UNIT 1100 BPD LPG CONTACTOR x 7.5 GPM CAUSTIC REGEN NGL/LPG PLANTS: 10–600 MMCFD AMINE PLANTS: 60–3300 GPM SULFUR PLANTS: 10–180 TPD FRACTIONATION: 1000–25,000 BPD HELIUM RECOVERY: 75 & 80 MMCFD NITROGEN REJECTION: 25–100 MMCFD MANY OTHER REFINING/GAS PROCESSING UNITS We offer engineered surplus equipment solutions.

Epoxy Resists Harsh Chemicals

4.25" wide x 4" high Epoxy Compound EP41S-1HT CMYK color

Bexar Energy Holdings, Inc.

Phone 210-342-7106 • Fax 210-223-0018 www.bexarenergy.com • Email: [email protected]

Select 203 at www.HydrocarbonProcessing.com/RS

The Operator’s Role in Achieving Equipment Reliability Series Presented by Heinz Bloch • Part 1: Introduction and Overview, and why there is no reliability without operator involvement • Part 2: Common Misunderstandings with Equipment Reliability Impact • Part 3: Avoiding Machinery Failures Price: $595

GulfPub.com / +1 713-520-4426

Designed to Meet Specific Application Requirements • Fuel, alcohol & solvent resistant • Serviceable from -60°F to +400°F

+1.201.343.8983 • [email protected]

www.masterbond.com Select 204 at www.HydrocarbonProcessing.com/RS

Visit HydrocarbonProcessing.com for daily news, trends and FREE e-newsletters. Select 205 at www.HydrocarbonProcessing.com/RS Hydrocarbon Processing | OCTOBER 201699

ADVERTISER INDEX  /  HydrocarbonProcessing.com The first number after the company name is the page on which an advertisement appears. The second number is the Reader Service Number. There are two ways readers can obtain product and service information: go to www.HydrocarbonProcessing.com/RS, follow the instructions on the screen, and your request will be forwarded for immediate action, or go online to the advertiser's website listed below.

Company

Page

RS#

Website

ABB Inc.............................................................. 37

(53) (73)

www.info.hotims.com/61392-73

ARCA Regler GmbH.............................................43

(156)

www.info.hotims.com/61392-156

Axens .............................................................. 104

(51)

www.info.hotims.com/61392-51

Bluebeam Software Inc....................................... 14 Burckhardt Compression AG ..................................6

(97)

www.info.hotims.com/61392-97

CB&I ..............................................................S–60

(58)

www.info.hotims.com/61392-58

Dyna-Therm .......................................................13

(153)

www.info.hotims.com/61392-153

Emerson Process Management ..............................2 Enersul .......................................................... S–62

Page

RS#

Website

www.info.hotims.com/61392-53

Ametek Process Instruments ............................... 16

Company

HP Marketplace...............................................99 HP Webcast—Aspentech.................................. 81 HP Webcast—Shell.......................................... 32 HP Webcast—Simsci........................................47 HPI Market Data 2017................................... S–70 Software—Instrucalc.......................................30 Texas Pipelines Map ........................................85 Honeywell Process Solutions .................................5

(79)

Gastech .............................................................78 Gulf Publishing Company Construction Boxscore Database ......................28 Circulation...................................................... 77 Events—IRPC .................................................95

Linde AG ............................................................20 Maire Tecnimont SpA ..........................................29

Petrotechnics .....................................................44 (71) (157) (81)

www.info.hotims.com/61392-159

(63)

www.info.hotims.com/61392-63

Prognost Systems GmbH..................................... 91

(158)

www.info.hotims.com/61392-158

Prosernat ...................................................... S–59

(60)

www.info.hotims.com/61392-60

Rosen Swiss AG .................................................. 33 (62)

(61)

www.info.hotims.com/61392-61

Shell Research Ltd ..............................................26 (155)

Silcotek ..............................................................12

(152)

www.info.hotims.com/61392-152

(84)

Spraying Systems Co .......................................... 22

(159)

ZymeFlow Decon Technology .............................. 53

www.info.hotims.com/61392-84

OHL ............................................................... S–67

(151)

Petroleum Economist Ltd ....................................92

www.info.hotims.com/61392-155

Merichem Company............................................ 18

(78)

www.info.hotims.com/61392-151

www.info.hotims.com/61392-62

www.info.hotims.com/61392-79

Paqell ...........................................................S–68 Pentair ...............................................................11

www.info.hotims.com/61392-81

Kobelco Compressors America, Inc .....................103

RS#

www.info.hotims.com/61392-78

www.info.hotims.com/61392-157

KBC Advanced Technologies Inc...........................49

Page

Website

www.info.hotims.com/61392-71

Idrojet ...............................................................82

Company

(67)

www.info.hotims.com/61392-67

(93)

www.info.hotims.com/61392-93

This Index and procedure for securing additional information is provided as a service to Hydrocarbon Processing advertisers and a convenience to our readers. Gulf Publishing Company is not responsible for omissions or errors.

Catherine Watkins, Publisher Phone: +1 (713) 520-4421 E-mail: [email protected] www.HydrocarbonProcessing.com SALES OFFICES—NORTH AMERICA IL, LA, MO, OK, TX Josh Mayer Phone: +1 (972) 816-6745 E-mail: [email protected] AK, AL, AR, AZ, CA, CO, FL, GA, HI, IA, ID, IN, KS, KY, MI, MN, MS, MT, ND, NE, NM, NV, OR, SD, TN, TX, UT, WA, WI, WY, WESTERN CANADA Ryan Akbar Phone: +1 (713) 520-4449 Mobile: +1 (713) 504-9695 E-mail: [email protected] CT, DC, DE, MA, MD, ME, NC, NH, NJ, NY, OH, PA, RI, SC, VA, VT, WV, EASTERN CANADA Merrie Lynch Phone: +1 (617) 357-8190, Mobile: +1 (617) 594-4943 E-mail: [email protected] SALES OFFICES—EUROPE FRANCE, GREECE, SPAIN, PORTUGAL, SOUTHERN BELGIUM, LUXEMBOURG, SWITZERLAND, GERMANY, AUSTRIA, TURKEY Hamilton Pearman Phone: +33 608 310 575 [email protected]

100 OCTOBER 2016 | HydrocarbonProcessing.com

ITALY, EASTERN EUROPE Fabio Potestá Mediapoint & Communications SRL Phone: +39 (010) 570-4948 E-mail: [email protected] RUSSIA/FSU Lilia Fedotova Anik International & Co. Ltd. Phone: +7 (495) 628-10-333 E-mail: [email protected] UNITED KINGDOM/SCANDINAVIA, NORTHERN BELGIUM, THE NETHERLANDS Michael Brown Phone: +44 161 440 0854 Mobile: +44 79866 34646 E-mail: [email protected] SALES OFFICES—OTHER AREAS AFRICA Tanya Mbaluli Twiga Media Partner Phone: +254 722 376 972 Email: [email protected] CHINA—Hong Kong Iris Yuen Phone: +86 13802701367 (China) Phone: +852 69185500 (Hong Kong) E-mail: [email protected] INDIA Manav Kanwar Phone: +91-22-2837 7070/71/72 Mobile: +91-98673 67374 E-mail: [email protected]

INDONESIA, MALAYSIA, SINGAPORE, THAILAND Peggy Thay Publicitas Singapore Pte Ltd Phone: +65 6836-2272 E-mail: [email protected] JAPAN—Tokyo Yoshinori Ikeda Pacific Business Inc. Phone: +81 (3) 3661-6138 E-mail: [email protected] KOREA Young-Seoh Chinn JES Media, Inc. Phone: +82 (2) 481-3411/3 E-mail: [email protected] MEXICO, CENTRAL AMERICA, SOUTH AMERICA Marco Antonio Monteiro Mobile: +55 21 99616-4347 E-mail: [email protected] CLASSIFIED SALES Gerry Mayer Phone: +1 (972) 816-3534 E-mail: [email protected] DATA PRODUCTS J’Nette Davis-Nichols Phone: +1 (713) 520-4426 E-mail: [email protected] REPRINTS Rhonda Brown, Foster Printing Service Phone: +1 (866) 879-9144 ext. 194 E-mail: [email protected]

ALISSA LEETON, CONTRIBUTING EDITOR [email protected]

Events OCTOBER Yokogawa Users Conference and Exhibition, Oct. 3–6, Renaissance Orlando at SeaWorld Hotel, Orlando, Florida www.yokogawausers conference.com High Horsepower (HHP) Summit, Oct. 11–13, McCormick Place, Chicago, Illinois P: 888-993-0302 [email protected] www.hhpsummit.com NAPE, Oct. 12–13, Colorado Convention Center, Denver, Colorado P: +1 817-847-7700 [email protected] www.napeexpo.com National Safety Council (NSC), Oct. 15–21, Anaheim Convention Center, Anaheim, California P: +1 630-285-1121 [email protected] www.congress.nsc.org ACC Annual Meeting, Oct. 16–19, Moscone Center, San Francisco, California P: +1 202-293-4103 www.acc.com Canadian Society of Chemical Engineering (CSChE), Oct. 16–19, Québec City Convention Centre, Québec, Québec City P: 418-644-4000 www.csche2016.com Gasification and Syngas Technologies Conference, Oct. 16–19, The Westin Bayshore Hotel, Vancouver, British Columbia P: 703-503-0738 www.gasification-syngas.org 49th GOMA Symposium, Oct. 19–21, Jure Hotel, Šibenik, Croatia P: +385 1-48-73-549 [email protected] www.fuels.goma.hr RIO Oil & Gas 2016 Expo and Conference, Oct. 24–27, Rio de Janeiro, Brazil P: +55 21-2112-9080 [email protected] www.riooilgas.com.br/en

Emerson Global Users Exchange, Oct. 24–28, Austin Convention Center, Austin, Texas EmersonExchange@ Emerson.com www.emersonexchange.org/ americas/ LARTC 5th Annual Meeting, Oct. 25–27, Mexico City, Mexico P: +44 0-20-7384-8022 [email protected] www.lartc.events.gtforum.com/

NOVEMBER Women’s Global Leadership Conference, Gulf Publishing Company Events, Nov. 1–2, Hyatt Regency Houston, Houston, Texas WGLconference.com (See box for contact information) The Abu Dhabi International Petroleum Exhibition & Conference (ADIPEC), Nov. 7–10, Abu Dhabi National Exhibition Centre P: +971 0-2-6970-500 [email protected] www.adipec.com Sulphur 2016 International Conference & Exhibition, Nov. 7–10, Hilton London Metropole, London, England P: +44 0-20-7903-2444 [email protected] www.crugroup.com International Society of Automation (ISA) Process Control and Safety Symposium, Nov. 7–11, Houston Marriott Westchase, Houston, Texas P: 919-549-8411 [email protected] www.isa.org API 11th Annual Cybersecurity Conference for the Oil & Natural Gas Industry, Nov. 9–10, Westin Houston Memorial City, Houston, Texas (See box for contact information) AFPM International Lubricants & Waxes Meeting, Nov. 10–11, Hilton Post Oak, Houston, Texas P: 202-457-0480 [email protected] www.afpm.org

AIChE Annual Meeting, Nov. 13–18, Hilton San Francisco Union Square, San Francisco, California P: 800-242-4363 www.aiche.org ERTC 21st Annual Meeting, Nov. 14–16, Epic Sana Hotel, Lisbon, Portugal P: +971 0-55-307-3332 [email protected] www.gtforum.com API Fall Refining and Equipment Standards Meeting, Nov. 14–17, Hyatt Regency New Orleans, New Orleans, Louisiana (See box for contact information) European Autumn Gas Conference, Nov. 15–17, The Hague, Netherlands P: +44 0-20-3772-6080 [email protected] www.theeagc.com Latin American Petrochemical and Chemical Association, Nov. 19–22, Sheraton Buenos Aires Hotel & Convention Center, Bueno Aires, Argentina P: +54 11-4325-0086 [email protected] www.apla.com.ar CIS Downstream Summit, Nov. 28–30, The Ritz-Carlton Hotel, Vienna, Austria P: +44 0-207-384-7980 www.cis-downstream.com Valve World Expo & Conference, Nov. 29–Dec. 1, Fairground Düsseldorf, Dusseldorf, Germany P: +1 312-781-5180 [email protected] www.valveworldexpo.com

DECEMBER

FEBRUARY 2017 Egypt Petroleum Show (EGYPS), Feb. 14–16, CICEC, Cairo Egypt P: +971 0-4445-3726 [email protected] www.egyptpetroleumshow.com

MARCH 2017 Corrosion 2017, Mar. 27–30, Ernest N. Morial Convention Center, New Orleans, Louisiana [email protected] www.nacecorrosion.org

APRIL 2017 Gastech Conference & Exhibition, April 4–7, Makuhari Messe International Convention Complex, Tokyo, Japan P: +44 0-203-772-6086 [email protected] www.gastechevent.com

JUNE 2017 Global Petroleum Show, June 13–15, Stampede Park, Calgary, Alberta, Canada P: +1 403-209-3555 [email protected] www.globalpetroleumshow.com

AUGUST 2017 HCSMRP 11TH Annual Maintenance and Reliability Symposium, Aug. 9–11, Moody Gardens, Galveston, Texas P: 281-452-9800 [email protected] www.hunterbuildings.com

Center for Chemical Process Safety (CCPS) Global Summit on Process Safety, Dec. 4–5, InterContinental Al Jubail, Jubail, Kingdom of Saudi Arabia [email protected] www.aiche.org

Hydrocarbon Processing/ Gulf Publishing Company Events P: +1 713-520-4475 [email protected] [email protected]

Power Generation Gas Turbine User Group, Dec. 6–7, Thinktank, Birmingham P: +44 0-207-973-1251 [email protected] www.imeche.org

American Petroleum Institute (API) P: +1 202-682-8195 [email protected] www.api.org

Hydrocarbon Processing | OCTOBER 2016 101

MIKE RHODES, MANAGING EDITOR [email protected]

People

Dr. Christoph Wegner will become president of BASF’s Information Services and Supply Chain Operations division, effective December 1, 2016. He succeeds Dr. Robert Blackburn, who is leaving the company. Dr. Wegner joined BASF in 1997 as a research scientist, and has since served in numerous management positions, including his most recent role as senior VP of BASF’s Regional Business Unit Amines Europe. Victoria Pope has joined Gulf Publishing Company as its global events director. After 12 years in the events industry, where she was responsible for strategy and management of global events in the downstream oil industry and financial markets, Ms. Pope will oversee a portfolio of international events in the global upstream, midstream and downstream markets. She will also work closely with Gulf’s information brands Hydrocarbon Processing, World Oil, Gas Processing and Petroleum Economist. JV Driver Group has expanded its US leadership team with the transfer of Todd Robinson from its corporate HQ in Alberta, Canada to the firm’s Gulf Coast office in Houston, Texas. Mr. Robinson will be responsible for oversight of all project operations, including both office and field services.

Siluria Technologies, a provider of process technologies for the energy and petrochemical industries, has appointed Robert Trout as its CEO. He will bring 30 years of experience to his new role, as well as to his new place on Siluria’s board of directors. Prior to joining Siluria, Mr. Trout spent his entire career at Royal Dutch Shell in numerous leadership roles. Most recently, he served as the president of Criterion Catalysts & Technologies, where he was responsible for the overall integrated strategy related to Shell’s catalyst and licensing businesses, and was directly accountable for the global refining catalyst business. aeSolutions global functional safety consultant Paul Gruhn has been appointed as co-chair to the ISA84 Committee. As co-chair, he will help build consensus among participants on ISA84 standards interpretations and modifications, and ensure that the committee operates within its scope and purpose. Mr. Gruhn is an ISA Life Fellow, a member of the ISA 84 standards committee for 26 years, a developer and instructor of ISA courses on safety systems, the author of two ISA textbooks, and the developer of the first commercial safety system software modeling program.

102 OCTOBER 2016 | HydrocarbonProcessing.com

The Measurement, Control & Automation Association (MCAA) has named Teresa Sebring as its new president. Ms. Sebring joined MCAA in 1997 as administrative assistant to the president. Since 2012, she has served as VP and has been responsible for MCAA operations that include communications, data programs and staff oversight, marketing and meeting/event planning, as well as membership development, retention and strategic planning. Pöyry PLC is reducing its group executive committee from 10 to seven members. Anja McAlister has been appointed head of transformation and strategy, and will lead the HR team on an interim basis. Richard Pinnock is now executive VP of the merging energy business group and global sales and project management. Juuso Pajunen has been appointed executive VP and CFO of Pöyry PLC. Martin à Porta will assume the role of president and CEO, and chairman of regional operations. Nicholas Oksanen is now an executive VP and the president of the industry business group, while Erik Olsson serves as an executive VP and the president of the management consulting business group. Pasi Tolppanen has been appointed as an executive VP and vice chairman of Northern Europe regional operations.

Stephen Williamson has assumed the role of president and chairman of the board for T.D. Williamson, a Tulsa, Oklahoma-based pipeline and services company. He will succeed his brother, Richard B. Williamson, who has retired but will maintain the position of chairman emeritus. Stephen Williamson has served on the board of directors since 1977, and most recently held the position of vice chairman. He joined T.D. Williamson in 1971 and has served in numerous functional and leadership roles during his career. In January 1989, Mr. Williamson purchased and ran a Canada-based division of T.D. Williamson until January 2007, when it was sold back to T.D. Williamson. He was active in the Canadian Gas Association (CGA) and served on the CGA board for 6 years. Frank Oehler has been named as VP of international sales for JMA Wireless, where he will oversee the sales organizations in the Europe/Middle East, Latin America and Asia-Pacific regions. Prior to joining JMA Wireless, Mr. Oehler held leadership positions at the Kathrein Group and Nokia Siemens Networks. He brings 20 years of experience in business development, account management and marketing to his new role.

Chevron Corp. has elected Dr. Dambisa Moyo (pictured) and Dr. Wanda Austin to its board of directors. Dr. Moyo’s appointment is effective October 11, and she will serve on the company’s audit committee. Dr. Austin’s appointment is effective December 1, and she will serve on the company’s board nominating and governance committee and public policy committee. Dr. Moyo has been the founder and CEO of Mildstorm LLC since 2015. From 2001 to 2008, she worked at Goldman Sachs in various roles. Prior to that, she worked at the World Bank in Washington, D.C. and served as a director of Lundin Petroleum AB from 2009 to 2012. She now serves on the boards of Barclays Plc., Barrick Gold Corp., SABMiller Plc. and Seagate Technology Plc. Dr. Austin joined The Aerospace Corp. in 1979, and served in numerous leadership positions before becoming president and CEO in 2008. She holds an adjunct research professor appointment at USC’s Viterbi School of Engineering. Prior to The Aerospace Corp., Dr. Austin worked at Rockwell Intl. on the technical staff. She serves on the boards of the Horatio Alger Assoc., the National Geographic Society and the University of Southern California.

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Technology and Business Information for the Global Gas Processing Industry

GasProcessingNews.com | SEPTEMBER/OCTOBER 2016

LNG TECHNOLOGY Reduce LNG costs, complexity with optimized expander system Solutions for nearshore FLNG

GAS COMPRESSION

Compact BOG recondenser minimizes equipment costs

PIPELINES

Optimize integrated gas production and distribution

Special Supplement to

CONTENTS

EDITORIAL COMMENT As of early 2016, approximately 140 metric MMtpy of LNG capacity were under construction. Of this volume, 62 metric MMtpy are located in the US and 50 metric MMtpy are in Australia. A total of 42 metric MMtpy of LNG capacity are slated for commercial startup in 2016. Nearly half of the remaining 98 metric MMtpy of LNG capacity under conADRIENNE BLUME, struction are tentatively scheduled to Editor come online by the end of 2019. Approximately 448 metric MMtpy of liquefaction capacity could be in operation by 2020, representing a 40% boost in worldwide LNG output over a five-year period. The large volume of LNG capacity starting up will create an oversupply that could last until 2024. This overhang could also weigh on final investment decisions for LNG projects through the end of 2017. Despite the potential for an oversupply of LNG, liquefaction terminal and FLNG vessel projects will continue to be approved and constructed. The need for LNG imports will grow in South America, Asia and Europe, and the booms in natural gas production in North America and the Pacific will continue to drive export opportunities. Successful LNG and FLNG projects will take advantage of optimized engineering and design to lower the complexity and costs of LNG production facilities, as discussed in this issue's Special Focus on LNG technology. For more LNG/ FLNG spending, construction, trade and technology analyses and forecasts, see Hydrocarbon Processing’s HPI Market Data 2017 annual report, available in October. GP

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9 SPECIAL FOCUS: LNG TECHNOLOGY 13

Dual-methane expander liquefaction reduces LNG costs and complexity G. W. Howe, G. F. Skinner and A. D. Maunder

17

Develop successful nearshore FLNG solutions— Part 2: Natural gas liquefaction S. Mokhatab, S. Basi and P. Hunter

23

Use integrated analyses in design and operation of LNG systems J. Valappil, R. Kumar and J. Mumm

29

Impacts of benzene and piperazine concentrations on LNG plant capacity K. K. Hwang and S. Kim

PIPELINES AND INFRASTRUCTURE 35

Optimize an integrated natural gas production and distribution network D. Aluma, N. Thijssen, K. M. Nauta, C. C. Pantelides and N. Shah

GAS COMPRESSION Super-compact BOG recondensing system minimizes equipment lifecycle costs K. Hayashi, K. Yarimizu and S. Furutani

Adrienne Blume Bob Andrew Lee Nichols

41

Sheryl Stone Angela Bathe Dietrich Ashley Smith David Weeks Amanda McLendon-Bass Cheryl Willis

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“No novelty is the new novelty” for NextDecade’s first LNG export project

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New in Gas Processing Technology Cover Image: The Świnoujście LNG import terminal in Świnoujście, Poland is operated by Polskie LNG, a subsidiary of Gaz-System. The facility, shown here during construction in October 2015, is now in operation. Świnoujście LNG received its first commercial LNG cargo on June 17, 2016 from Qatar.

GAS PROCESSING NEWS

BOB ANDREW, Technical Editor

Cryogenic valves for new vessels

Parker Bestobell Marine received a major new order from Daewoo Shipbuilding and Marine Engineering (DSME) in South Korea to supply cryogenic valves for three new vessels. The company’s valves will be used in the cargohandling systems, including the main discharge line that controls the initial flow of LNG from the cargo tanks when pumping starts. The three vessels are part of the Yamal series of Arc7 ice-class LNG carriers that will be operated in Arctic winter conditions. Parker Bestobell Marine previously supplied cryogenic globe and check valves to the first vessel in the series, owned by Russia’s Sovcomflot. The “on-deck” valves will be subjected to extreme sub-zero temperatures, which is not an issue with cryogenic valves designed to operate at a temperature as low as –196°C. For added protection, Parker Bestobell Marine will supply covers for the headworks to protect exposed valve parts. Due to the arctic conditions in which the LNG carriers will operate, it is not possible to use actuators, which operate via hydraulic oil due to viscosity issues. As a result, Parker Bestobell Marine innovated within its valve design to ensure that the valves could operate efficiently using electric actuators. The LNG carriers will be the first to have electric actuators fitted to globe valves. Fifteen ships are planned to be built in this series by DSME for three different owners: Mitsui O.S.K Lines (Japan), Teekay (Canada) and Dynagas (Greece). The ships are designed specifically for the LNG Yamal project in Russian Siberia. The 15 Arc7 ice-class gas carriers will operate in Arctic conditions, with temperatures as low as –54°C. They will be required to independently navigate ice more than 2 m thick.

GTL conversion without CO2 emissions A team of scientists from CoorsTek Membrane Sciences, the University of Oslo in Norway, and the Instituto de Tecnología Química in Spain has developed a new process to use natural gas as a raw material for aromatic chemicals. The process uses a ceramic membrane to make the direct, non-oxidative conversion of gas to liquids possible for the first time. This process reduces cost, eliminates multiple process steps and avoids carbon dioxide (CO2) emissions. The resulting aromatic precursors are source chemicals for jet fuel, insulation materials, plastics and textiles, among other products. Direct activation of methane has been a key goal of the hydrocarbon research community for decades. By using a ceramic membrane that simultaneously removes hydrogen and injects oxygen, the team was able to produce liquid hydrocarbons directly from methane in a one-step process. Temperature and pressure have historically been the main parameters chemists and engineers have used to control reactions. Catalysts can improve speed and selectivity without promoting reactions beyond their chemical equilibrium limit. Integrating a ceramic ion-conducting membrane into the reactor enables an increase in productivity of industrially appealing processes that are otherwise impractical due to strong thermodynamic constraints. The ceramic membranes are made from abundant materials like barium and zirconium, found within large sand deposits, with the addition of thin electrocatalytic layers of plentiful metals like nickel and copper. While the reactor costs will be standard, the results enabled by this new process have the potential to improve both the financial and environmental costs of chemical production.

Compressor repair savings with unique chemistry

A midstream gas plant in Texas was experiencing severe black powder fouling resulting in compressor, valve and piston failures on two compressors each week. The plant looked to U.S. Water’s energy team for a system solution. After a plant survey, the team discovered that the formation of black powder was the cause of unscheduled compressor downtime. The first recommendation was to conduct a trial antifoulant chemistry on the second-stage compressor inlet to remove the blockage and measure the debris concentration located in the gas streams. The results from the debris analysis allowed U.S. Water to formulate a custom inhibitor package for the two separate trains. Working together with the engineering and equipment team, chemical feed equipment was designed to allow the plant to deliver the antifoulant by atomization throughout the entire system. Since the plant introduced U.S. Water’s solution into its system, one compressor has been running uninterrupted, and the other has been on a six-month run. In addition to continuous production, the plant was able to save $670,800.

4 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

Automation and safety systems for Goldboro LNG Honeywell Process Solutions will provide full automation and safety systems and serve as the integrated main automation contractor (I-MAC) for a new LNG facility being built in Eastern Canada to process North American natural gas for international export. Located on the eastern shore of Canada in Nova Scotia, Pieridae Energy’s Goldboro LNG project will include a natural gas liquefaction terminal and facilities for LNG storage and marine export. The facility will be able to process 10 metric MMtpy of LNG and have a storage capacity of 690 Mm3. Startup is expected in 2021. Honeywell is responsible for designing, delivering and installing the distributed control systems, safety instrumented systems, fire and gas systems and operator training simulator for the project. The company will also be responsible for helping Pieridae integrate all plant infrastructure to the business enterprise systems. Specific key deliverables include a number of Honeywell’s patented technologies. By leveraging these integrated solutions, Honeywell will reduce risks and minimize potential schedule delays for both Goldboro LNG and its engineering, procurement and construction contractor during the facility startup.

Regasification system for ship conversion Wärtsilä has been contracted to supply the regasification system for an FSRU conversion project that Höegh LNG plans to carry out on a modern LNG vessel. The Wärtsilä regasification system to be supplied will feature regasification technology using water glycol as the intermediate medium, instead of propane. This provides a more compact solution, as it is approximately 15% smaller and lighter than the propane-based system. The Wärtsilä scope of supply for this project comprises the water glycol regasification module, the water glycol/seawater heaters and the pumps. Delivery is scheduled for autumn 2017. Wärtsilä previously supplied eight regasification systems to Höegh LNG.

Additional news items can be found online at GasProcessingNews.com.

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US INDUSTRY METRICS

A. BLUME, Managing Editor

US natural gas spot prices at Henry Hub and NGL spot prices at Mont Belvieu, $/MMBtu 25 Natural gasoline Isobutane Butane NGPL composite Propane Ethane Natural gas spot prices (Henry Hub)

$/MMBtu

20 15 10 5 0

Sept. Oct. Nov. Dec. Jan. Feb. Mar. April May June July Aug. Sept. 2016 2016 2015 2015 2016 2016 2016 2016 2016 2016 2016 2016 2016 Source: US EIA

6 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

US gas production (Bcfd) and prices ($/Mcf) 100

7

Production, Bcfd

5

60 40 20 0

4 3 Monthly price (Henry Hub) 12-month price avg. Production J A S O N D J F M A M J J A S O N D J F M A M J J 2014 2015 2016

2 1 0

Gas prices, $/Mcf

6

80

Production equals US marketed production, wet gas. Source: EIA.

US natural gas plant field production of NGL, LPG, ethane/ethylene and propane/propylene, Mbpd 120 US gas plant field production, Mbpd

In the US, Henry Hub natural gas spot prices increased sharply in June and July from suppressed levels earlier in the 2Q, as domestic production decreased slightly. However, natural gas storage inventories were at an all-time high of 2,480 Bcf at the beginning of the injection season on April 1. Inventories were still at record levels as of early September, despite a summer heat wave over a large portion of the US that led to higher power burn. Meanwhile, output of NGL slid in July from record-high levels in May, as gas processing plants marginally curbed production. GP

100 80 NGL LPG Ethane/ethylene Propane/propylene

60 40 20

June- July- Aug.- Sept.- Oct.- Nov.- Dec.- Jan.- Feb.- Mar.- April- May- June2015 2015 2015 2015 2015 2015 2015 2016 2016 2016 2016 2016 2016 Source: US EIA

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EXECUTIVE Q&A VIEWPOINT

“No novelty is the new novelty” for NextDecade’s first LNG export project

SHAUN DAVISON, Senior Vice President of Development and Regulatory Affairs, NextDecade, The Woodlands, Texas

SHAUN DAVISON leads commercial project development for North America, specifically the Gulf of Mexico, for NextDecade LLC. An entrepreneurial professional with more than 20 years in the energy industry, Mr. Davison is a business development and infrastructure specialist with broad-based expertise in all aspects of the natural gas and LNG industry. He has extensive knowledge of the energy/natural gas value chain: pipelines, storage, US and global markets, NGL, and LNG terminals and shipping.

Texas-based NextDecade LLC is working to develop and manage new land-based and floating LNG projects in the US and abroad. At present, the company is focusing on a land-based LNG plant, Rio Grande LNG (FIG. 1), to be located in Brownsville, Texas. In addition to this project, NextDecade continues to explore and develop additional opportunities, including but not limited to the creation of natural gas infrastructure, transportation and storage of LNG, and natural gas/LNG trading. The company is creating new business models for LNG that compete on a fundamentally lower cost basis, providing market players and international end-users access to Henry Hub-indexed or oil-indexed LNG. Gas Processing spoke with Shaun Davison, senior vice president of development and regulatory affairs for NextDecade, about the Rio Grande project’s status and NextDecade’s future development plans. GP. Can you provide a brief background on the Rio Grande LNG project’s conception and development? What is the status of Rio Grande LNG and the Rio Bravo pipeline, and what export opportunities exist for the project?

Davison. Coming off involvement in FLNG developments in Australia and downstream development of FSRUs globally, members of the NextDecade team recognized the potential for smaller and more nimble entrepreneurial companies in the LNG value chain. Experiencing firsthand how FSRU technology has changed and reshaped the traditional LNG industry, our team set its sights on opportunities to challenge the status quo on the upstream side through liquefaction development. After initially reviewing an FLNG solution for a US-based project, we recognized the likely regulatory hurdles that were legacy issues from the original

FSRU/offshore regasification projects. Working with our industry partners, we developed an alternative path optimizing a land-based solution that capitalizes on the US’ abundant gas supply, vast pipeline network and transparent regulatory process for traditionally engineered landbased projects. For NextDecade’s first US export project, our technology mantra has been “no novelty is the new novelty.” Rio Grande LNG is our proposed LNG export facility at the Port of Brownsville in Texas. The project, planned for a 1,000acre industrial site, includes the 140-mi Rio Bravo Pipeline. The project is now going through the rigorous Federal Energy Regulatory Commission (FERC) permitting process. On May 5, 2016, we submitted our full National Environmental Policy Act (NEPA) application (assigned docket numbers CP16-454 and CP16-455, respectively). We expect to receive the Draft Environmental Impact Statement (DEIS) before the end of 2016. A final investment decision (FID) is expected to take place sometime in mid-2017. In November 2015, we announced that NextDecade had signed non-binding agreements for 14 MMtpy of LNG with customers from across Asia and Europe. Since then, that number has grown to 30 MMtpy, demonstrating a continued desire for US-produced LNG from customers around the world—and, specifically, a desire for LNG from the Rio Grande facility. GP. What benefits has NextDecade experienced in working with domestic LNG solutions providers?

Davison. Our team has great experience across the full LNG value chain; however, we recognize and appreciate the tremendous value of partnering with the best energy and LNG companies out there to deliver safe, sustainable, economically sound and environmentally responsible projects. We are proud to have closeGas Processing | SEPTEMBER/OCTOBER 2016 9

EXECUTIVE Q&A VIEWPOINT uefaction plant development, which does not necessarily make it easy for FLNG to break into the existing value chain. GP. What will be the impact of low commodity prices on the global LNG market? How will it affect US LNG export projects?

FIG. 1. Aerial view of the proposed Rio Grande LNG facility at the Port of Brownsville in Texas.

ly partnered and built strong relationships within all of the key support industries necessary to develop, permit, construct and operate LNG projects. For the Rio Grande LNG and Rio Bravo Pipeline projects, these include Ecology & Environment, CH-IV, Norton Rose Fulbright, CB&I, and Moffat & Nichol, among others. The wealth of expertise and insight that these companies bring to our projects is evidenced by the quality of the products they deliver, and visible throughout our FERC and regulatory filings. GP. What opportunities does NextDecade see for FLNG?

Davison. NextDecade’s core team has been very involved with the creation and evolution of the floating regasification industry. NextDecade’s CEO, Kathleen Eisbrenner, was the founder and CEO of the original FSRU leader, Excelerate Energy, spearheading the development of this segment of the valve chain and dramatically altering the LNG industry. We have seen some minor growth in the upstream/liquefaction component of the FLNG industry per the development of Shell’s Prelude project offshore Australia, at least one Malaysian project from Petronas, and the construction of Exmar’s small-scale floating liquefaction module previously aimed at Colombia’s Pacific Rubiales-led project. However, at NextDecade, we are also looking for other opportunities. We do believe there will be some opportunities in the FLNG space, but these are likely to be explored by national oil companies, or international oil companies like Petronas and Shell, rather than by smaller players. Despite the smaller scale of FLNG, there remains a high-dollar barrier to entry, and there are limited market areas that would allow for smaller developers like a NextDecade to participate. At the same time, the US is expected to remain the incremental low-cost opportunity for liq-

Davison. Buyers are pushing back on long-term contracts and pricing agreements that were executed when commodities were high, while suppliers are making concessions due to low commodity pricing and current oversupply. Liquefaction projects that have been proposed or are under development have, in many cases, delayed or deferred their FIDs. In some cases, they have been canceled altogether. At the same time, on the downstream side, there is an aggressive push to develop regasification capacity to take advantage of this low pricing environment. As we move through the present phase of the market cycle, with liquefaction projects curtailing and imports expanding, the industry will inevitably need to rebalance. For US projects that continue to advance with an eye toward the 2021–2023 market—at which point it is quite possible that there will be a global undersupply—discussions with customers are in a state of suspension. Few customers are ready to execute sales-and-purchase agreements (SPAs). However, at the same time, they are pushing hard for pricing at preCheniere-Train-3 (i.e., sub-$3/MMBtu) levels, with an eye to 2017–2018 as they begin to plan for the 2020s. GP. Do you foresee a glut of LNG supply impacting the success of Rio Grande LNG or other US LNG export terminals scheduled to come online over the next few years? Should US LNG producers be investigating other export outlets?

Davison. At present, the conventional wisdom is that, by 2021–2023, the existing LNG supply overhang will have been absorbed, and the market will be looking for the next tranche of LNG supply. In that regard, we do not see the existing supply overhang impacting the success of Rio Grande LNG; however, that does not mean that all current US export projects will be standing when the market comes back. Only well-developed, well-engineered and economically sound projects will remain.

10 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

The downstream LNG industry has been changing quite rapidly with the development of the FSRU. Floating regasification technology has opened up new markets previously shut out of the LNG space. Countries like Argentina, Indonesia, Brazil, Kuwait, Lithuania, Pakistan and Jordan have become LNG importers due to economical FSRU technology. The potential for FSRUs to provide broader LNG market penetration opens up entirely new markets, and US LNG producers must be willing to work with these customers rather than relying on the traditional players from Europe, South Korea and Japan, for example. GP. How does NextDecade foresee LNG contracts being renegotiated in the present volatile market?

Davison. It is evident that the traditional Asian players, especially Japan, are seeking more destination flexibility with their historic long-term supply. As far as new US projects are concerned, the present oil price climate has put downward pressure on the US formula, and most customers are pushing for a sub-$3/ MMBtu liquefaction fee. However, with no new SPAs executed, there is still somewhat of an unknown component to US pricing for the second wave of projects. GP. What actions or activities might help energy prices to recover to $80/bbl for oil and $4/Mcf for gas?

Davison. I, along with many others, wish we had the silver bullet for that question! My response assumes the recognition that oil (Brent) and gas (Henry Hub) are not linked and have very different drivers. Brent is largely affected by international markets, and Henry Hub is driven by US/North American supply and demand factors. For oil, greater alignment between OPEC and others could certainly have the greatest and quickest impact to oil prices, but it’s anyone’s guess as to the likelihood of that occurring anytime soon. For US natural gas prices, there is always the specter of US policy issues tied to new regulations for exploration, production and/or power generation that can potentially impact prices, making gas more costly. Additionally, the market demand response could, to some degree, lift US gas prices, with LNG exports being the greatest driver of this demand. GP

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SPECIAL FOCUS: LNG TECHNOLOGY

Dual-methane expander liquefaction reduces LNG costs and complexity G. W. HOWE, G. F. SKINNER and A. D. MAUNDER, Gasconsult Ltd., London, UK

LNG producers have sought to enhance project returns through higher plant capacities to achieve economies of scale. Many complex multi-refrigerant plants were constructed to realize this objective. However, the unprecedented capital costs and financial risks associated with these mega-scale plants may be unsustainable in an era of distressed energy prices. Some operators are looking for more flexible project development and commercial models that alleviate risk; they are seeking means to monetize smaller gas reserves with lower-cost schemes. Dual-methane (DM) expandera liquefaction offers a differentiated solution for mid-scale and floating LNG (FLNG) applications. Unlike conventional processes, it uses no external refrigerants, utilizing instead natural gas feed as the refrigerant medium in an optimized system of expanders. This setup eliminates refrigerant storage and transfer systems used in MR cycles, as well as the additional process equipment used to extract refrigerant components from the feed gas. Makeup refrigerant is low-cost natural gas, as opposed to nitrogen or a mixture of hydrocarbons, thereby avoiding complex supply logistics and reducing operating costs. The absence of liquid hydrocarbon refrigerant also makes for safer operations. The methane process requires significantly less power than other “safe” systems, such as multiple-expander nitrogen processes, allowing reduced capital cost through lower installed compressor power or increased LNG production from a selected compressor driver. Single-train capacities exceeding 2 MMtpy are possible, facilitating phased project development and lower initial capital requirements, yet still allowing a progressive buildout of significant LNG capacity. The avoidance of equipment to produce, handle, store and process external refrigerants reduces cost, weight and footprint, making the technology particularly attractive for FLNG schemes. Where appropriate, freed-up deck space could be used to install additional productive liquefaction capacity, which would enhance project returns. A number of variants of the basic configuration have also been developed for low-pressure (LP) feeds, the removal of heavy hydrocarbons and to facilitate benefits from high-speed (HS) compression. These variants further differentiate the technology and are described in this work. Process configuration. A simplified schematic of a proprietary DM expander process is shown in FIG 1. Refrigeration is effected in two expander circuits, a warm circuit indicated in red and a low-temperature circuit shown in blue. Chilled gases

from expanders CX1 and CX2 are routed to the cold box for cooling duty, and then returned to the expanders by the recycle compressor CP1. Flash gas is also routed through the cold box for cooling duty and recaptured to the system by a small compressor, CP2, which feeds the suction of the recycle compressor. The expanders are configured as companders and operate in series with the recycle gas compressor (FIG. 2), providing approximately 35% of the total compression power. The methane cycle is similar in concept to nitrogen expander schemes. However, it enjoys a fundamental advantage, as methane has a higher specific heat than nitrogen. This factor significantly reduces circulating gas flows, which, in turn, reduces power consumption and pipe sizes. A patented feature of the described process is that partial liquefaction takes place in the low-temperature expander CX2—this efficiently converts latent heat directly into mechanical work and also permits a reduction in heat-transfer area and cost of the main heat exchanger HX1. An optional liquid turbine, TU1, in the LNG rundown line also improves efficiency by providing a significant chilling effect. These features, together with the optimized distribution of flows, temperatures and pressures in the expander circuits, makes for a highly energy-efficient system consuming approximately 300 kWh/metric t of LNG in temperate climates. This performance is equivalent or better than that involving singlemixed-refrigerant (SMR) processes, and 15%–30% lower than the more sophisticated variants of dual- and triple-expander nitrogen schemes. To fuel CP1

CP2

Recycle gas

Zone 1

Cold box HX1 Zone 2

Flash gas

Zone 3

TU1

Pretreated feed gas CP: Compressor SP: Separator CX: Compander TU: Liquid turbine

SP1

SP2

SP3

NGL CX1

CX2

LNG

FIG. 1. Simplified schematic of proprietary DM expander process. Gas Processing | SEPTEMBER/OCTOBER 2016 13

SPECIAL FOCUS: LNG TECHNOLOGY Alternative configurations. A number of variants of the

technology have been investigated. These variants have the potential to further reduce capital cost and/or increase operating efficiency. Open methane cycles lend themselves to advantageous configurations for LP feed gases, removal of heavy hydrocarbons and HS rotating equipment.

IPL for LP feed gases. All liquefaction technologies consume more power at lower feed gas pressures. The integrated pressure liquefaction (IPL) variant (FIG. 3) of the described process boosts LP feed gas by routing it after liquids separation in SP1 back to an interstage suction point on the recycle gas compressor, instead of to Zones 2 and 3 of the liquefaction section of the cold box, as shown in FIG. 1. This process provides a higher inlet pressure to the cold box independent of the feed gas pressure, TABLE 1. Basis of design for the DM process Gas composition, mol%

CH4 of 95; C2H6 of 4; C3H8 of 1

Gas pressure at liquefaction inlet

As indicated

Feed gas pressure

As indicated

Process streams cooled to, °C

–40 and 40

Heat leak to cold box

Integrated heavies removal (IHR). With heavier feed gases

0.50%

Minimum cryogenic approach temperature, °C

above or close to their critical pressures, adequate removal of C5+ and aromatics may require an upstream NGL unit. Typically, this expands the feed gas to a sub-critical pressure, condenses the heavy material and then recompresses the depleted gas for liquefaction.

3

Recycle gas compressor polytropic, η

85%

Expander adiabatic, η

87%

enhancing liquefaction efficiency without the need for a separate feed gas compression plant. For 40°C ambient conditions on 25-bar pipeline gas (as might prevail, for instance, in the US Gulf Coast region), IPL operation at 80 bar achieves a reduction in power demand exceeding 20% of that for the basic DM expander system (FIG. 4). This capability is only available to open methane cycles. Nitrogen or SMR schemes require an additional compression facility to enhance liquefaction cycle efficiency; unlike open methane cycles, they do not have a methane compressor in their basic configuration. FIG. 5 provides the authors’ computations of the relative power demand measured in kWh/metric t of the DM (in IPL mode), SMR and dual-nitrogen processes over a pressure range of 20 bar–80 bar. This data is based on normalized machine efficiencies and provides an indication of the relative merits of the technologies in a warm climate and feed gas precooled scenarios (40°C and –40°C “cooled to” temperatures, respectively). The DM process is advantaged over the full data range. The FIG. 5 data assumes the design basis provided in TABLE 1.

30

Recycle gas

To process Reduction in power

LT compander

Compressor driver

25

HT compander

15 10 5

Recycle compressor

0 20

FIG. 2. The expanders operate in series with the recycle gas compressor, providing 35% of total compression power.

30

40 Feed gas pressure bar

50

60

FIG. 4. IPL operation at 80 bar achieves a reduction in power demand exceeding 20% of that for the basic DM expander system.

To fuel CP1

20

CP2

700 Cold box HX1 Zone 1

Zone 2

500

Zone 3

TU1

Pretreated feed gas SP1

SP2

SP3

kWh/metric t

Recycle gas

Dual N2 -40°C SMR -40°C Dual CH4 -40°C

600

Flash gas

40°C

400 300

Dual N2 40°C SMR 40°C Dual CH4 40°C

-40°C

200 100

NGL CX1

CX2

LNG

FIG. 3. The IPL variant boosts LP feed gas by routing it after liquids separation back to an interstage suction point on the recycle gas compressor.

14 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

0 20

30

40

50 Feed gas pressure bar

60

70

FIG. 5. Computations of the relative power demand of the dualmethane (in IPL mode), SMR and dual-nitrogen processes at a pressure range of 20 bar–80 bar.

80

SPECIAL FOCUS: LNG TECHNOLOGY

TABLE 2. Economically matched equipment configurations LNG production, metric MMtpy

0.9

1.1

1.5

2.2

PGT25 + G4

LM6000PF

LM6000 + MD

Frame 7

2BCL800/29.8

2BCL1007/34.8

2BCL1400/46.7

2BCL1400/81.5

LT expander/power MW

EC50-1/5.4

EC50-1/6.7

EC50-1/9.5

EG50-1/13.4

HT expander/power MW

EC60-1/11.1

EC50-1/13.8

EC50-1/8.6

EC60-1/13.8

EC50-1/8.6

EC60-1/13.8

Gas turbine Compressor/absorbed MW

HT expander/power MW

In its IHR variant, the described DM cycle process removes heavy components by passing the feed gas and recycle gas through the warm circuit gas expander CX1 (FIG. 6), and separates the condensed heavy material from the expander outlet at subcritical pressure, around 10 bar–15 bar. This solution decouples the vapor/liquid separation and feed gas pressures and saves a large part of the equipment and cost of a separate expander-based NGL removal unit. Specifically, the compander (CX1), recompression facilities (CP1) and associated bulk materials already exist in the basic DM configuration, avoiding additional capital cost. Weight and footprint are also reduced, which is particularly relevant to FLNG schemes. High-speed compression. Methane can be compressed at

a significantly higher rotational speed than the higher-molecular-weight hydrocarbons found in mixed refrigerant cycles. This permits the use of HS driver/compressor combinations that are significantly lower in cost and weight than those used in conventional processes. In a recent study conducted around an HS 46-MW output gas turbine (FIG. 7), the main recycle compressor is direct-driven and runs at a much higher speed (6,600 rpm) than conventional MR compressors. Since the turbine is a single-shaft machine, a small starter/helper electric motor, M1, is provided. The study, developed with support from the OEM, demonstrated an ability to achieve a capacity of 1.5 MMtpy/train with a power demand of approximately 310 kWh/t of LNG, based on a feed gas pressure of 60 bar, ambient air temperature of 30°C and seawater temperature of 23°C. Significant weight savings (> 70 metric t) for the gas turbine and the recycle compressor were demonstrated compared to a typical aeroderivative based solution, with cost savings for this equipment measured at 20%–30%. Although the gas turbine subject of the study was an industrial machine, the changeout capability at 48 hr is comparable to aero-derivative machines. The turbine also has longer equivalent operating hours (60,000 hr) between major overhauls than aero-derivatives. The energy efficiency at > 38% was respectable, with DLE < 15 ppm NOx . Technology advantages. In addition to its low power de-

mand, reduced equipment count and low footprint, a further set of advantages accrue to methane cycles from the absence of external refrigerants. Many of these advantages have particular relevance for FLNG schemes, where weight, deck space, operational simplicity and safety are important factors: • No refrigerant logistics issues are present in remote or offshore locations. Neither shipments of light and

To fuel CP2

CP1

Cold box HX1

Flash gas

Recycle gas

TU1

Pretreated feed gas SP2

SP1

SP3

inc C5+ aromatics CX1

LNG

CX2

FIG. 6. In its IHR variant, the DM cycle process removes heavy components by passing the feed gas and recycle gas through the warm circuit gas expander. Fuel gas

Natural gas

High-speed GT

CP1

M1

M2 Cold box HX1 Zone 2

Zone 1

Zone 3

TU1

CP2

Flash gas

Pretreated feed gas Recycle gas CX1 CX2

SP1 NGL

CP: Compressor SP: Separator CX: Compander TU: Liquid turbine

SP2

SP3

CX1 CX2

LNG

FIG. 7. In an HS 46-MW output gas turbine, the main recycle compressor is direct-driven and runs at a higher speed than conventional MR compressors.

• • • •

heavy hydrocarbons, nor segregated storage to facilitate blending a mixed refrigerant, are required. Absolute security of refrigerant supply is ensured. No propane or other liquid hydrocarbon refrigerants are present, which offers a major safety advantage relative to MR schemes. Single-phase refrigerant (always a gas) makes the system motion-tolerant. Operational benefits relative to MR schemes are present. These benefits include no refrigerant makeup cost, no refrigerant composition adjustments to maintain Gas Processing | SEPTEMBER/OCTOBER 2016 15

SPECIAL FOCUS: LNG TECHNOLOGY cycle efficiency, shorter startup time from warm condition and reduced flaring. Project returns. Most liquefaction schemes are built around

a preselected compressor driver. An economically matched set of ancillary process equipment is assembled around this driver. Once the compressor driver is selected, the power available for liquefaction is set. Then, the overwhelmingly dominant factor determining LNG production is the liquefaction cycle efficiency. For mid-scale projects, the differential can run to some hundreds of millions of dollars, as measured by NPV. FIG. 8, developed by the authors from a case study, plots cumulative NPV vs. time for the DM cycle, a dual-nitrogen cycle Dual CH4 SMR Dual N2

IRR 44%

IRR 51%

IRR 39%

NPV, $MM

4,000 3,500 3,000 2,500 2,000 1,500 1,000 500 0 -500 -1,000 -1,500

0

1

2

3

4

5

6

7 8 Years

9

10

11

12

13

14

15

FIG. 8. Cumulative NPV vs. time for the DM cycle, a dual-nitrogen cycle and a basic SMR scheme for a 4-metric-MMtpy, five-train FLNG project.

and a basic SMR scheme for a nominal, 4-metric-MMtpy, fivetrain FLNG project monetizing a 2-Tcf gas field. The DM cycle earns higher returns over a shorter period of time because its superior efficiency supports a higher production capacity. Technical validation. All equipment in the methane cycle process is fully proven in operation, and the process steps are well established in dozens of cryogenic gas processing plants. BP and three engineering companies (under nondisclosure agreements) have reviewed the described design, from simulations through key process and detailed design parameters. All of the companies confirmed the energy efficiency and key performance parameters. Leading equipment vendors have confirmed the mechanical design/configuration viability and that all equipment operates within a window of proven operating experience. Work performed in conjunction with a leading OEM established economically matched rotating equipment configurations around various gas turbine drivers for single-train capacities in the range of 0.9 metric MMtpy–2.2 metric MMtpy. An important outcome from this work was confirming an achievable train capacity of > 2 metric MMtpy. For the particular compressor driver, this capacity is significantly higher than that achievable by nitrogen or the simpler SMR processes, providing an economy-of-scale advantage for the methane cycle. Takeaway. For mid-scale operation, the DM expander process combines high energy efficiency with a fundamental simplicity, low equipment count and low investment cost. Elimination of external refrigerants provides OPEX, CAPEX and logistics advantages and simplifies operations. The open-methane cycle allows advantageous variants to the basic process. Single-train capacities exceeding 2 metric MMtpy of LNG allow substantial production capacity buildout on a phased basis, decreasing upfront costs and reducing project risk. No equipment supply is tied to the technology licence, and all equipment is available from multiple vendors, allowing for fully competitive procurement with cost and schedule benefits. GP a

NOTE The dual-methane expander process described in this article is the ZR-LNG (Zero Refrigerant LNG) system that is owned, patented and licensed by Gasconsult Ltd. The gas turbine cited in the high-speed compression variant is a Siemens SGT800. The various economically matched equipment configurations in the capacity range 0.9 metric MMtpy–2.2 metric MMtpy were worked up with GE Oil & Gas and are as shown in TABLE 2, with a basis of design as per TABLE 1 (but with process streams cooled to 20°C).

BILL HOWE is the CEO of Gasconsult Ltd. He graduated as a chemical engineer from Birmingham University and has spent 30 years in the E&C industry, mainly with Foster Wheeler. He was managing director of Foster Wheeler’s South African affiliate and, subsequently, director of sales and marketing at Foster Wheeler Reading in the UK. GEOFF SKINNER is a director at Gasconsult Ltd. He graduated from Oxford University and joined Foster Wheeler in the UK in 1965. From 1981 to 1986, he was technical director of Foster Wheeler Synfuels Corp. in Livingston, New Jersey. Upon his return to the UK, Mr. Skinner acted as a consultant to several multinational companies and has registered a number of patents, including LNG liquefaction processes. TONY MAUNDER is a director at Gasconsult Ltd. He holds degrees in mechanical sciences and chemical engineering from Cambridge University. After working with ICI General Chemicals, he spent 16 years in the E&C industry, including with Foster Wheeler. From 1980 to 1993, he worked for BP Research and BP Engineering on natural gas conversion to liquids, synthesis gas and fuels. He has registered a number of LNG liquefaction process patents.

16 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

SPECIAL FOCUS: LNG TECHNOLOGY

Develop successful nearshore FLNG solutions— Part 2: Natural gas liquefaction S. MOKHATAB, Gas Processing Consultant, Dartmouth, Nova Scotia, Canada; and S. BASI and P. HUNTER, KBR, London, UK

Nearshore small- to mid-scale (0.5 MMtpy–2.5 MMtpy) FLNG production technology, which is viable for monetizing nearshore gas fields or converting pipeline-quality gas into LNG, provides another push to the FLNG market. As such, there is great interest in developing successful solutions for nearshore FLNG projects that may have advantages over competing onshore projects. Part 2 of this series discusses various natural gas liquefaction processes to allow a selection that best fits the needs of FLNG applications. In addition, this paper addresses some of the design considerations that can reduce capital costs, thereby improving the overall economics of nearshore FLNG projects.

FLNG DESIGN BACKGROUND It is clear that the design criteria for an FLNG facility is quite different from a land-based LNG plant in terms of process safety, storage, compactness, design flexibility, sensitivity to motion and simplicity of operation. Sensitivity to motion, which can cause process upsets or mechanical failures for an open-ocean offshore FLNG facility, is a lesser concern for a nearshore FLNG located in a relatively calm marine environment. Further protection for a nearshore FLNG vessel can be provided by breakwaters or by fixing the FLNG vessel to a gravitybase structure or jetty structure. When the FLNG vessel is located adjacent to the shoreline, referred to as at-shore FLNG, opportunities exist to decongest the topside design by locating some non-hazardous elements onshore. A successful nearshore barge-mounted FLNG project requires a proven liquefaction technology that is safe, simple, easy to operate and maintain, flexible for variations in gas composition and turndown, scalable and efficient in terms of space and cost. A number of design factors must also be addressed, such as machinery configuration, heat exchanger type, and utility systems. Each FLNG facility must be tailored to sitespecific conditions to determine the optimal nearshore FLNG plant configuration. FLOATING LIQUEFACTION CYCLES Three main types of refrigeration cycles (cascade, mixed refrigerant and turboexpander) have been proposed for FLNG applications. Based on a proven track record with onshore facilities, mixed refrigerant (MR) and turboexpander-based technologies have been qualified by numerous operating companies for FLNG applications. Although the Conoco Phillips Optimized Cascade Process is proven in mid- to large-scale onshore service, it is being con-

sidered for offshore service only by the technology licensor. As a result, the ConocoPhillips Optimized Cascade process is not covered in detail in this review. Mixed-refrigerant cycles. MR technology has been assessed for offshore liquefaction for both single-MR (SMR) and dualMR (DMR) cycles. The SMR process (FIG. 1) benefits from operational simplicity and flexibility, as well as reduced equipment count; however, these benefits come at the expense of lower efficiency compared to the DMR cycle. With two refrigerant loops, the DMR cycle (FIG. 2) better matches the MR boiling curves to the overall feed gas condensation curve. The DMR process has been successfully applied to large-scale onshore LNG projects, while the SMR process is a proven solution for smaller onshore LNG facilities. Onshore SMR cycles are provided by Air Products and Chemicals Inc. (APCI), Linde and Black & Veatch, while DMR cycles are provided by APCI and Shell. The APCI propane precooled single MR (C3-MR) process is a two-refrigerant loop process in which precooling is performed in multiple kettle-type heat exchangers with a propane refrigeration loop. This feature helps the process achieve a higher efficiency than the SMR process, due to the ability to better match the MR boiling curve to the feed gas condensation curve. However, the large inventory of propane and the relatively large plot space that is required for the propane evaporators make the C3-MR process less attractive for FLNG projects than other cycles. The advantage of SMR and DMR technoloCompressor Treated feed gas Cooler MR loop

LNG FIG. 1. Typical process scheme of an SMR cycle. Gas Processing | SEPTEMBER/OCTOBER 2016 17

SPECIAL FOCUS: LNG TECHNOLOGY gies is the use of compact heat exchangers, which require significantly less topside footprint than the C3-MR technology. Another variant of MR technology is the Mixed Fluid Cascade (MFC) process developed by Linde. This process has a similar efficiency to the DMR technology, but it utilizes three discrete MR refrigeration loops within two or three spiral-wound heat exchangers in series. This configuration increases the equipment count, congestion and hydrocarbon inventory, which is undesirable for small- to mid-scale FLNG facilities. Expander cycles. Turboexpander refrigeration cycles, which

use nitrogen (N2) as the refrigerant, have been widely used for small-scale LNG production (up to 0.8 MMtpy per train). While the single-expander cycle has relatively low efficiency, higher process efficiencies can be achieved by adding a second (FIG. 3) or third expander. Adding a precooling cycle, based on propane or other refrigerants (e.g., hydrofluorocarbons, carbon dioxide or methane), can improve overall efficiency and significantly increase throughput for a given barge size. Increased FLNG facility complexity, reduced overall reliability and the need for increased refrigerant storage are potential disadvantages of adding a precooling refrigeration circuit. Many of the safety concerns can be negated by the selection of non-flammable refrigerants, and reliability concerns can be addressed by the use of parallel equipment or trains on larger FLNG facilities. Several developments utilizing methane (sourced from the feed gas) and/or its combination with N2 have been proposed based on the dual-N2 expander cycle, but none of these schemes have been proven at small- to mid-scale capacities. LNG

Treated feed gas

Precooling MR compressor

Coolers

Cold MR compressor

FIG. 2. Typical process scheme of a DMR cycle.

Process comparison. A comparison of liquefaction cycles is necessary to select an appropriate process for a given FLNG opportunity. TABLE 1 compares the liquefaction processes, taking into consideration typical criteria that influence the technical and commercial acceptance of small- to mid-scale FLNG projects. Of the processes compared in TABLE 1, the Cascade, C3-MR and DMR cycles are well-proven within large onshore LNG facilities. In adapting these technologies to a small- to mid-scale FLNG design, factors such as hydrocarbon inventory, equipment count and equipment selection must be addressed. None of these challenges are insurmountable or preclude the use of these technologies for FLNG; however, for a facility producing up to 2.5 MMtpy, other technologies may be more suitable. The main disadvantage of the MR cycles (relative to the expander processes) is that they have a large hydrocarbon inventory at elevated pressure and require storage for significant quantities of liquid hydrocarbon refrigerants. These features raise additional safety concerns within a confined FLNG facility. For all hydrocarbon-based refrigerant cycles, the sourcing of refrigerants may be an issue when processing lean feed gases (e.g., pipeline gas) that do not contain significant quantities of LPG. In these circumstances, the refrigerants can either be imported or extracted from the feed gas with the aid of a frontend NGL extraction scheme. However, this additional scheme will increase the weight, congestion, hydrocarbon inventory, cost and complexity of the FLNG topside. The dual-N2 expander cycle provides many benefits for offshore and nearshore FLNG applications. A major advantage of using N2 as the cycle fluid is that it is inert, non-flammable and inherently safe. A process using inert refrigerant allows for compact equipment spacing on the FLNG vessel, as long as appropriate safeguards are taken to mitigate the asphyxiation risk in the event of a large leak. Using N2 eliminates the need for a C2/C3/C4 fractionation process for refrigerant makeup or for onboard refrigerant storage, since the N2 production needed for equipment purging and inerting can be augmented. N2 is circulated in the gaseous phase at all points of the refrigeration cycle, so maldistribution in the heat exchangers is not a concern, unlike other refrigeration cycles that use multicomponent refrigerants. As a result, the dual-N2 process performance is less sensitive to marine-induced motion. The dual-N2 design is simple, and the reduced complexity will require less operator intervention than MR-based cycles. The control of specific temperatures is not as important for

TABLE 1. Comparison of liquefaction processes for small- to mid-scale FLNG projects Thermal efficiency

Dual-N2 expander

SMR

C3-MR

DMR

Cascade

Medium

Medium

High

High

High

Equipment count

Medium

Low

Medium

Medium

High

Hydrocarbon refrigerant inventory

None

Medium

Large

Medium

Large

Small-scale FLNG suitability (nearshore/at-shore)

High

High

Low

Low

Low

Mid-scale FLNG suitability (nearshore/at-shore)

Medium

Medium

Low

Medium

Low

Compactness

High

Medium

Low

Medium

Low

18 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

SPECIAL FOCUS: LNG TECHNOLOGY dual-N2 as with MR cycles, and the process is inherently stable and robust. Another important attribute of the dual-N2 cycle is the ability to quickly start up and shut down in a safe and controlled manner. The dual-N2 and SMR cycles are considered more appropriate technologies for FLNG facilities at the smaller end of the capacity range, as the relative inefficiency of each cycle results in a low LNG-production-to-topside-area ratio.

DESIGN CONSIDERATIONS In addition to designing a compact, flexible and energyefficient gas liquefaction process, several design considerations must be properly addressed at the conceptual stage so that fundamental decisions can be made for developing a nearshore FLNG project. The following sections describe some of these considerations. Liquefaction system. Key design requirements of the liquefaction system must be considered for the development of any FLNG project. This section will review main process equipment, train size and utility systems. Main process equipment: Liquefaction heat exchangers. The main cryogenic heat exchanger, a critical piece of equipment in the liquefaction system, will be either a plate-fin heat exchanger (PFHE) or a spiral-wound heat exchanger (SWHE), depending on the liquefaction process selected. PFHEs, often referred to as brazed-aluminum heat exchangers, consist of an aluminum core of alternating layers (passages) of corrugated fins. The layers are separated from each other by parting sheets. Each core will typically comprise no more than five streams. The design of each layer is optimized by varying fin type, fin height and pass geometry to maximize the heat transfer coefficient from streams in close proximity to each other. Layers are stacked within a core up to a maximum manufacturing limit. For liquefaction service, it is possible to manifold up to 6–8 cores together within a cold box. PFHEs are specialized equipment manufactured by several vendors. The main advantages of PFHEs over SWHEs are their compactness, low equipment weight, small footprints, shorter lead times and lower capital costs. One disadvantage of PFHEs is that they are vulnerable to mechanical damage or damage from thermal shocks due to transient operating conditions, particularly in two-phase service. Exchangers subjected to repeated thermal excursions could fail, resulting in refrigerant or hydrocarbon leaks to atmosphere. This risk is partially mitigated by using PFHEs for single-phase, smaller-scale expander liquefaction cycles. Distributing two-phase refrigerant across multiple cores could prove problematic for an FLNG facility, even with mild marine motion. The process must be operated to ensure a temperature difference between the different passes of no more than 50°F (28°C) to minimize thermal stress. Since PFHEs are contained within a self-supporting, insulated cold box with internal distribution headers, accessibility for in-place cleaning and repairing is difficult and time-consuming. SWHEs are essentially vertical, helically wound shell-andtube exchangers designed with a high heat transfer area, which allows them to operate with a larger temperature gradient. The advantages of an SWHE include its proven tolerance to ther-

mal shocks (resulting from transient refrigerant/load imbalances) and containment of tube leaks within the exchanger shell. These features make the use of SWHEs an appropriate selection for MR cycles in FLNG applications. Compared to PFHEs, the SWHE is higher in capital cost, size and weight. The number of SWHE suppliers is limited, with each exchanger having a longer delivery time compared to a single, large cold box. Technical design limits the available heat exchanger size for FLNG service. A single cold box typically can be designed to liquefy up to 0.75 MMtpy of LNG. However, a higher-train capacity would require multiple cold boxes to be grouped together in parallel to provide the necessary heat exchange area. For FLNG, a single SWHE in SMR service can be designed to produce up to 1.5 MMtpy. In precooled MR service, it can be designed for up to 4 MMtpy of LNG. In general, PFHEs can be a good choice for small- to mid-scale floating liquefaction due to their lightweight, compact and highly efficient design for simultaneous heat exchange between multiple streams. For FLNG projects that are designed to utilize either PFHEs or SWHEs, the exchanger must be designed for the transient structural loads and vessel motions in addition to the thermal and hydraulic loads. Both PFHEs and SWHEs are well proven within onshore LNG plants, and PFHEs have been successfully used offshore for many years. However, both types of exchangers are currently unproven in nearshore or offshore LNG service, with the first FLNG facility due to come onstream in 2016. Compressor drivers. The refrigeration compressors in LNG plants are commonly driven by gas turbines. Electric power for LNG facilities is commonly provided by gas turbine generators. In some cases, electric motors or steam turbines have been used as compressor drivers. However, selecting the best driver for an FLNG facility is a challenging issue. Several parameters must be considered when addressing this issue, Compressor Cooler Treated feed gas

Warm expander

Cold expander LNG FIG. 3. Typical process scheme of a dual-expander cycle. Gas Processing | SEPTEMBER/OCTOBER 2016 19

SPECIAL FOCUS: LNG TECHNOLOGY such as thermal efficiency, ease of operation, availability, economics, space and weight limitations, maintenance and safety considerations. Heavy-duty industrial gas turbines, commonly used in onshore LNG plants, are difficult to deploy on an FLNG vessel as they are large in size, require significant maintenance periods and have limited speed variation capabilities. Aeroderivative gas turbines, which are lighter and more compact than heavy-duty frame industrial gas turbines, are better suited for floating installations, although they have been used only in LNG projects at power levels below 45 MW. Aeroderivative turbines have relatively high availability and reliability and are easy to service and maintain. In addition, they are more thermally efficient than industrial gas turbines (TABLE 2), resulting in lower fuel consumption and lower carbon emissions per unit of power. Multiple aeroderivative gas turbines may be installed in parallel to achieve the required compression power demand for large LNG trains. With the many advantages of the aeroderivative turbines and the potential capital cost and energy savings over industrial machines, their use is being accepted for process drivers, particularly in smaller FLNG facilities. Commercially available packages, in which aeroderivative turbines are coupled with centrifugal compressors, provide a new solution for FLNG applications. These packages feature faster installation, enhanced reliability and availability, ease of maintenance, and a reduced footprint and weight relative to frame machines. It should be noted that all gas turbine options have some inherent limitations that require detailed evaluation to determine the optimal selection for each FLNG facility. Electric motor drive systems for the main refrigerant compressor(s) are more expensive than gas turbines, but they are an attractive option due to their higher availability, ease of operation and lower emissions. They may be an option for atshore FLNG applications, as some of the larger electrical elements (e.g., switchgear, transformers) can be located off-barge. Some elements, however, must be located on the barge (e.g., variable-speed drive cabinets, harmonic filters). In the event of importing electric power to a nearshore facility, motor drives will result in a shorter project schedule than for a project that includes a power generation unit. Electric motors of 65 MW have been used for an onshore LNG project, and 78-MW motors have been tested for a project under construction. While large electric motors offer high reliability, they also require a complex electrical system for startup and control. The economic feasibility of using large electric drives

is dependent on the availability of a local, low-cost, reliable, high-voltage power supply. Traditional steam turbine drivers are the best choice for overall reliability, availability and safety (no open ignition source). However, steam turbines necessitate more infrastructure (steam boilers, large heavy pipes, etc.) and include complex operations (requiring freshwater makeup and wastewater/ chemical disposal). In addition, steam systems have a large footprint, more weight and the highest maintenance burden of the driver choices. The early onshore LNG facilities featured steam turbine drivers, but in the 1980s facilities switched to industrial gas turbine drivers, which proved to be more cost- and plotefficient than steam turbines and supporting systems. As such, the steam turbine is not usually an economic driver selection for refrigeration compressors on an FLNG facility. Turboexpanders. Today’s turboexpander technology can deliver robust designs with a small footprint and high performance (with typical isentropic efficiencies of nearly 90%) for FLNG applications. However, a limit exists on the size of current companders, meaning that multiple compander trains would be used if a large facility is required. Liquefaction train size. For any type of LNG facility, the total facility capacity is often dictated by gas deliverability from dedicated gas fields or supply pipeline gas availability. The liquefaction train size is determined based on evaluating economies of scale for a large train vs. the availability benefits of multiple identical trains. For mid-scale FLNG, the facility may opt for either the single- or multi-train concept. In a multi-train concept, the facility has the capability to continue production when one of the trains is down for maintenance or an unexpected shutdown, while a single train may result in the lowest capital cost per unit of production. Utility systems. For an at-shore FLNG facility, common utilities (i.e., refrigerant storage, amine storage, cooling water, seawater desalination, fuel gas, instrument air, N2 supply and storage tank systems) can be installed onshore and connected to support operation of the process units. No absolute definition is given of which utilities can be installed onshore; it is a matter of site-specific economics, regulations and other factors. However, a few utility systems should remain onboard to improve the safety, reliability and energy efficiency of the FLNG facility. Cooling system. The interstage cooling for the refrigerant compressors on the nearshore FLNG facility can be provided by coastal seawater, freshwater or ambient air. For an at-shore

TABLE 2. Power output and efficiency of available gas turbines Heavy-duty-frame gas turbines Model

Siemens SGT 700

GE5D

GE 6B

GE 7EA

GE 9E

ISO power, MW

33.7

32.6

43.5

86.2

130.1

Efficiency, %

38.2

29.4

33.3

33

34.6

GE PGT25+G4

GE LM6000

GE LMS100

Siemens (RR) RB211 DLE

Siemens (RR) Trent DLE

ISO power, MW

34.3

43.9

100.2

33

54.2

Efficiency, %

41.2

43

44.1

40.5

43.9

Aeroderivative gas turbines Model

20 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

SPECIAL FOCUS: LNG TECHNOLOGY FLNG facility, air cooler(s) can be installed onshore where space is available, but higher-approach temperatures to ambient air will result in lower process efficiency. Within FLNG, design safety concerns are paramount; therefore, only an indirect closed cooling water loop would be considered for onshore air cooling service. The water cooling system can be an open or closed loop. In an open-loop system, a huge amount of seawater is drawn in, pumped directly through shell-and-tube heat exchangers (which are constructed of expensive, corrosion-resistant material, such as titanium) in the refrigeration systems, and then discharged back to the sea. Invariably, the nearshore FLNG facility will be located in shallow water close to shore, where warmer seawater will be encountered. This positioning will impact the process efficiency and adversely impact LNG production. Also, the water discharged into the sea must meet environmental requirements (in terms of maximum allowable return temperature, maximum temperature rise and residual chemical treatment) to minimize impact on ocean life nearshore. In the closed-loop system, fresh cooling water is pumped through a closed circuit to the plate-and-frame heat exchangers (which are constructed of less-expensive carbon steel material), in which heat is rejected to seawater. This option allows for the possibility of controlling the inlet and outlet temperatures of the heat exchangers, thereby stabilizing the process. The introduction of an intermediate loop ultimately increases the process to seawater approaching temperature, with a consequential increase in the weight and loss of process efficiency. Heating system. To improve the safety case of the FLNG facility, the heating system should not use fired heaters, if possible. As such, waste heat from gas turbine exhaust is recovered to provide process heating. The recovered heat is used mainly to meet the amine unit reboiler duty. Applying a heat-recovery steam-generation system on the exhausts of the gas turbines can also produce steam, which is used to generate additional electric power and to provide high-quality heat for regenerating molecular sieve beds. For operational simplicity, a heat transfer fluid is preferred to steam, as it avoids the operational complexity, large footprint, cost and weight of a steam generation system. Power generation and emergency power. Although power to nearshore/at-shore FLNG facilities can be provided from the local power grid or an efficient combined-cycle power plant located onshore, the facility should include an onboard power-generation system to provide essential power to support the entire facility. The facility should also have its own diesel generators as a minimum to ensure safe shutdown, or a larger generator as a backup to normal operation. Modularization. The modular building of the FLNG facility potentially reduces delivery time and moves costly site construction hours to fabrication yards, where cost and quality are best controlled. When practical, modules are designed up to 2,000 t to facilitate crane and transport availability. Larger modules (up to 4,000 t) have been installed on FLNG ships when the economic or schedule benefits were justified. Potentially significant cost savings and advantages exist in modular design, but the application to LNG facilities has not yet been successful in reducing the overall cost. However, innovation, development, improved project management and

redefinition of the workflow are necessary for more cost-effective modularization projects. These elements include offsite modularization at sites with lower labor costs and prefabrication of as many FLNG components as possible.

TAKEAWAY Evaluating criteria for the commercial acceptance of FLNG projects has shown that the dual-N2 expander liquefaction process is an appropriate selection for both offshore and nearshore small-scale FLNG projects where layout area is limited. N2 expansion technology is a proven, simple, compact solution, and is undoubtedly the safest design among the well-known refrigeration cycles. The SMR process, in which a low equipment count gives a more compact design, is well-suited for small- to mid-scale nearshore FLNG with better efficiency than the dual-N2 cycle, but with a higher risk profile. SMR provides operational simplicity while maintaining high process efficiency and low operating costs. The viability of LNG projects can be improved by maximizing production, but FLNG has a finite limit on barge size and equipment size. Therefore, increased production can be realized by switching to a more efficient DMR liquefaction process, even though this will introduce additional hydrocarbon inventory and increase the equipment count. Each FLNG facility will have unique aspects that require the process design to be tailored to achieve the optimum configuration. In addition to selecting the correct liquefaction capacity and technology, the execution aspects of the FLNG project (modularization strategy, construction philosophy, project execution and schedule) are also fundamental to achieving the overall objective of minimizing the cost/t of LNG. End of series. Part 1 appeared in the July/August issue. GP SAEID MOKHATAB is an internationally recognized gas processing consultant who has been actively involved in several large-scale gas field development projects, concentrating on design, precommissioning and startup of processing plants. He has presented on gas processing technologies worldwide and has authored or coauthored nearly 250 technical publications, including two well-known and frequently referenced Elsevier handbooks. He has held technical advisory positions for leading professional journals, societies and conferences in the field of gas processing, and has received a number of international awards in recognition of his outstanding work in the natural gas industry. SUKHPAL BASI is a chief technical adviser with the LNG technology group at KBR in London, UK. Having started at M. W. Kellogg in 1995, he has worked on LNG studies, FEED and EPC projects across a range of onshore and offshore applications throughout his career.

PHILIP HUNTER is the senior vice president for global LNG, FLNG and GTL technology and development at KBR in London, UK. He has more than 38 years of experience in the LNG business, and is considered one of the world’s leading technical experts in the natural gas liquefaction area. He has been involved in different sectors of major LNG and gas processing projects, from feasibility study, technology management, conceptual design, and project engineering and construction, to plant commissioning, startup and operation. He has authored or coauthored several technical papers on LNG, and given numerous technical presentations at prestigious international conferences. Gas Processing | SEPTEMBER/OCTOBER 2016 21

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SPECIAL FOCUS: LNG TECHNOLOGY

Use integrated analyses in design and operation of LNG systems

J. VALAPPIL, R. KUMAR and J. MUMM, Bechtel Oil, Gas and Chemicals, Houston, Texas

The LNG industry has seen tremendous growth in recent years, with several grassroots plants being built around the world. Many of these facilities are being constructed onshore, although the gas field may be located offshore—in some cases, hundreds of km away. The transportation of the gas and condensate from these offshore fields to the gas processing and LNG facilities introduces several challenges. The pipeline from the offshore field to the shore is characterized by long distances and elevation changes. Any operational disturbances in the upstream pipeline can adversely impact the operation of the gas processing or LNG facility. These impacts include the formation of slugs that can cause severe operational challenges for the inlet facilities of the onshore plant. Accounting for these upstream disturbances in the design of slug catcher and inlet facilities is critical to ensuring reliable facility operation. The slug catcher in the onshore facility is designed to handle the slugs that result from various operational scenarios. In addition, the inlet facilities, including the stabilizer system, are designed to account for the condensate liquids that must be processed. The proper designs of these units are important for the reliable operation of the entire facility under all conditions. The unit designs are a balance between operational flexibility and capital cost. A design that can handle the volume of slug for all operational scenarios without impacting the operation can be prohibitive from a capital-cost perspective. Therefore, some operational flexibility is sacrificed to design a facility with reasonable CAPEX. The slug catcher and stabilizer facilities are designed to handle the slug generated during various modes of operation and upset conditions. This is usually accomplished by flow assurance analyses in the early stages of the facility design (front-end engineering or earlier).1 This analysis uses a rigorous model to determine the effects of various operational scenarios. These scenarios include pigging, startup after shutdown, flow rampup, minimum turndown and other relevant scenarios. Transient simulation of the pipeline model is utilized to determine the slug arrival time, slug volume and other pertinent information. Additional benefits are gained by integrating a rigorous dynamic model of the pipeline with a model of the onshore facility. This model helps eliminate assumptions regarding the conditions at the interface between offshore and downstream facilities. The integrated model can also be valuable to establish operational strategies on both sides to account for the various upstream scenarios. This process helps verify the design of the

entire natural gas/LNG value chain (from wells to the delivery point). The tradeoffs established during design between capital cost and operational flexibility can be verified in a rigorous manner with this methodology. A case study is presented where a rigorous pipeline model is linked with a dynamic process simulation model of an onshore LNG plant. This integrated model is used to test various operational scenarios, including pigging and flow ramp-ups. The results are valuable for verifying the design of the individual equipment and units, and for identifying operational recommendations for the entire offshore/onshore gas transport infrastructure and LNG facility.

DESIGN AND OPERATION OF SLUG CATCHER AND STABILIZER SYSTEMS At onshore gas processing facilities, the first step is to separate the gas from the liquids. Slug catchers are used onshore to catch large slugs of liquid in pipelines and to provide a temporary storage buffer. These liquids are then delivered to the liquid processing facility at a suitable rate. These slug catchers may be of the vessel type or the finger type. The finger-type slug catchers are normally preferred for the larger volumes typical in natural gas and LNG facilities. FIG. 1 shows a typical inlet facility in an LNG plant with a three-phase slug catcher. The gas from the slug catcher is sent for further purification and liquefaction. The condensate stabilization system is designed to process the NGL from the slug catcher and produce condensate that meets Reid vapor pressure (RVP) and H2S specifications. The process schematic of a condensate stabilization unit typically Acid gas removal/ dehydration

Vapors Upstream pipeline feed

Gas to liquefaction

Slug catcher Liquids

Three-phase separation

Condensate stabilization unit

Aqueous liquids

Hydrocarbon liquids

FIG. 1. A typical slug catcher inlet facility configuration in a natural gas/LNG plant. Gas Processing | SEPTEMBER/OCTOBER 2016 23

SPECIAL FOCUS: LNG TECHNOLOGY used in LNG plants is shown in FIG. 2. Here, the hydrocarbon liquids from the slug catcher are let down to an intermediate pressure in the flash drum. The flash gases are recovered and sent back to the feed gas; then, the liquids are routed to the condensate stabilizer column. The overhead vapor from the stabilizer is compressed and returned back to the feed gas system and the amine unit. The bottom product from the stabilizer column is the condensate product. Upstream scenarios affecting onshore facilities. The design of the slug catcher and inlet facilities must account for the various operational scenarios and disturbances in the upstream pipeline feeding the LNG plant. Several scenarios impact the feed to onshore facilities, especially the slug volume: 1. Pigging. Pigging operation is necessary in multiphase flow systems for the purpose of inspection/maintenance and ensuring that the pipeline is clean. When pigging, a liquid slug is created ahead of the pig, which can be challenging for receiving facilities (i.e., the LNG plant). 2. Flow ramp-up. The increase of pipeline flow from turndown rates to full rate is a major disturbance for the onshore plant. This disturbance is due to the fact that the liquid holdup in the pipeline is higher at lower rates. As the flow is ramped up, the extra liquid is swept onshore, creating a surge of liquid at the onshore facility. To AGRU

Second-stage compressor

First-stage compressor Water

Condensate stabilizer

Slug catcher

Feed gas

C4+ from NGL unit

Water

Reboiler

MP flash drum

Stabilized condendate

0.9

0.9

0.8

0.8

0.7

0.7

0.6

Volume in slug catcher Net liquid flow into slug catcher Liquid flow into slug catcher Accumulated liquid Liquid flow out of slug catcher

0.5 0.4 0.3

0.6 0.5 0.4 0.3

Accumulated liquid

Removable liquid volume in slug catcher, flowrate

FIG. 2. A typical condensate stabilization system in a natural gas/ LNG plant.

0.2

0.2

0.1 0.0

0.1 0.0 50

0

5

10

15

20

25 Time, hr

30

35

40

45

FIG. 3. Various factors that go into sizing the slug catcher for a flow ramp-up case.

24 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

3. Gas field switchover. In several cases, a single LNG plant can be fed from multiple pipelines and wells. Switching of these pipelines or production wells can cause a compositional disturbance that changes the volume of liquid in the pipeline. 4. Onshore plant operational upsets. Various operational upsets in the onshore plant can cause the pressure at the slug catcher to vary significantly. This pressure variance can result in changes in the volume of liquid arriving at the onshore facility. Sizing of slug catcher and stabilizer systems. The sizing of

the slug catcher in a natural gas facility is intended for process stabilization, phase separation and storage of the liquid during various operational scenarios. The slug catcher is intended to dampen the effects of flowrate surges and deliver a steady supply of liquid to the production facilities. The size of the slug catcher is based on the liquid accumulation rate. This can be calculated as Liquid input mass rate – Liquid discharge mass rate = Liquid accumulation mass rate. The liquid input mass rate is the liquid flowrate entering into the slug catcher. It is dependent on the system operating philosophy, strategies and procedures adopted. Liquid discharge mass rate is a function of the capacity of the control valve and the stabilizer unit. The liquid accumulation rates can be calculated to determine an appropriate size for the slug catcher. Significant CAPEX savings can be realized by optimizing the size of the slug catcher that is required to handle the slugs and prevent upsets on the downstream equipment. Analysis using multiphase simulation tools early in the design phase can be valuable in optimizing the size. Stabilizer systems are sized to handle the liquids from the slug catcher to produce NGL products, and are sized for the normal liquid rates, within an appropriate margin. It is impractical to size the stabilizer systems for the maximum liquid rates expected during various upstream scenarios. It is possible to oversize the stabilizer system beyond the normal liquid rate. This scenario reduces the required size of the slug catcher, thereby establishing a tradeoff between the capital cost of the slug catcher vs. the stabilizer system. The removable liquid in the slug catcher (FIG. 3) is the volume above the minimum volume required for the slug catcher. The pipeline was initially operating at the steady-state flowrate. It was then shut down and gradually ramped up to the final flowrate. Tradeoff between capital cost and operability. The design of the slug catcher and inlet facilities should account for the various upstream operational scenarios that can impact the natural gas facility. In many cases, it is not economically feasible to size the slug catcher to handle the maximum slug arrival rate, considering all the scenarios. In these cases, the facility must sacrifice the operability to reduce the slug catcher size and, therefore, the capital cost. Placing an economic figure on the operability of the plant may prove difficult, but an attempt should be made to account for this figure in economic calculations. The overall design is, therefore, a tradeoff between capital cost and operability.

SPECIAL FOCUS: LNG TECHNOLOGY

DESIGN AND ANALYSIS METHODOLOGIES A brief overview of the design methodology normally used for multiphase pipeline systems is outlined in the following sections. Design using model of upstream pipeline. A rigorous sim-

ulation model of the upstream pipeline is normally utilized in the design of the slug catcher in the onshore plant. The model should offer the capability to predict the multiphase flow characteristics to accurately estimate the slug sizes. All the relevant aspects of the pipeline, including elevation changes, are included in the model. Normally, the model is developed as a part of the flow assurance study performed during the early design phase. The rigorous pipeline model is then used to run steadystate and transient simulations of the various aforementioned operational scenarios. The main information used for sizing is the volume of liquid flows coming from the pipeline to the slug catcher. In establishing the design and operation of the multiphase pipeline, valuable data can be obtained from such an analysis: 1. Steady-state pressure profiles. These profiles are critical to ensure that an adequate flow can be fed to the onshore facility at various pressures. The pressure drop is calculated using multiphase flow correlations.

2. Steady-state liquid holdups. The steady-state liquid holdup in the pipeline is important in establishing the liquid volume during turndown conditions. At lower flowrates, the liquid volume is greater. FIG. 4 demonstrates the relationship between flowrate and liquid holdup. 3. Slug sizes during various scenarios. Transient analysis of the upstream pipeline model is used to establish slug sizes during various scenarios, including pigging and flow changes. 4. Methodology for startup. Ramping up pipeline flow to a higher rate (low to full) is impractical due to the large volume of liquids that will be swept onshore. A more practical approach is to establish a staged startup, where the flow increases in a step fashion, holding at intermediate rates to sweep out the liquid. A maximum flow ramp-up rate from turndown conditions can also be established, beyond which the slug catcher capacity will be exceeded. Integrated upstream and onshore facility models. It is im-

portant to analyze the impact of upstream disturbance on the onshore equipment and facilities. Sequential simulations for the upstream pipeline and the downstream plant can be performed. The flowrates for the gas/condensate and compositions as a 1.2 .01 Scaled liquid holdup

This tradeoff is demonstrated in FIG. 4, which shows the amount of liquid holdup in the upstream pipeline as a function of flowrate. At lower flowrates, the holdups can be large. Performing pigging operation at these low flowrates, or ramping up from these low rates to full rate in a reasonable time, requires a slug catcher that is larger than an optimal one (FIG. 4). This is cost prohibitive, so the more practical approach of performing a pigging operation and ramp-up rate change (continuous, step-mode) should be performed to meet the slug catcher boundary conditions. During ramp-up operation in step mode, slug catcher inventory is reduced after each step. This adds to the ramp-up time, affecting operability. If permissible prior to the pigging or ramping-up operation, the liquid inventory in the system can be lowered by various modes of operation. To account for many of the upstream scenarios, the onshore facility can be prepared in a number of ways: • Prepare slug catcher. The slug catcher can be drained to the lowest level and then set at the proper operating pressure to reduce the impact. • Initiate condensate stabilization. The condensate stabilization unit can be ramped up to maximum production before the arrival of a slug. • Reduce time in turndown conditions. This decreases the liquid inventory in the pipeline, minimizing the slug volume arriving at the onshore facility. • Change compositions fed to the pipeline. If feasible, the upstream source composition can be modified to minimize the liquid inventory. This normally means using lean wells before pigging. • Tailor flow ramp-up to satisfy the slug catcher. The increase in flow from turndown or shutdown to full rate may be tailored to reduce the slug size.

Operational management

0.8 0.6 0.4

Slug catcher size

0.2 0.0 0

0.1

0.2

0.3

0.4

0.5 0.6 Scaled flowrate

0.7

0.8

0.9

1.0

FIG. 4. Liquid holdup in pipelines as a function of flowrate.

Pressure

Pipeline operation process simulation model

OPC connection

Server

Process simulation model Client

Temperature Phase flows P-F derivatives Component phase flows FIG. 5. Integrated upstream and onshore process models used for analysis. Gas Processing | SEPTEMBER/OCTOBER 2016 25

SPECIAL FOCUS: LNG TECHNOLOGY function of time can be entered into the downstream facility model to evaluate the effect on inlet facilities. The main disadvantage is that a sequential approach does not account for the interaction between the upstream and onshore units. For example, the change in flow to the onshore plant would, in turn, impact the volume of liquid and the composition of the pipeline flow. This method would only be appropriate if the pipeline flow is independent of the pressure, or if the pressure is well-controlled at the slug catcher. Benefits of integrating the upstream pipeline and downstream natural gas facility models include: 1. The integrated model captures the effect of upstream operations on the downstream facilities and control system, and vice versa, more accurately. This visibility eliminates the need to assume unrealistic boundary conditions. For example, the change in pressure at the slug catcher impacts upstream behavior. This pressure, in turn, is affected by what is delivered from the pipeline and the behavior of the LNG plant. 2. The integrated model allows for development of operational strategies for both upstream and downstream. 3. The integrated model allows for the development of control schemes that are best suited to both upstream

Normalized liquid volumetric flowrate

1.00 0.80 0.60 0.40 0.20 0.00 1,400

1,420

1,440

Time, min.

1,460

1,480

1,500

FIG. 6. Liquid volumetric flowrate into the slug catcher as the slug exits the pipeline during pigging.

Normalized liquid volumetric flowrate

1.00 0.80 0.60 0.40 0.20 0.00 0

500

Time, min.

1,000

1,500

FIG. 7. Liquid volumetric flowrate at the exit of the pipeline as the flowrate into the pipeline is increased.

26 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

and downstream operations. For example, the integrated model would be beneficial in designing and validating the compressor controls in the downstream units. Case study setup. The dynamic model for this analysis consisted of a rigorous offshore pipeline model linked to a separate dynamic process simulation model of the onshore LNG plant. Separate simulation software packages were utilized for the pipeline and LNG plant models to accurately capture the transient behaviors unique to each section, including transient multiphase flow within the pipeline, and operability and control within the LNG plant. A link between the two models was established using the native pipe unit operation in the plant process simulation model, which allowed the pipeline model’s outlet stream to be connected to the plant model’s inlet stream. To establish the connection, the pipeline model was first configured to be compliant with open platform communications (OPC). Variables controlling time and speed were exposed from within the pipeline model software to start/stop the model and to allow for synchronization of simulation time (FIG. 5).2 The scope of the pipeline model extended from the export of the offshore wells to the inlet of the slug catcher. Production from the wells was controlled via a source node (export) with a fixed flowrate and composition. The outlet boundary node (ending at the inlet to the slug catcher) was specified with a fixed pressure. For each case, the initial steady-state temperature, pressure, flow, compositional and phase profiles along the pipeline were configured by fixing the pressure at the outlet boundary node and running a steady-state preprocessor within the pipeline simulation software. During transient simulations, the pressure at the outlet node was calculated by the dynamic plant model and sent back to the pipeline model as a fixed value at the beginning of each integration step. The flow conditions along the pipeline were then calculated and tracked within the pipeline simulation software, and the resulting conditions at the outlet node were sent back to the plant model. During integration, both models were executed in parallel; as a result, there was a one-step time delay when passing boundary pressure and flow information. To prevent oscillation, pressure-flow derivatives were communicated between the models to predict the boundary response during transient conditions. Temperature and phase flow information were also retrieved by the plant model to set the composition and temperature of the inlet. For the flow ramp-up, pressure change and pigging simulation cases, a non-compositional pipeline model was used. In these cases, a fixed total composition was assumed in the pipeline model, meaning that specific component fractions were neither calculated nor tracked. For the composition change case, a separate pipeline model with composition tracking was used. Upstream operational scenarios. Three cases were investigated as part of this analysis: pigging, flowrate ramp-up and inlet composition changes. Pigging is a common practice in pipeline systems, wherein a pig is propelled through a pipeline by the pressure of the product flow for inspection, maintenance and cleaning. During pigging, a liquid slug is often created ahead of the pig, which can result in a sudden surge of liquid entering the slug catcher. Typically, slug catchers are designed to receive a cer-

SPECIAL FOCUS: LNG TECHNOLOGY

1.00 Normalized liquid volumetric flowrate

tain net liquid volume over a specified period of time. However, if they are under-designed, the arrival of a slug can result in liquid overflow to those downstream processes where liquid could disrupt production. For this analysis, the arrival of such a slug during pigging was simulated to determine the impact on the receiving facilities, and to verify that the sizing of the slug catcher was sufficient to receive and process the full volume of the slug. As the pig traveled through the pipeline, a liquid slug formed ahead of the pig. FIG. 6 shows the total liquid volumetric flowrate into the slug catcher. As the slug exited the pipeline, the flowrate increased and was maintained at a high rate for nearly 1 hr. The area under this curve represents the total volume of the slug. In response, flow to the condensate stabilization system was increased to maximum capacity. To analyze the capability of the stabilization system and to process the sudden increase in liquid, the relative volatility of the condensate product was tracked via the RVP. A temporary spike occurred in the RVP as the stabilizer capacity was suddenly increased. Such a disruption could have been minimized by ramping up the stabilizer capacity at a slower rate; for this case, the slug catcher would have provided a sufficient buffer volume to allow for more gradual corrective action on the stabilization system. When the flow through a pipeline is ramped up, the increase in flowrate sweeps out the liquid within the pipeline and results in large slugs of liquid volume entering the slug catcher. To determine the impact of a ramp-up and the volumes of resulting slugs, a dynamic simulation was performed in which the inlet flowrate was increased at a predefined maximum rate from turndown conditions to the maximum production rate. For the flowrate ramp-up case, the total mass flowrate at the inlet of the pipeline was ramped up from turndown conditions to full capacity over a period of 1 hr. FIG. 7 shows the liquid phase flowrate at the exit of the pipeline. As the flowrate increased, residual condensate built up within the pipeline was pushed out, resulting in a large surge of liquid into the slug catcher. For this case, the peak liquid flowrates into the slug catcher were lower than in the pigging case; therefore, flow to the stabilizer was increased more gradually. Eventually, the stabilizer reached maximum capacity well before the increase in the slug catcher liquid level subsided. Inlet composition changes. The transition of the plant inlet composition from the pipeline was also simulated to analyze any potential operational issues. Compositional change can occur due to switching or changing of production wells with different well fluid compositions. For this case, the inlet composition from the well was assumed to transition from a case with a high nitrogen content to a case with a higher concentration of heavier components, meaning more condensate formation along the pipeline. After changing the feed composition, the fraction of liquid in the feed increased as heavier hydrocarbons condensed along the pipeline. FIG. 8 shows the volumetric flowrate of the liquid phase arrival at the inlet to the slug catcher. The wave of liquid took approximately 27 hr to traverse the entire length of the pipeline, resulting in the sudden increase shown in FIG. 8. This increase in liquid volume was found to be within the capacity of the slug catcher, and no action was required.

0.80 0.60 0.40 0.20 0.00 1,400

1,500

1,600

1,700 Time, min.

1,800

1,900

2,000

FIG. 8. Liquid volumetric flowrate into the slug catcher after the feed composition is changed to one with heavier components.

TAKEAWAY The impact of upstream operation on the onshore facilities can be significant and must be taken into account in design and operation. The sizing of the facility is a tradeoff between capital cost and operability. This balance can be established early in the design phase to eliminate any issues during the facility lifetime. Analysis using an integrated model is beneficial in various ways, including for verification of the integrated operation of the entire facility. Operating procedures can be developed to handle a variety of upstream transients using these analyses and methodologies. GP LITERATURE CITED Hagesaether, L., K. Lunde, F. Nygard and H. Eidsmoen, “Flow assurance modeling: Reality check and aspects of transient operations of gas/condensate pipelines,” Offshore Technology Conference, Houston, May 2006. 2 Heum, J. R., “Steady state vs. dynamic simulation and OLGA,” e-Field Seminar, Dubai, UAE, November 1, 2009. 1

JALEEL VALAPPIL is a principal process engineer and team lead for Bechtel Oil, Gas & Chemicals’ advanced simulation group in Houston, Texas. His areas of expertise include process engineering, simulation, control and optimization. He is responsible for developing and deploying advanced technical solutions during design, commissioning and operation of various Bechtel projects, including LNG. Dr. Valappil holds a BS degree from the Indian Institute of Technology in Kharagpur, India and a PhD in chemical engineering from Lehigh University in Bethlehem, Pennsylvania. RAKESH KUMAR is an engineering supervisor in Bechtel Corp.’s pipeline division. His field of experience includes process plant operation, construction, precommissioning, commissioning, detailed design and operations support for system startup. His responsibilities include flow assurance, thermal hydraulic design and assessment of gas, liquid and multiphase systems. Mr. Kumar holds a BS degree in chemical engineering from Panjab University, Chandigarh, India and an MS degree in chemical engineering from Lamar University in Beaumont, Texas. JESSE MUMM is a process engineer in Bechtel Oil, Gas & Chemicals’ advanced simulation group. His responsibilities include the hands-on development of dynamic plant models and utilizing simulation for the design, validation and optimization of LNG and gas processing plants. He holds a BS degree in chemical engineering from the University of Minnesota and has several years of experience in dynamic process simulation. Gas Processing | SEPTEMBER/OCTOBER 2016 27

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SPECIAL FOCUS: LNG TECHNOLOGY

Impacts of benzene and piperazine concentrations on LNG plant capacity K. K. HWANG and S. KIM, SK E&C, Houston, Texas

Benzene emissions can have an effect on LNG plant capacity when a vent gas incinerator is not present. Here, a typical amine sweetening unit for acid gas removal, featuring a typical lean feed gas composition, is modeled using commercial amine sweetening software to calculate annual benzene emissions rates. The maximum allowable benzene content in a given feed gas rate is set by US Environmental Protection Agency (EPA) benzene emissions limits. Two parameters—benzene concentration in the feed gas stream and piperazine content in the methyldiethanolamine (MDEA) solution—are used to determine LNG plant capacity. LNG plant capacity is described by an empirical curve fit equation. This equation provides a method to predict LNG plant capacity when a vent gas incinerator is not present. Understanding the effects of benzene emissions rates on LNG plant capacity provides operators and designers a basis for the interpretation and manipulation of a broad range of feed gas compositions in LNG plant design. Parameter overview. Benzene, toluene, ethylbenzene and xylene (BTEX) are present in natural gas streams and are picked up in the exit CO2 stream of the amine unit. Typically, the vent gas is incinerated through a thermal oxidizer to meet EPA BTEX emissions limits. BTEX components are listed by the EPA in the Clean Air Act of 1990 as some of the 188 hazardous air pollutants. The EPA sets a standard of 25 tpy for total aromatic compounds emitted by any given plant. A 10-tpy limit also exists on each individual aromatic compound that can be emitted. The reasons1 for the restrictions are as follows: 1. Benzene is a human carcinogen that promotes leukemia 2. Toluene exposure can lead to reproductive or developmental effects 3. Ethylbenzene affects the blood, kidneys and liver 4. Xylene exposure can affect the central nervous system, leading to respiratory and cardiovascular problems. Therefore, most gas operations incinerate BTEX, which are then absorbed in the amine unit and eventually released to the atmosphere to resolve the disposal issue of BTEX components in the vent gas stream. However, the vent gas incinerator (thermal oxidizer) is a major source of flue gas, such as NOx and CO. The NOx and CO emissions rates from the thermal oxidizer could pose a major issue in obtaining local environmental air permits. In addition, the vent gas incineration is a thermal oxidation process in which the BTEX components are combusted at a temperature

of 1,500°F. Since most of the vent gas stream is water-saturated CO2 , a considerable amount of fuel gas is required for this operation. In other words, the vent gas incineration process can be operating cost-prohibitive from a fuel-usage standpoint.2 The other challenge posed by the thermal oxidizer comes from a construction viewpoint. According to National Fire Protection Association Standard 59A, process equipment containing LNG, refrigerants, flammable liquids or flammable gases shall be located at least 50 ft from the ignition source. The separation distance could provide another problem, since most small-scale LNG feed gas treatment systems inclusive of amine units are designed by modular fabrication. The modular fabrication provides benefits such as an accelerated schedule, lower installed cost and increased reliability for LNG plant operations. Therefore, efforts to replace the vent gas incinerator with adsorption technology have drawn attention, and are under study. In this work, the authors provide the example of the removal of vent gas incineration from amine regeneration in the feed gas treatment portion of an LNG plant, and explain how this modification would affect LNG plant capacity. More specifically, the authors review the configuration of the feed gas treatment system and explain how benzene emissions rates from the amine regenerator can impact the capacity of an LNG plant without a vent gas incinerator. In this study, the authors use a commercial process simulator to model a typical acid gas removal unit (AGRU) without a vent gas incinerator. The model allows for the predicted measurement of annual benzene emissions rates for the two parameters, which are benzene contents in the feed gas stream and piperazine concentrations in the amine solution. Requirements of feed gas treatment. In an LNG plant, natural gas from the gathering system must be treated before liquefaction can take place. Four major categories of contaminants in the raw feed gas are considered potentially damaging to the liquefaction process: mercury, heavy hydrocarbons, acid gas and water. Mercury is known to cause stress cracking in brazed aluminum heat exchangers that are used in the cryogenic section. To prevent this stress cracking from occurring, the typical mercury specification for LNG is set at 10 Ng/m3. Mercury can be easily removed by conventional methods, such as a non-regenerable metal oxide guard bed. In addition, a sulfur-impregnated mercury-sulfur guard bed is used to meet the EPA H2S emissions limit. Gas Processing | SEPTEMBER/OCTOBER 2016 29

SPECIAL FOCUS: LNG TECHNOLOGY The optimal location of the lead-lag beds is upstream of the AGRU, since the vent gas stream with H2S from the amine regenerator is not supposed to be incinerated. Also, the heavy hydrocarbon (HHC) removal unit should be placed upstream of the amine unit so that a significant amount of BTEX is not released into the acid gas along with the CO2 from the amine regenerator. CO2 removal from natural gas using amine to very low levels (< 50 ppmv) is required to prevent freezing in the cold box. In a typical commercial amine process, an aqueous alkanolamine solution is in counter-current contact with natural gas containTABLE 1. Acid gas stream composition and operating conditions Gas component

mol%

C1

94

C2

2

C3+

0.3

Benzene

4 ppmv

CO2

2

Inert

Balance

Operating condition

Measurement

Temperature, °F

95

Pressure, psig

900

Flowrate, MMscfd

320

TABLE 2. Amine solution composition and operating conditions Amine solution

wt%

MDEA

45

Piperazine

5

Water

50

Operating condition

Measurement

Temperature, °F

113

CO2 loading to MDEA solution, mol

0.58 Vent gas to ATM

Sweet gas to dehydration

ing CO2 in an absorber column. Finally, water must be removed from the gas stream prior to liquefaction to avoid freezing in the cold box. These treating facilities are an essential part of LNG plants, helping ensure reliable LNG production. As discussed in the previous section, this case study highlights the AGRU to address benzene emissions without a vent gas incinerator. Modeling of AGRU. In this study, a typical amine sweetening unit (FIG. 1) is modeled using a commercial process simulator. An aqueous MDEA solution is in counter-current contact with natural gas containing CO2 in an absorber column. The MDEA reacts with the acidic CO2 gas to form a dissolved salt, allowing purified natural gas to exit absorber. The rich amine solution is regenerated in a stripper column to produce an acid gas stream concentrated with CO2, and is eventually vented to atmosphere without incineration. The lean solution is then cooled and returned to the absorber, allowing the process to repeat in a closed loop. The acid gas stream, MDEA solution and operating parameters used for the study are summarized in TABLES 1 and 2. Typically, feed to LNG plants is composed primarily of methane, together with ethane, propane, butane and heavier components. A typical lean gas feed composition is used for the higher content of CO2 gas in the study. All stream properties of the amine unit are calculated using amine sweetening software in a commercial process simulator. Benzene content and composition of amine solution are varied for the sensitivity studies, irrespective of the performance credit of the HHC removal unit. Results and discussion. Benzene is the least soluble aromatic

component in LNG and poses the highest risk of freezing in the cold box. Therefore, benzene content is used as a target specification for the operation of the HHC removal unit in the LNG plant. The benzene content in the AGRU feed can vary, depending on the performance of the HHC removal unit in conjunction with CAPEX and OPEX. Benzene contents in the vent gas stream from the amine regenerator and sweet gas stream from the amine contactor are shown in FIG. 2 as a function of the benzene concentration of the AGRU feed gas stream.

Amine contactor Acid gas from HHC unit

Amine regenerator

Steam Rich amine flash

Benzene content in outlet streams of amine unit, ppmv

80

60 50 40 30 20 10 0

Amine solution FIG. 1. AGRU process scheme.

30 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

Benzene content in vent gas Benzene content in sweet gas

70

0

5

25 10 15 20 Benzene concentration of feed gas stream, ppmv

FIG. 2. Benzene content in outlet streams of the AGRU.

30

35

SPECIAL FOCUS: LNG TECHNOLOGY mately 8 ppmv of benzene is theoretically soluble in liquid methane at the typical liquefaction temperature of –250°F. Consequently, the maximum benzene content of 4.5 ppmv meets the LNG specification and the EPA benzene emissions limit, as well. The benzene content is determined by the performance of the HHC removal unit, which is located upstream of the AGR unit. For instance, the performance of temperature swing adsorption removing HHC inclusive of benzene is dependent on adsorbent volume (CAPEX) and regeneration cycle (OPEX). Specifically, the benzene content of the AGRU feed gas stream could be reduced at the cost of CAPEX and OPEX. Eventually, the reduced benzene content can increase LNG plant capacity, meeting the EPA benzene emissions limit. To investigate the dependency of LNG plant capacity on benzene concentration, the same calculation methods described in FIG. 3 are conducted, varying the feed gas flowrate. In this calculation, 20

16 14 12 10 8 6 4 2 0

0

2

3 4 5 6 7 Benzene concentration of feed gas stream, ppmv

8

9

-100 Estimated annual benzene emissions EPA emissions limit

18

-120 -140

14

-160

12

-180

Temperature, °F

16

10

-200

8

-220

6

-240

4

-260

2

-280

0

1

FIG. 4. Calculated yearly benzene emissions at 4 × 320 MMscfd.

20

Benzene emissions, tpy

Estimated annual benzene emissions EPA emissions limit

18

Benzene emissions, tpy

Bullen et al.3 reported that the solubility of benzene in 50 wt% MDEA solution at 110°F should be approximately 0.017 scf/gal. However, as shown in FIG. 2, a very small amount of benzene is absorbed in the MDEA solution due to low partial pressure of benzene in the feed gas stream. By increasing the benzene concentration of the feed gas stream, benzene contents in sweet gas and vent gas streams display an increase. From the benzene contents in the vent gas stream of the amine regenerator, represented by a diamond symbol in FIG. 2, annual benzene emissions to the atmosphere are calculated based on a feed gas flow of 320 MMscfd. FIG. 3 reports the calculated annual benzene emissions rate in tpy, in terms of benzene concentration of the feed gas stream, for the case study of 320 MMscfd of feed gas flow. The calculated annual benzene emissions rate and the EPA emissions limit are designated by the blue diamond symbol and the red dotted line, respectively. The EPA limit of 10 tpy is taken to indicate that approximately 18 ppmv (maximum) of benzene content is allowed in the feed gas stream. Since LNG plants are typically designed with multiple trains for larger production rates or operational flexibility to meet client needs and commercial requirements, the feed gas flow of the case study is extended to 4 × 320 MMscfd. As expected, the benzene loading at the feed gas flow of 4 × 320 MMscfd shifted to a lower concentration in comparison to the feed gas flow of 320 MMscfd to meet the EPA emissions limit, as shown in FIG. 4. In other words, the HHC removal unit upstream of the AGRU must reduce the benzene content to approximately 4.5 ppmv (maximum) to meet the EPA benzene emissions limit. As stated earlier of the study result, there is a high risk of precipitation of benzene in the LNG product, due to its higher freezing point compared to cyclohexane and other aromatic components. Therefore, the maximum benzene content of 4.5 ppmv, meeting EPA limits, must be examined so that freezing does not occur in the LNG product. Experimental solubility data for benzene in methane from Neumann4 are shown in FIG. 5. According to the solubility data, approxi-

-300 0

5

10 15 20 25 Benzene concentration of feed gas stream, ppmv

FIG. 3. Calculated yearly benzene emissions at 320 MMscfd.

30

35

0

10

20

30 40 50 60 70 Concentration of benzene in CH4, ppm (mol)

80

90

FIG. 5. Solubility of benzene in methane (reproduced from Neumann4). Gas Processing | SEPTEMBER/OCTOBER 2016 31

SPECIAL FOCUS: LNG TECHNOLOGY the benzene emissions rate of 9 tpy is taken as a threshold point determining the benzene content in the feed gas stream, considering a 10% safety margin on the EPA emissions limit. FIG. 6 reports estimated LNG plant capacity as a function of the benzene concentration of the feed gas stream. The LNG plant capacity decreases as benzene content in the feed gas stream increases to meet the EPA emissions limit. The dependency of the LNG plant capacity (Ca) on the benzene concentration (Bz) of the feed gas stream is well described (R2 = 0.99) by the empirical curve fit (solid line) shown in Eq. 1: Ca =

5,326 Bz

(1)

where the fit constant has a unit of ppmv × MMscfd. The other parameter affecting the benzene emissions rate of the vent gas is the amine solution. Amine’s reactivity with CO2 and benzene’s solubility in amine solution could factor into the 12,000

Estimated LNG capacity, MMscfd

10,000

8,000

6,000

4,000

benzene emissions from an AGRU without a vent gas incinerator. MDEA is preferably used as an absorbent in applications involving CO2 removal from natural gas in LNG production due to its low regeneration energy. However, its reaction with CO2 is extremely slow, and the absorption process is controlled entirely by resistance to mass transfer in the solvent phase. For this reason, 5% of piperazine in MDEA solution was used in the modeling study of the AGRU. Piperazine is a cyclic diamine that reacts with CO2 approximately 10 times faster than MDEA. To further investigate piperazine’s effect on the benzene emissions rate of the vent gas stream, a commercial amine sweetening simulation was conducted, varying piperazine concentration in the MDEA solution and maintaining the same benzene concentration of 4 ppmv in the feed gas stream. FIG. 7 shows the benzene content in the vent gas stream as a function of piperazine concentration. According to the simulation result, benzene content in the vent gas stream decreases as piperazine concentration increases. From the benzene contents in the vent gas stream shown in FIG. 7, yearly benzene emissions to atmosphere are calculated based on a feed gas flow of 4 × 320 MMscfd and benzene content of 4 ppmv in the feed gas stream. FIG. 8 shows that the yearly benzene emissions rate decreases as piperazine loading increases. As discussed previously, piperazine is a promoter absorbing CO2 rather than benzene. As a result, benzene content in the vent gas stream decreases as piperazine’s strength in amine solution (50 wt%) increases. In other words, it is likely that benzene’s solubility decreases as TABLE 3. Piperazine concentrations and curve fit constant

2,000

Piperazine concentration, wt%

Curve fit constant, A, ppmv × MMscfd

0

3,150

0 0

2

4 6 8 Benzene concentration of feed gas stream, ppmv

10

12

FIG. 6. Estimated LNG plant capacities.

5

5,326

8

6,509

10

7,292

16

45

Estimated annual benzene emissions EPA emissions limit

35 Benzene emissions, tpy

Benzene content in vent gas stream, ppmv

14

25

12

10

8

15 6

5

0

0

1

2

3

4 5 6 7 8 Piperazine content in MDEA solution, wt%

9

10

FIG. 7. Benzene content in the vent gas stream at 4 ppmv of benzene in the feed gas stream.

32 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

11

0

1

2

3

4 5 6 7 8 Piperazine content in MDEA solution, wt%

9

10

FIG. 8. Estimated yearly benzene emissions at 4 ppmv of benzene in the feed gas stream.

11

SPECIAL FOCUS: LNG TECHNOLOGY the MDEA fraction in the amine solution decreases. The study result is enhanced by reference study results3,5 showing that the solubility of benzene in the MDEA solution is reduced by as much as 30% as MDEA concentration decreases by 10%. More detailed analysis on benzene’s solubility in piperazine should be conducted in a separate study. To investigate the dependency of LNG plant capacity on benzene and piperazine concentrations, the same calculation method and the same EPA emissions criteria described in FIG. 6 are conducted, varying the feed gas flowrate. As shown in FIG. 9, LNG plant capacity increases as piperazine concentration in the MDEA solution increases at the same level of benzene concentration of the feed gas stream. The dependency of LNG plant capacity (Ca) on benzene concentration (Bz) of the feed gas stream is well described (R2 = 0.99) by the empirical curve fit (Ca = A/Bz) and summarized in TABLE 3 for all piperazine loading rates. The curve fit constant linearly correlated with the piperazine concentration. The slopes and intercept of the linear regression were 418.9 MMscfd . ppmv . wt%–1, and 3,150 MMscfd . ppmv–1, respectively (R2 = 0.99). By compiling the fit constant into the empirical curve fit (Ca = A/Bz), the LNG plant capacity can be expressed in terms of benzene and piperazine concentrations using the curve fit shown in Eq. 2: Ca =

3,150 + 418.9Pi Bz

(2)

designers to interpret and manipulate a broad range of feed gas compositions in LNG plant design. GP LITERATURE CITED Majumdar, D., A. K. Mukherjeea and S. Sen, “BTEX in ambient air of a metropolitan city,” Journal of Environmental Protection, Vol. 2: 2011. 2 Morrow, D. and K. Lunsford, “Removal and disposal of BTEX components from amine plant acid gas streams,” Proceedings from the 76th annual GPA convention, 1997. 3 Bullin, J. and W. Brown, “Hydrocarbons and BTEX pickup and control from amine systems,” Proceedings from the 83rd annual GPA convention, 2004. 4 Neumann, A., R. Mann and W. Von Szalghary, “Solubility of solid benzene in liquid hydrocarbons,” Kaeltetech Klim, Vol. 24, 1972. 5 Critchfield, J., H. Holub and F. Mather, “Solubility of hydrocarbons in aqueous solutions of gas treating amines,” Proceedings from the Laurance Reid Gas Conditioning Conference, 2001. 1

KENNETH K. HWANG is a registered professional chemical engineer with more than 17 years of experience in the oil and gas industry, and in research. He received his PhD in chemical engineering from Texas A&M University. Dr. Hwang works for SK E&C as a senior process engineer. He has expertise in natural gas processing, LNG, LPG and ethylene plant design. SUCKHEE KIM is a principal process engineer with more than 22 years of engineering, procurement and construction (EPC) and front-end engineering design (FEED) experience in the oil and gas industry. Mr. Kim’s specific areas of technical expertise include LNG liquefaction technology, NGL processing and natural gas treatment. He works for SK E&C as a principal process engineer.

To the authors’ knowledge, neither benzene nor piperazine concentration dependency on LNG plant capacity has been previously reported. Ultimately, with knowledge of the benzene concentration (0 < Bz ≤ 8 ppmv) and piperazine concentration (0 ≤ Pi ≤ 10 wt%), Eq. 2 provides a powerful method for making predictions of LNG plant capacity where a vent gas incinerator is not present. In other words, the contribution of benzene emissions without a vent gas incinerator can also be extended to typical gas processing and treating plants. Ultimately, understanding the effects of benzene emissions rates on LNG plant capacity provides a basis for operators and TRI-CON

9,000 8,000 7,000 Estimated LNG capacity, MMscfd

TRI-CHECK

TRI-BLOCK

Piperazine = 0 wt% Piperazine = 5 wt% Piperazine = 8 wt% Piperazine = 10 wt%

6,000

WWW.ZWICK-VALVES.COM

5,000

THE NEXT GENERATION

4,000 3,000 2,000 1,000 0 0

2

4 6 8 Benzene concentration of feed gas stream, ppmv

FIG. 9. Estimated LNG plant capacities at various benzene and piperazine loadings.

10

TRI-SHARK

TRI-CONTROL

TRI-JACK

Gas Processing | SEPTEMBER/OCTOBER 201633

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PIPELINES AND INFRASTRUCTURE

Optimize an integrated natural gas production and distribution network

D. ALUMA and N. THIJSSEN, Shell International Exploration and Production, The Hague, The Netherlands; K. M. NAUTA and C. C. PANTELIDES, Process Systems Enterprise Ltd., London, UK; and N. SHAH, Imperial College London, London, UK

Gathering significant volumes of natural gas usually involves connecting to wells in different oil and/or gas fields that can be spread across vast geographical areas. Supplying processed gas and associated liquids, such as LPG and NGL, to consumers often requires an extensive network of degassing stations, compressor stations, pipelines and gas processing plants, as well as storage facilities and shipping terminals for the liquids. Natural gas networks have peculiar characteristics that complicate supply chain management. For example, there is usually no intermediates storage. For associated natural gas, it is generally impossible to curtail well production, as this leads to loss of oil revenue. Therefore, a drop in consumer gas demand or an outage in the gas network typically results in significant gas flaring. The penalty for a constrained or inefficient gas network is environmentally and economically severe, more so than for liquids- or solids-based process networks. The use of network modeling and optimization technologies for strategic decision-making yields substantial benefits, not only in economic and environmental terms but also in an improved understanding of the interaction between the various components of the process and the overall business.1 Key benefits of the application of systematic optimization to natural gas network production include: • Increased profitability: Given the tight market conditions, even small improvements in the gas supply chain can have an impact on profitability, and could be critical to long-term business viability. • Improved reliability: A network optimizer provides a tool that can be used, for example, to rapidly reallocate production upon equipment failure to satisfy

customers and honor contractual delivery commitments. • Investment planning: Comprehensive network models can be used to gain insights into network bottlenecks and to enable planning for debottlenecking or expansion projects. • Flare reduction: Network optimization techniques can be used to minimize environmental and economic penalties from inefficient operations. Even greater benefits can be realized by considering multiple production periods through the simultaneous optimization of production, sales and inventory over consecutive time periods, with demands, prices and costs varying across periods.

APPROACHES TO SUPPLY NETWORK OPTIMIZATION Conventional approaches to the optimization of large supply networks tend to rely on simple models of the individual nodes (e.g., production facilities and processing plants) in these networks, often taking the form of simple (frequently linear) relations between the flowrates of the various materials entering and leaving each node. While this greatly simplifies the solution of the underlying mathematical optimization problem, it may be problematic in important ways. In particular, the resulting solutions may not be implementable in practice, as they may exceed plant capabilities. This problem may be addressed by performing the optimization calculations in a conservative manner—for example, by introducing safety margins to ensure that the solution does not violate critical constraints. Alternatively, some appropriate adjustments may be applied, often at the local level of individual processing plants, a posteriori to the

“optimal” solution to restore the feasibility of any constraints that it violates. However, such pragmatic adjustments almost always lead to suboptimal solutions. Given the substantial revenue flows in such large networks, these solutions may translate into a significant loss of opportunity. An alternative approach to natural gas supply chain optimization is to use a higher level of physical detail in describing the operation of the individual production and processing nodes, thereby ensuring that any obtained solution satisfies all important constraints on the operation of plant equipment. Until recently, this approach was considered to be impractical because of two significant obstacles. First, the construction of detailed models for the individual plants is a non-trivial exercise. Secondly, the resulting optimization problem was often outside the capabilities of existing numerical solvers. However, detailed mathematical modeling is increasingly being used for the simulation and optimization of individual plants, which means that the required plant models are already available in many cases. Also, the continual evolution of computer hardware and process modeling technology is now bringing the solution of models of the required size within the scope of the available tools. A key development in this context has been the significant progress made over the past decade by equation-oriented process modeling frameworks.2 These frameworks now allow large-scale models, comprising fairly detailed, smaller models of individual equipment items within wide system envelopes, to be constructed and reliably solved with minimal user intervention. Moreover, coupling these models with rigorous mathematical optimization solvers allows the effective and efficient exploration of deGas Processing | SEPTEMBER/OCTOBER 2016 35

PIPELINES AND INFRASTRUCTURE cision spaces spanning large numbers of decision variables. In practical terms, it is now possible to perform the optimization of models comprising several hundreds of thousands of nonlinear equations that are subject to dozens of decision variables.

INTEGRATED NATURAL GAS PRODUCTION AND PROCESSING NETWORKS The technological developments outlined in the previous section, coupled with improving process knowledge and developments in the global natural gas inWest Qurna field (three compressor stations)

13 12

61

9

25

62

11

23

dustry, are now driving rapid innovation in natural gas/NGL supply chain modeling and optimization. This work describes how detailed models of such networks are used to optimize profitability, as well as other business objectives, creating significant economic benefit without the need for additional CAPEX. The natural gas network of Basrah Gas Co. (BGC)a is used as a case study. The BGC network in southern Iraq, shown schematically in FIG. 1, processes associated natural gas from four oilfields situated in West Qurna, North and South Rumaila and Zubair. BGC produces dry gas and

10 24

8

North Rumaila field (five compressor stations)

6 2 4 29C 34

27

7 5 1 3

GTU

North Rumaila NGL plant (slug catchers, liquids strippers, booster stations, acid gas removal units and NGL fractionators)

26 Pipelines legend Raw gas Dry gas Condensate Broad cut C3 C4 Gasoline Slops

28

42B 91A

39 South Rumaila field (four compressor stations)

40 37 38 35 36 47

Petchem plant and other users 14

33 Zubair field (three compressor stations and two dehydrator stations)

29B 15 V1

42A

18 20

91 58

19 21 22

43

90

10 31 29A

41

32

49

Khor al Zubair slug catcher

44 45 46 47 Khor al Zubair plant (two NGL trains and three LPG trains)

Um Qasr storage terminal 50 52 54 56 51 53 55 57

59 Um Qasr marine terminal 60

FIG. 1. BGC gas network schematic.

FIG. 2. BGC network model topology showing top-level models and their connectivity. Each model is actually an entire plant (see FIGS. 3–5).

36 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

LPG for the domestic Iraqi market. LPG and condensate are also exported after satisfying local demand. At each oilfield, gas is received at different pressures following its separation from oil. The gas is then compressed at a number of compressor stations and transported via an extensive (650-km) pipeline network to an NGL processing plant at North Rumaila and to NGL and LPG plants at Khor al Zubair. An important characteristic of the network is that gas from the South Rumaila field may be sent to either of the two processing locations, providing an additional degree of overall operational flexibility. Moreover, the two processing facilities are coupled via a broadcut liquid stream comprising propane, butane and gasoline, that is sent from the North Rumaila plant to Khor al Zubair for further processing. Dry gas from both plants is introduced into the domestic Iraqi network via a gas transmission unit (GTU), while liquid products from Khor al Zubair are sent to the Um Qasr storage and marine terminals.

NETWORK MODEL A detailed model of the BGC natural gas network was developed using a commercially available, equation-based process modeling and optimization tool. The model was constructed in a hierarchical manner, with models of the individual production and processing facilities embedded within a top-level model shown in FIG. 2. The unit operation models were taken from standard libraries supplied within the environment. The overall model comprises approximately 280,000 nonlinear equations. To provide suitable fidelity for optimization calculations, detailed pipeline models, compressor models with performance curves and multicomponent vapor-liquid-equilibrium (VLE) models were used throughout. Murphree tray efficiency was applied to the tray-by-tray distillation column to align predicted column performance with actual data from plant operations. Component and mixture properties were supplied by the built-in physical property package, using the PengRobinson equation of state to describe the pressure-volume-temperature (PVT) properties of the natural gas system. To establish a reliable basis for the optimization studies, the network model

PIPELINES AND INFRASTRUCTURE was validated against operational data from the BGC network. Compressor station modeling. Detailed steady-state models were implemented for the compressor stations, utilizing the actual compressor performance curves. Interpolation methods are used to determine polytropic head and efficiency from the flow data. The compressor station models also contain air-cooled heat exchangers, vaporliquid separators and a triethylene glycol (TEG) dehydration unit to meet the pipeline specification for H2O. The TEG unit was modeled as a simple component splitter to remove water from the compressed gas. FIG. 3 shows the model topology for a three-train compressor station. Pipeline modeling. The pipelines are modeled using a distributed model to describe pressure drop in the line and heat transfer to surroundings. This level of detail is required to properly account for multiphase flow and liquid dropout in the extended pipeline system. Pressure drop is calculated based on the Haaland friction factor correlation, with appropriate corrections applied to account for pipeline bends, elevations and inline deposits. NGL plant modeling. The NGL plant comprises a slug catcher, gas booster units, an acid gas removal unit (AGRU), cooling loops, a low-temperature separator and a deethanizer. Detailed models were used for the cooling loops, lowtemperature separator and deethanizer. The AGRU was not modeled, as the H2S treatment function does not significantly affect network flowrates and production. Slug catchers. The slug catcher removes the transient slugs that develop in the pipeline and ensures a steady flow of gas to the NGL plant. The liquids that form in the pipeline as a result of ambient cooling are also removed. The operation is modeled as a simple flash drum affecting the necessary vapor/liquid separation. Liquid stripper units. The liquid from the slug catcher goes to a liquid stripper unit, which causes any gas in the liquid to desorb, thereby increasing gas recovery while stabilizing the resulting liquid stream. The latter is pumped to broadcut storage, where it mixes with the broadcut (C3+) from the deethanizer bottoms.

The liquid stripper unit is modeled as a trayed column with a reboiler. The column model incorporates detailed phase equilibria. Cooling loops and low-temperature separator. The gas cooling cycle was modeled in detail (FIG. 4). The cycle involves a number of recycle loops comprising chillers and heat exchangers for the integration of heat between the inlet and outlet streams of the low-temperature separator (LTS). After cooling to about –30°C, the fluid flows into the LTS, where it is flashed. The gas phase is dry gas (mostly methane), which is fed into the export grid after it is used to cool the inlet stream. The liquid phase from the LTS is also used to cool the inlet streams before flowing to the deethanizer fractionator. Deethanizers. Dry gas (methane and ethane) is extracted from the top of the deethanizer distillation column while the C3+ is drawn from the bottom. An impor-

tant constraint during network optimization is the avoidance of flooding in the various columns. Failure to take proper account of this constraint may result in significant over-estimation of the plant’s processing capacity. The authors’ optimization calculations required the actual vapor velocity to not exceed 85% of the flooding velocity, as determined via the fair correlation.

LPG PLANT MODELING The LPG plant at Khor al Zubair comprises three parallel trains, each including depropanizer and debutanizer columns (FIG. 5). Each depropanizer receives two feeds (one with and one without reheating) of C3+ components from the NGL trains. A condensed, propane-rich stream is extracted from the top, while the C4+ components are drawn from the bottom and go to the corresponding debutanizer. The latter produces a butane-rich top product

FIG. 3. Three-train compressor station model.

FIG. 4. NGL plant model showing cooling cycles, LTS and deethanizer. Gas Processing | SEPTEMBER/OCTOBER 2016 37

PIPELINES AND INFRASTRUCTURE and a natural gasoline (C5+ components) bottom product. As in the case of the NGL plant deethanizers, flooding constraints were implemented in both columns. The top products from the depropanizers in the three trains, and the top and bottom products from the debutanizers, are combined to form the propane, butane and gasoline product streams of the LPG plant.

OPERATIONAL OPTIMIZATION A number of different operational optimization cases were studied. The full network model described in the previous section was used in each case, and the optimization was carried out with respect to 77 decision variables, including feed gas rates from oil production, gas-to-NGL plant splits, distillation column reflux ratios and column boilup ratios. The optimization also took account of 32 constraints, including product specifications and equipment limits. The optimization studies were closely related to the business objectives listed

in the beginning of this work, and covered both normal and abnormal operating scenarios. Optimization under normal operating scenarios. Several normal opera-

tional scenarios were considered: • Maximize operating profit, taking account of the values of the various products (dry gas, propane, butane and natural gasoline) based on market-driven contractual sales prices, as well as variable production costs, including raw gas and utilities (electricity, and heating and cooling utilities) • Maximize total production (i.e., the total mass flowrate of all product streams), irrespective of any economic factors • Maximize specific product yields (i.e., the ratio of that product’s mass flowrate divided by the overall mass flowrate of gas entering the system)

TABLE 1. Operational optimization results for normal operation Base value

Objective Maximize profitability, $yr–1

Optimal value

Improvement over base operating point, %

Not disclosed

+4.9%

Maximize total production, kg s–1

346.9

359.2

+3.5%

Maximize propane yield, kg kg–1

0.23

0.34

+47.8%

Minimize flaring, kg s–1

9.85

1.26

–87.2%

FIG. 5. LPG plant showing depropanizer and debutanizer columns.

Initial product inventory

Period 1

Product inventory at end of period 1

Product sales

FIG. 6. Multi-period optimization schematic.

38 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

Period 2

Product sales

Product inventory at end of period 2

• Minimize flaring (i.e., the difference between the total mass flowrate of gas entering the system and that of all product streams). The performance of the network under optimal operation was compared with that of a base point representing design operating conditions. The magnitudes of the resulting improvements in operation are listed in TABLE 1. Profitability optimization results in a potential 4.9% profit increase compared to the base case. A more detailed analysis of the results indicates that this increase arises largely from the optimized routing of raw gas from the different compressor stations to the two NGL plants and, to a lesser but significant extent, from applying optimal reflux and boilup ratios in the columns. Significant improvements are also obtained under the other normal operating scenarios. In particular, there is scope to increase total production by 3.5%. Additionally, the results show that it is possible to increase the yield of a key product by nearly 50%, if required, and flaring can be reduced by 87% if this is posed as the primary objective. Optimization under abnormal operating scenarios. The following “abnor-

mal” scenarios were also considered: • Maximize production under equipment failure, resulting in the unavailability of one of the two NGL trains at the Khor al Zubair plant • Maximize production under field decline, resulting in lower gas flowrates being received from the West Qurna field. TABLE 2 compares the optimal production rates with base values representing the rates that would result from the above failure scenarios leading to deviation from normal design operation. The potential benefits from a rigorous mathematical optimization approach are even more pronounced than those under the normal scenarios shown in TABLE 1. In such cases, the optimizer provides a tool that can be used to significantly reduce the effects of abnormal situations on the network’s operation.

MULTI-PERIOD OPTIMIZATION OF NETWORK PROFITABILITY The profitability maximization study considered in the previous section focused on a single operating period with given product prices and unit costs. The production rate for each product was as-

PIPELINES AND INFRASTRUCTURE

Optimal value

Production improvement over existing operating practices, %

Maximize production under equipment failure (one KAZ NGL train down), kg s–1

233.5

295.3

+26.5%

Maximize production under field decline, kg s–1

230.5

343.2

+48.9%

TABLE 3. Profitability over two-period operation Integrated network profitability (relative units) Period 1

Period 2

Total

Profitability improvement over base case, %

1

1.19

2.19

0%

Sequential optimization

1.02

1.21

2.23

2%

Multi-period optimization

0.26

2.07

2.32

6.2%

Base operation

sumed to be equal to its sales rate; therefore, the potential of building up product inventories, or of using existing inventories to support sales, was not considered. In this section, the simultaneous optimization of production, sales and inventory over consecutive time periods is considered, with demands, prices and costs potentially varying across periods. The goal is to maximize the overall profit for the entire time horizon by varying, in each period, the process operating variables, sales rates (subject to contractually committed levels and maximum market demands) and the inventory levels (subject to storage capacities). Overall, this multi-period optimization (MPO) problem, illustrated schematically in FIG. 6, aims to combine operational optimization with commercial planning using the full network model. To illustrate the potential benefits of an MPO approach, a two-period problem is considered, with product prices in the second period being higher than in the first. The optimization makes use of a restricted subset of the decision variables, excluding the column reflux and boilup ratios. Limitations on product storage tank capacities are taken into account, as well as the constraints introduced in the previous section. TABLE 3 compares network profitabilities achieved via 1) a base case operation, 2) application of the optimization approach described in the previous section to each of the two periods independently (“sequential optimization”), and 3) the MPO approach. It is evident that MPO results in significantly improved profitability (6.2% over the base case) compared with that achieved via the

sequential application of even a sophisticated single-period optimization approach. In particular, when product prices are expected to be higher in future periods, the MPO approach deliberately reduces product sales rates, instead increasing inventories that can be sold more profitably in subsequent periods. In contrast, the sequential optimization approach myopically optimizes the network operation based only on the present period’s product demands and prices.

TAKEAWAY Significant economic benefits can be achieved via the holistic optimization of natural gas production and processing networks. The use of detailed, physics-based models of processing plants in this context significantly increases the probability of the solutions obtained being both optimal and practically implementable. Moreover, notwithstanding the underlying mathematical complexity, such optimizations are now feasible using current process modeling technology. Even higher benefits may be achieved via the extension of this approach to supply chain optimization over multiple periods, thereby bridging commercial planning with operational decisions. However, the computational cost increases significantly with the number of periods, and specialized MPO solvers are likely to be necessary for solving problems of practical interest. GP

romance.hoerbiger.com

Base value

Objective

...R...O...MA...NCE

TABLE 2. Operational optimization results for abnormal scenarios

ONLINE EXCLUSIVE Complete literature cited, author’s note and author bios available at GasProcessingNews.com Gas Processing | SEPTEMBER/OCTOBER 2016 39 325037_Hoerbiger_AZ_Romance_PTQ_59x270mm_RZ.indd 1

11.03.16 11:08

‫ﻳﻘﺎم ﺗﺤﺖ اﻟﺮﻋﺎﻳﺔ اﻟﺴﺎﻣﻴﺔ ﻟﻔﺨﺎﻣﺔ اﻟﺮﺋﻴﺲ ﻋﺒﺪ اﻟﻔﺘﺎح اﻟﺴﻴﺴﻲ‬ ‫رﺋﻴﺲ ﺟﻤﻬﻮرﻳﺔ ﻣﺼﺮ اﻟﻌﺮﺑﻴﺔ‬ HELD UNDER THE PATRONAGE OF HIS EXCELLENCY PRESIDENT ABDEL FATTAH EL SISI PRESIDENT OF THE ARAB REPUBLIC OF EGYPT

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GAS COMPRESSION

Super-compact BOG recondensing system minimizes equipment lifecycle costs K. HAYASHI, K. YARIMIZU and S. FURUTANI, JFE Engineering Corp., Tokyo, Japan

Processing boiloff gas (BOG) is a challenge for the hydrocarbon processing industry. The industry has been addressing the issue by installing BOG recondensing systems, which reduce power consumption by the BOG compressor.1 To further minimize the lifecycle cost of BOG processing equipment, the authors propose a considerable reduction in the size of the BOG recondenser. Here, an innovative BOG recondenser with a volume of only 2% of conventional ones is introduced. Patents were filed for the recondenser. The unit’s recondensing performance and benefits, as well as the control system that ensures the safe and stable operation of the entire system, are also described. Since it began importing LNG from Alaska in 1969, Japan has become the world’s biggest importer and consumer of LNG. Japan also has the largest number of LNG importing terminals in the world, along with rich experience in the design, construction and operation of these terminals. During this period of nearly 50 years, the country has confronted various technical issues regarding LNG, and it has demonstrated the methods of overcoming these problems to the industry. The technology introduced here is one of the products based on this experience.

Several types of BOG recondensing systems have been commercialized. The packed-bed type (FIG. 2) is most widely applied. However, it requires a large and heavy recondenser, which consists of two portions with a skirt. The upper portion is the packed-bed section where the BOG is liquefied as it contacts with the subcooled LNG. The lower portion is the holdup section, which serves as a suction drum for the secondary pump. The skirt is required to provide the net positive suction head (NPSH) for the secondary pump. Due to its design, this type of recondenser is very high, as shown in FIG. 3. This height requires operators to climb up platforms for regular inspections and maintenance, which increases their workload and risk. It also requires a large crane whenever the operators need to replace the packing. Another type of recondenser is the heat exchanger type (FIG. 4), which is more common in Japan. LNG flows into the shell, the BOG flows into the tubes, and the gas is condensed

Conventional BOG processing: Power consumption. The

To pipeline

tanks, piping and equipment that contain LNG are always subject to heat input because the temperature of the liquefied gas, at approximately –160°C, is far below ambient temperature. LNG is circulated through the piping and equipment to remove the input heat and maintain the equipment at cryogenic temperature. As a result, the heat is gathered to the LNG tank, and the LNG in the tank is partly evaporated by the collected heat and the direct heat input at the tank. This evaporated gas is the BOG (FIG. 1). Due to its evaporation in the tank, the BOG is discharged to the connected pipeline to keep the tank under the design pressure. The BOG compressor pressurizes and sends the BOG out to the pipeline. However, the compressor requires huge power, as the pipeline operating pressure is generally very high (5 MPa or more). This increases the OPEX of the terminal. Conventional BOG recondensing system. A BOG recondensing system, which liquefies BOG by utilizing the cold energy of LNG, will greatly reduce the power consumption of the compressor, as pumping the BOG in the liquid phase requires considerably less power than compressing the same in the gas phase.

BOG compressor (high pressure) LNG carrier

P Primary pump

LNG vaporizer Secondary pump

FIG. 1. Process flow of LNG import terminal without BOG recondensing system. BOG compressor (high pressure) BOG compressor (low pressure)

P

BOG recondenser (packed bed)

Primary pump

LNG vaporizer Secondary pump

FIG. 2. Process flow of packed-bed BOG recondensing system. Gas Processing | SEPTEMBER/OCTOBER 2016 41

GAS COMPRESSION by subcooled LNG. This type of recondenser is also large because it needs a wide heat transfer area. The large shell and the dense-allocated tubes inside add to its weight. As shown in FIG. 4, this type of recondenser requires two additional pumps in the system. One is for pumping LNG to the heat exchanger, and the other is for discharging the liquefied BOG to the secondary pump. This configuration will increase OPEX.

The dominant advantage of the new recondenser is its compactness. The recondenser is smaller in size and weight than the conventional types, which brings benefits in the construction and operation of the recondensing system.

Inspection and maintenance are also challenging with this type of configuration. The operator must discharge all LNG in the shell to detach the tube bundle for inspection. Although LNG and BOG are clean and noncorrosive, authorities require scheduled inspections in some places, such as Japan. New BOG recondensing system: Process description. This recondenser is a vertical pipe structure (FIG. 5). LNG flows into the recondenser, and BOG is introduced into the LNG through proprietary nozzles inside. The nozzles convert the BOG into very fine bubbles, and the bubbles are recondensed immediately. The nozzles were developed through trial-and-error laboratory tests with water and steam. Significant experiments were carried out to find the optimum arrangement of the nozzles and other parameters. The dominant advantage of the new recondenser is its compactness. As shown in FIG. 6, the recondenser is smaller in size and weight than the conventional types. In a typical case (BOG recondensing capacity = 14 tph), the diameter is only 0.3 m and the length is 3 m. A heat exchanger-type BOG recondenser with the same capacity is 1 m in diameter and 12 m in height. This means that the new recondenser is only 2% of the volume of the conventional recondenser, and is approximately 1/15 the weight. The significant compactness of the new recondenser design brings benefits in the construction and operation of the reconBOG compressor (high pressure) BOG compressor (low pressure) BOG recondenser (static mixer)

P Primary pump

LNG vaporizer Secondary pump

FIG. 5. Process flow of new BOG recondensing system.

FIG. 3. Packed-bed BOG recondenser. BOG compressor (high pressure) BOG compressor (low pressure) BOG recondenser (heat exchanger)

P Primary pump

LNG vaporizer Secondary pump

FIG. 4. Process flow of heat exchanger BOG recondensing system.

42 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

FIG. 6. Conventional recondenser designs vs. new recondenser design.

GAS COMPRESSION densing system. In the aforementioned typical case, CAPEX is reduced by approximately 20%, and the footprint is decreased by approximately 40%. OPEX is expected to be 60% lower than that of the heat exchanger-type recondenser because fewer pumps are required. The new recondenser system is also less labor-intensive in many ways. It is virtually maintenance-free because of its simple internal structure. The control system is straightforward and does not require manual operation. System configuration and control. The configuration of the system is shown in FIG. 7. The low-pressure BOG compressor sends the BOG to the recondenser, and the recondenser mixes the BOG with LNG to liquefy it. The BOG/LNG mixture pump sends the mixture to the secondary pump. The system control is quite simple. The LNG flowrate to the BOG recondenser is maintained at the set value by controlling the outlet flowrate of the BOG/LNG mixture pump. The set value is not affected by the BOG flowrate and is unchanged in principle. The value is reduced only when the LNG regasification rate becomes lower than the initial set value. The maximum allowable BOG flowrate (i.e., the amount of BOG that can be recondensed) is calculated from the process data shown in FIG. 8. The flowrate of the low-pressure BOG compressor is always kept at or below the calculated allowable value. If the BOG generated is greater than the allowable value, the excess BOG is sent to the pipeline by the high-pressure BOG compressor.

1

BOG flowrate, tph

14 at 40°C

LNG flowrate, tph

140 at –157°C

Pressure, MPaG

and 10) was installed in the Chita-Midorihama works of Toho Gas Co. Ltd. Its specifications are shown in TABLE 1. The commissioning of the plant was completed in June 2016, and its performance was confirmed to be in accordance with the design. Details of the test results are provided in the following sections. BOG recondensing performance. The actual BOG recondensing performance, represented by the ratio of LNG/ BOG mass flowrate, is far better than the typical heat exchanger-type recondenser and close to the theoretical limit (FIG. 11). The theoretical limit is the minimum LNG/BOG ratio where all BOG can be recondensed completely, assuming 100% heat exchanger efficiency. The dots connected by the red lines in FIG. 11 indicate the state in which the operating LNG/BOG ratio is gradually reduced toward the theoretical limit while maintaining other conditions, such as pressure, temperature, 9

LNG regasification rate

LNG flow control (FLC)

Process data (P, TG, TL)

LNG composition Calculated maximum allowable BOG flowrate BOG compressor flowrate control

TABLE 1. Specifications of the first commercial plant (operating targets) Number of system

First commercial plant. The first commercial plant (FIGS.

FIG. 8. Basic control philosophy of new BOG recondensing system.

0.76

Allowable N2 content in BOG, mol%

20 (maximum)

BOG compressor (high pressure)

BOG compressor (low pressure) FG

P

BOG recondensing system

Newly developed BOG recondensor

TG

LNG vaporizer

FLC P Primary pump

TL

BOG/LNG mixture pump Secondary pump

FIG. 7. System configuration of new BOG recondensing system.

FIG. 9. Overview of new BOG recondensing system. Gas Processing | SEPTEMBER/OCTOBER 2016 43

GAS COMPRESSION etc. The series of dots demonstrate that the new recondenser, as a device, has the potential to perform at conditions very close to the theoretical limit. Another beauty of the recondenser is the very small pressure loss. If the pressure loss is large, the pressure at the outlet of the recondenser will be reduced, and the LNG sensible heat for the recondensation will also be reduced. Therefore, the BOG recondensing capacity will be diminished. Safe operation. The behavior of the entire recondensing system was tested in case the operating conditions deviated suddenly. When the BOG flowrate or the LNG pressure changed suddenly, the system maintained safe and stable operations.

The emergency shutdown sequence was also tested. The system is designed to initiate an emergency shutdown when the BOG is not completely recondensed and discharged downstream of the recondenser. The test conditions were created by manually reducing the LNG flowrate. The behavior of the system conformed to design specifications, and the system made an emergency shutdown in a safe manner. Noise and vibration have been negligible throughout the commissioning and initial commercial operation. The continued absence of noise and vibration problems will minimize any potential issues related to these elements throughout the recondenser’s long lifecycle. Takeaway. The BOG recondensing system that has been developed is much smaller and lighter than conventional recondensing systems. This advantage contributed to the reduction of lifecycle cost and footprint for the recondenser. Additionally, its maintenance-free structure and simple control philosophy will minimize the workload of the operators. The first commercial plant has demonstrated excellent BOG recondensing performance that is close to the theoretical limit. The plant has also demonstrated continuous, safe and stable operation, even during a sudden fluctuation in operating conditions. Noise and vibration were measured at negligible levels. GP ACKNOWLEDGMENTS The authors express their gratitude to the codeveloper, Toho Gas Co. Ltd., which cooperated and contributed greatly to the design, construction and commissioning of the first commercial plant.

1

FIG. 10. The new BOG recondenser.

KANETOSHI HAYASHI is group manager of the research center of engineering innovation for JFE Engineering Corp. He joined JFE in 1990 and has been responsible for the research and development (R&D) of processes and equipment in the energy and environmental areas since that time. His areas of expertise include heat transfer, mass transfer and fluid dynamics. His R&D project involvement has included thermal energy storage, energy conversion and multi-phase/phase-change processes. He was a visiting researcher for Argonne National Laboratory in the US between 1997 and 1999. He earned his MS degree in mechanical engineering from the University of Tokyo, Japan in 1990.

12

Typical performance of heat exchanger type

Ratio of LNG/BOG mass flowrate (operation)

11

10

Theoretical limit

9

8

O.67 MPaG O.72 MPaG

7

O.76 MPaG O.80 MPaG

6 6.0

6.5

LITERATURE CITED Lemmers, S. P. B., “Simplify BOG recondenser design and operation—Part 1,” Gas Processing, June 2014; and Lemmers, S. P. B., “Simplify BOG recondenser design and operation—Part 2,” Gas Processing, August 2014.

7.0 7.5 8.0 Ratio of LNG/BOG mass flowrate (theoretical limit)

8.5

9.0

FIG. 11. BOG recondensing performance of the first commercial plant.

44 SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

KEIJI YARIMIZU is a project engineer in the energy division of JFE Engineering Corp. and is responsible for project management of LNG import terminals. He has 20 years of experience in the design of underground LNG storage tanks and in the process engineering and commissioning of LNG import terminals. He has been involved in various LNG import terminal projects for power and city gas companies in Japan. Mr. Yarimizu earned his MS degree in mechanical engineering from Tokyo Metropolitan University, Japan in 1995, and joined JFE in the same year. SHIGEYA FURUTANI is a senior engineer in the energy division of JFE Engineering Corp., responsible for business development in the oil and gas midstream area, including the LNG value chain. He joined JFE in 1994 and has over 20 years of experience in engineering, construction and project management of onshore and offshore pipeline projects in Japan, Vietnam, Saudi Arabia and Hong Kong. He earned his MS degree in civil engineering from Kobe University, Japan in 1994, and is a registered professional engineer in Oregon, US.

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GasProcessingNews.com 46  SEPTEMBER/OCTOBER 2016 | GasProcessingNews.com

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Register Today! November 1–2, 2016 Houston, Texas WGLConference.com/2016

2016 WOMEN’S

Explore the Dynamics Impacting the Industry + Get Valuable Advice on How to Become a Leader Dear Gas Processing Readers, I’d like to invite you to join me and hundreds of the energy industry’s leading and aspiring professionals for the 13th Women’s Global Leadership in Energy Conference (WGLC) and explore the challenges faced by all segments of the industry (upstream, midstream and downstream). We will also discuss the opportunities arising from these challenges for those stepping into leadership positions to affect positive change and drive performance. The event’s opening session will focus on the present state of the industry, as well as provide a look into the future. Dr. Kenneth B. Medlock III, Senior Director, Center for Energy Studies at Rice Adrienne Blume University’s Baker Institute for Public Policy, will commence the session with a presentation titled, “What does the future of energy look like?” Following his presentation, I will moderate a panel discussion on the State of the Industry: Upstream, Midstream and Downstream. Joining me are panelists Marianne Kah, Chief Economist, Corporate Planning and Development, ConocoPhillips; Frank Cassulo, President, Chevron Pipeline; and Liz McDaniel, Vice President Environmental, Health & Safety, Polyurethanes Division, Huntsman. In our panel discussion, we will address key topics influencing the state of the upstream, midstream and downstream energy industries, including: • Opportunities for energy efficiency and cost savings in oil and gas production and processing operations • When a shift in the energy pricing down cycle may occur, and the impacts that could follow from it • Growth in market share for natural gas, and the changing landscape of energy usage around the world • Emerging technology applications for smaller-scale and modularized energy production and distribution Additional session topics include: tackling challenges to create opportunity; supplier diversity; women in the workplace; networking effectively; creating opportunities for women and minorities; big data initiatives; stepping into a leadership position; building a more sustainable tomorrow; and women in the C-suite. You can download the full agenda and register at www.WGLConference.com/2016. If you have questions or would like more information, please contact Tranessa Hunt, Events Program Manager, at [email protected] or +1 (713) 520-4470. Sincerely, Adrienne Blume Editor Gas Processing Executive Editor Hydrocarbon Processing

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