Indian Oil Corporation Ltd. Company History The Indian Oil Corporation Ltd. operates as the largest company in India in terms of turnover and is the only Indian company to rank in the Fortune "Global 500" listing. The oil concern is administratively controlled by India's Ministry of Petroleum and Natural Gas, a government entity that owns just over 90 percent of the firm. Since 1959, this refining, marketing, and international trading company served the Indian state with the important task of reducing India's dependence on foreign oil and thus conserving valuable foreign exchange. That changed in April 2002, however, when the Indian government deregulated its petroleum industry and ended Indian Oil's monopoly on crude oil imports. The firm owns and operates seven of the 17 refineries in India, controlling nearly 40 percent of the country's refining capacity.
IndianOil Major Projects IndianOil continues to lay emphasis on infrastructure development. Towards this end, a number of schemes have been initiated with increasing emphasis on project execution in compressed schedules as per world benchmarking standards. Schemes for improvement and increased profitability through debottlenecking / modifications / introduction of value added products are being taken up in addition to grassroots facilities. Project systems have been streamlined in line with ISO standards. 1) GRASSROOTS REFINERY PROJECT AT PARADIP (ORISSA) 2) RESIDUE UPGRADATION AND MS/HSD QUALITY IMPROVEMENT PROJECT AT GUJARAT REFINERY 3) NAPHTHA CRACKER AND POLYMER COMPLEX AT PANIPAT (HARYANA) 4) MS QUALITY UPGRADATION PROJECT BARAUNI REFINERY (BIHAR) 5) MS QUALITY UPGRADATION PROJECT AT GUWAHATI REFINERY (ASSAM) 6) MS QUALITY UPGRADATION PROJECT AT DIGBOI REFINERY (ASSAM) 7) DADRI-PANIPAT R-LNG SPUR PIPELINE 8) PANIPAT REFINERY EXPANSION FROM 12 MMTPA TO 15 MMTPA 9) BRANCH PIPELINE FROM KSPL, VIRAMGAM TO KANDLA 10) DIESEL HYDRO-TREATMENT (DHDT) PROJECT AT BONGAIGAON REFINERY (ASSAM) 11) MS QUALITY UPGRADATION PROJECT AT BONGAIGAON REFINERY (ASSAM) 12) PARADIP-NEW SAMBALPUR-RAIPUR-RANCHI PIPELINES 13) DE-BOTTLENECKING OF SALAYA-MATHURA CRUDE PIPLEINE 14) INTEGRATED CRUDE OIL HANDLING FACILITIES AT PARADIP
Down the memory lane
1958
Indian Refineries Ltd. formed in August with Mr. Feroze Gandhi as the Chairman.
1959
Indian Oil Company Ltd. established on 30th June 1959 with Mr. S. Nijalingappa as the Chairman. MV Uzhgorod carrying the first parcel of 11,390 tonnes of Diesel for IndianOil docked at Pir Pau Jetty in Mumbai on 17th August 1960. Agreement for supply of Kerosene and Diesel signed with the then USSR Construction of Barauni Refinery commenced
1960
1962 1964
1967
1998
Barauni Refinery commissioned Indian Refineries Ltd. merged with Indian Oil Company with effect from 1st September, 1964, and Indian Oil Company renamed as Indian Oil Corporation Ltd. Haldia Barauni product pipeline commissioned. 1965 Barauni-Kanpur product pipeline and Koyali- Ahmedabad product pipeline commissioned Haldia-Barauni crude oil pipeline completed.
BARAUNI REFINERY Barauni Refinery was built in collaboration with Russia and Romania. Situated 125 kilometres from Patna, it was built with an initial cost of Rs 49.40 crore. Barauni Refinery was commissioned in 1964 with a refining capacity of 1 Million Metric Tonnes per Annum (MMTPA) and it was dedicated to the Nation by the then Union Minister for Petroleum, Prof. Humayun Kabir in January 1965. After de-bottlenecking, revamping and expansion project, it's capacity today is 6 MMTPA. Matching secondary processing facilities such Resid Fluidised Catalytic Cracker (RFCC), Diesel Hydrotreating (DHDT), Sulphur Recovery Unit (SRU) have been added. Theses state of the art eco-friendly technologies have enabled the refinery to produce environment- friendly green fuels complying with international standards. IndianOil is the highest ranked Indian company in the prestigious Fortune 'Global 500' listing, having moved up 19 places to the 116th position in 2008. It is also the 18th largest petroleum company in the world.
Awards/Accolades Barauni Refinery achieved safety award in gold category of “Green Tech Foundation Safety Award” on 04.05.09. BR bagged 2nd prize in Golden Jubilee Indian Oil Album in Aug 09. Barauni Refinery accredited in Oct 09 with prestigious “Jawaharlal Nehru Centenary Awards” (3rd prize) for Energy Performance in Refinery for the year 2008-09 by MoPNG. o Suggestion Fortnight declared and inaugurated by ED, BR on 09.12.09. o Barauni Refinery has been accredited first prize in the refinery sector for “National Energy Conservation Awards-2009” by Ministry of Power. Award received by ED, BR on 14.12.09. Barauni Refinery was initially designed to process low sulphur crude oil (sweet crude) of Assam. After establishment of other refineries in the Northeast, Assam crude is unavailable for Barauni . Hence, sweet crude is being sourced from African, South East Asian and Middle East countries like Nigeria, Iraq & Malaysia. The refinery receives crude oil by pipeline from Paradip on the east coast via Haldia. With various revamps and expansion projects at Barauni Refinery, capability for processing high-sulphur crude
has been added — high-sulphur crude oil (sour crude) is cheaper than lowsulphur crudes — thereby increasing not only the capacity but also the profitability of the refinery. Crude oil is separated into fractions by fractional distillation. The fractions at the top of the fractionating column have lower boiling points than the fractions at the bottom. The heavy bottom fractions are often cracked into lighter, more useful products. All of the fractions are processed further in other refining units. Different boiling points allow the hydrocarbons to be separated by distillation. Since the lighter liquid products are in great demand for use in internal combustion engines, a modern refinery will convert heavy hydrocarbons and lighter gaseous elements into these higher value products. Oil can be used in a variety of ways because it contains hydrocarbons of varying molecular masses, forms and lengths such as paraffins, aromatics, naphthenes (or cycloalkanes), alkenes, dienes, and alkynes. While the molecules in crude oil include different atoms such as sulfur and nitrogen, the hydrocarbons are the most common form of molecules, which are molecules of varying lengths and complexity made of hydrogen and carbon atoms, and a small number of oxygen atoms. The differences in the structure of these molecules account for their varying physical and chemical properties, and it is this variety that makes crude oil useful in a broad range of applications. Once separated and purified of any contaminants and impurities, the fuel or lubricant can be sold without further processing. Smaller molecules such as isobutane and propylene or butylenes can be recombined to meet specific octane requirements by processes such as alkylation, or less commonly, dimerization. Octane grade of gasoline can also be improved by catalytic reforming, which involves removing hydrogen from hydrocarbons producing compounds with higher octane ratings such as aromatics. Intermediate products such as gasoils can even be reprocessed to break a heavy, longchained oil into a lighter short-chained one, by various forms of cracking such as fluid catalytic cracking, thermal cracking, and hydrocracking. The final step in gasoline production is the blending of fuels with different octane ratings, vapor pressures, and other properties to meet product specifications.
HIGHLIGHTS Barauni Refinery achieved highest ever crude processing of 6.2 MMT (outlook) during the year. Previous best was 5.94 MMT during the year 2008-09. Achieved highest ever Low Sulphur crude processing of 5.55 MMT (outlook) during the year surpassing the previous best of 5.16 MMT during the year 2008-09. Achieved highest ever CRU throughput of 288.3 TMT (outlook). Previous best was 277 TMT during the year 1999-00.
Achieved highest ever RFCCU throughput of 1.497 MMT (outlook) during the year surpassing previous best of 1.454 MMT during the year 2008-09.
Annual Production (Outlook for the year 2009-10) Product Qty (TMT) Previous best (TMT) LPG 291.8 284.5 (2008-09) MS (Total) 763.3 703.2 (2008-09) SKO 954.6 894.3 (2005-06) HSD (Total) 3119.7 3087.7 (2008-09) RPC 207.8 176.9 (2007-08) FO 26.9 -
ATMOSPHERIC AND VACUUM DISTILLATION UNIT (AVU-I / II)
INTRODUCTION There are two Atmospheric and Vacuum Distillation Units in Barauni Refinery numbered as AVU-I and AVU-II, each were designed for 1 MMT/year crude processing. Subsequently another distillation unit without vacuum distillation facility was added. This unit was designed for 1 MMT/year of crude and known as AU-3. Crude Processing capacity of both units AVU-I & AVU-II was increased to 1.6 MMT/year by HETO project (Heat Exchanger Train optimization) in 1990. The above modification (HETO project job was designed by EIL (Engineer's India Limited) and fabrication/erection job was completed by M/s. PETHON ENGG. LTD, Mumbai. The units were again revamped in 1998 (M & I) when the capacity was expanded to 2.1 MMT/year of each of the two units. Through these units were designed on the basis of evaluation data of Naharkatiya crude, presently the units have switched on to imported crude due to none availability of Assam crude.
PROCESS DESCRIPTION
Crude oil (imported) is received from Haldia by pipeline and is pumped from tanks through Heat Exchangers after exchanging heat with various hot stream, the crude streams attain a temperature of approx. 120oC to 130oC. After attaining temperature about 120oC to 130oC the two crude flows combine together and enter in Desalter for separation and removal of water and salt. Bi electric desalter is having two energised electrodes. A distributor head splits crude between the upper and lower pair of electrodes. Crude oil separated from water between the centre and lower electrodes passes through the upper electrode in a converging countercurrent flow with the separating water from upper set of electrodes. This creates a second washing zone for half of the feed in a strong electrical field thereby causing maximum salt removal efficiency. The two desalter in AVU-I &II are PETRECO BIELECTRIC type which were commissioned in the year 2001. POST DESALTER:- At the outlet of Desalter there are two booster pumps which boost up the crude at discharge pressure around 15 kg/km2 . Pre-topping column has 20 Trays (All valve trays with a bed of packing between 9th & 10th tray) and operates operating conditions Present Pressure Kg/cm² (g) Top temperature (Deg. C) Bottom temperature (Deg. C)
As per design
2.4
4.0
112 – 118
120
210 to 215
240
Pretopped crude stream passes through heat exchangers. After exchanging heat with various hot products the pretopped crude flows combine and it is segregated again near furnace in two pass flows before entering the atmospheric heater for further heating and finally fed to 6th tray of main column through two entry nozzles at 340oC.The Furnace is provided with Air Preheater. Main Fractionating Column has 43 double pass valve trays. Following are the operating parameters of the main column. Present
As per design
Pressure Kg/Cm²(g)
0.3-0.5
0.8
Top temperature oC
115
130
Bottom temperature oC
330
330
As per design two types of gas oil, one light and other heavy were supposed to be withdrawn light gas oil from 6th and 18th tray and heavy gas oil from 8th/10th trays at 140-300oC and 300-350oC respectively. At present gas oil is withdrawn as 250-370oC cut from 16th/18th tray. The existing 7th to 14th double pass channel trays were replaced with valve trays in HETO,1990. Since 1970 heavy gas oil withdrawal was stopped. Main Column bottom is feed to vacuum column . VACUUM DISTILLATION :- Vacuum distillation is a method of distillation whereby the pressure above the liquid mixture to be distilled is reduced to less than its vapor pressure (usually less than atmospheric pressure) causing evaporation of the most volatile liquid(s) (those with the lowest boiling points). This distillation method works on the principle that boiling occurs when the vapor pressure of a liquid exceeds the ambient pressure. Vacuum distillation is used with or without heating the solution.
Vacuum distillation increases the relative volatility of the key components in many applications. The higher the relative volatility, the more separable are the two components; this connotes fewer stages in a distillation column in order to effect the same separation between the overhead and bottoms products. Lower pressures increase relative volatilities in most systems. A second advantage of vacuum distillation is the reduced temperature requirement at lower pressures. For many systems, the products degrade or polymerize at elevated temperatures. Vacuum distillation can improve a separation by:
Prevention of product degradation or polymer formation because of reduced pressure leading to lower tower bottoms temperatures, Reduction of product degradation or polymer formation because of reduced mean residence time especially in columns using packing rather than trays. Increasing capacity, yield, and purity.
Another advantage of vacuum distillation is the reduced capital cost, at the expense of slightly more operating cost. Utilizing vacuum distillation can reduce the height and diameter, and thus the capital cost of a distillation column.
Reduced crude from main column bottom at a temperature of approx. 330oC is pumped through Furnace. The Furnace coil outlet (4 passes) combines in one header and enter into vacuum column at 4th plates through two entry nozzle. Coil outlet temperature is maintained at about 380oC. Operating condition of Vacuum Column are as follows : Attributes
Present
As per Design
Pressure, mm of Hg abs
60
60
Top Temp.o C
60
100
Bottom Temp o C
345
385
Coil outlet temp o C
380
420
STABILIZER COLUMN:-Unstabilised gasoline is pumped to 16th/20/24th tray of Stabiliser. Feed temperature is about 110oc. The column has 35 valve trays. Operating conditions of Stabiliser are:Pressure
8.0 Kg/cm²
Top temperature
60°C
Bottom temperature
150oC
. LPG CAUSTIC WASH:LPG caustic wash facilities were provided in AVUII and was first commissioned in Sept,1984 where LPG of AVU–I, AVU-II is washed with caustic solution of 10-12%, strength. Heavy Naphtha is drawn from main column ,36th tray, through stripper. Operating parameters Top Temp.
90 Deg celsius
Top Pressure
0.3 kg/cm² (g)
Heavy Naphtha withdrawal rate
8 – 9 M³/hr
Product Streams Ex AVU-I/II
Streams
Distln. (Deg C)
Flash (deg. C)
LPG
W= +2 max.
E1 Gasoline
FBP = 110 – 130
E2 Gasoline
FBP <= 165
Heavy Naptha
FBP <= 210
MTO
FBP <= 205
>35
SK
FBP <= 280
>35
Mixed Gas Oil
95% = 370 (max.)
1st CR
Pour (Deg. C)
<=3 0-30
Wide Cut
>150
Short Residue
>150
TEMPERED WATER FACILITIES: (HETO - 1990) During HETO, tempered water facility was provided in AVU-2 . Tempered water is used as cooling media in S.R. cooler instead of Pressurised cooling water as is being used in conventional coolers. This facility is common for both AVU-1 & AVU-2. Tempered water is steam condensate, received from condensate recovery system of refinery, with neutral ph value after chemical treatment. Use of tempered water in cooler prevents the sealing and corrosion in cooler tubes thus ensuring the very efficient cooling of product (S.R.) and minimising to a great extent the maint. of the cooler. Tempered water facility is essentially a closed circulating system in which the loss of tempered water during circulation is very negligible. CORROSION CONTROL:- Ammonia is injected in the form of aquous solution for preventing HCL corrosion in pretopping and main column overheads. Recent modification of this system is the installation of on line pH meters for measuring pH in both the units. AHURALAN INJECTION:- Ahuralan is the trade name of an organic inhibitor compound, used for preventing corrosion of condenser shell. It prevents corrosion by forming a thin protective layer on the equipment. A 5% W/V and 2% W/V solutions are prepared in AVU-I and AVU-II respectively. Injection rate in both the units is 5 PPM of overhead contents.
MAJOR EQUIPMENT
TUBULAR FURNACES Tubular furnace is cylindrical type for pretopping and vacuum sections. It is box type for the main distillation column. The furnaces have sections called "Radiation Section" and convection section. A part of the tube in convection zone is for super-heating steam( used in the process) and the rest is used for heating the oil in tubes. The inside walls of the furnace are protected against the temperature effects by a refractory insulation to reduce the outside heat losses. The bottom bed show openings in which burners are placed. The flue gases go out of the furnace thorough the stack.
The stack is protected inside, in its lower part where the flue gases are still very hot by a wall of refractory bricks. A damper is located at its base to allow the regulation of the draft. This damper is built with steel suitable for the flues gases temperature. BURNERS The burner is conceived to burn either gas or oil. Gas burners are of two types: either with pre-mixing or without premixing. In the first type a part of the combustion air is mixed with the fuel gas before this has reached the injector nozzle of the burner. The burners without premixing give a diffusion flame, the combustion air entering the furnace in a parallel direction with the gas jet and slowly diffusing in it. AVUs gas burners are of this type. They give a longer and more luminous flame than those with premixing. AVUs burners are of inside-mix type. In these, the steam and oil are mixed in a chamber within the burners, and they issue together from the burner as a single stream. Foam formed in the mixing chamber is directed by the shape and direction of the burner tip so that the flame is of proper shape and size for the furnace box.The burners with spraying by steam have a flexibility much higher than those with mechanical spraying.
ATMOSPHERIC AND VACUUM DISTILLATION UNIT III
THE PROCESS
Crude preheat Crude is pumped to desalter through two parallel passes in PreDesalter Heat Exchanger Train-1. The first pass consists of four nos. of heat exchangers: The second pass of heat exchangers also has four nos. of heat exchangers: Both the passes combine in a single header and enter the desalter.
Desalter circuit:-A static Mix valve and a control valve is provided for mixing water and demulsifier with crude prior to entry into the desalter. The desalter pressure is controlled at around 9.0 Kg/cm2 (g) The desalted crude is pumped to Pretopping Column . Heat exchanger train ii:-The discharge is through a series of heat exchangers (6 Nos.). In this network of heat exchangers, crude is heated by outgoing products to a temperature of around 230 °C. Pretopping column:- The desalted crude at 230ºC enters the columns for withdrawal of unstabilised gasoline and heavy naptha. Heat exchanger train iii:- The bottom product at a temperature of around 250 °C is pumped to furnace through heat exchangers and then after combining is routed in parallel streams through pre-heat exchangers (3 Nos). The preheat temperature at the exit is around 270 °C. Part of the bottom product coming out of the heat exchanger train is sent to Pretopping Column as heat input. The coil outlet temperature of the furnace is maintained at around 360 °C. Pre topping column bottom product is sent through the main furnace. The coil outlet temperature is maintained at around 360°C. Main fractionator:-The main column is provided with: Valve trays in the top section, KERO / LGO section, Structured packing in LGO / HGO section, The bottom stripping section and Over-flash section. The column bottoms (RCO) is flashed into the Vacuum Column.
Stabiliser section:-Part of the condensed overhead gasoline is pumped through heat exchanger to stabiliser section.
Lpg caustics wash:-LPG goes to LPG caustic wash-vessel after mixing with caustic. Vacuum section The Vacuum column is provided with structured packing in LVGO pumparound section, LVGO/HVGO fractionation section, HVGO pumparound section, and Wash section and valve trays in the bottoms stripping section. The column is operated at a top pressure of 70 mm Hg. K-301 top is provided with a demister to minimize the entertainment of liquid droplets in the vapour going to overhead-condenser. The side streams of main vacuum column are as under : Stream First Second Third Bottom
Product LVGO & CR HVGO & CR SLOP & OVERFLASH SHORT RESIDUE
This Reboiler Furnace is a vertical cylindrical heater with convection and radiant section. The heater houses 12 nos. of horizontal tubes in convection section and 48 nos. bare tubes 6" NB of A335 P9 material. These 48 tubes are arranged in double pass arrangement giving material total radiant heat transfer area of 248.8 m2. The firing of this heater is done by 4 nos. combination fuel fired forced draft burners provided with pilot burners having automatic electric ignition system. Refractory material used in the radiant sanction of this heater is ceramic fiber blanket. Crude heater is a vertical cylindrical heater with convection and radiant sections. The radiant section of the heater houses 88 nos. bare tubes of 6" NB of A335 P9 material. These 88 tubes are arranged in four-pass arrangement giving total heat transfer area of 856.8 M2. In the horizontal convection section there are 24 nos. bare tubes of A335 P9 material and 64 nos. of studded tubes with an extended surface area of 950 M2. In the convection section, there are also 12 nos. of extended surface tubes for steam superheat with an extended surface area of 70 M2. The firing of this heater is done by 8 nos. combined fuel fired forced draft burners provided with pilot burners having automatic electric ignition system. Refraction material used in the radiant section of this heater is ceramic fiber blanket
Vacuum heater F-301 is a vertical cylindrical heater with convection and radiant section. In each pass of the furnace, there is arrangement for introducing turbulising steam at convection section inlet and convection section outlet. The radiant section of the heater houses 54 nos. bare tube of 6" each NB A335 P9 material. These tubes are arranged in two parallel passes giving total heat transfer area of 330 M2. In the convection section, there are 12 nos. of bare tubes of A335 P9 material of total surface area of 34.81 M2 & 44 nos. of studded tubes of A335 P9 material of total exposed surface area of 509 M2. The firing of this heater is done by 4 nos. of combined fuel fired forced draft burners provided at the floor with pilot burners having automatic electric ignition system. Refractory material used in the radiant section is ceramic fiber blanket. Air preheater :-During normal operation, combustion air for all furnaces is supplied by forced draft fans. Air is preheated at 236oC in a common air pre-heater. Air preheating is based on heat exchange between hot flue gas and combustion air. Hot flue gas leaving the convection section of the furnaces at 323oC is mixed together before going to shell side of the APH (annular spaces between the finned modules). The cast iron HT/HTA tubes have integral fins on the inside (air) and outside (flue gas) surfaces. Air preheater is provided with glass tubes in the lowest pass in order to avoid corrosion due to acid condensation in cold flue gases.
LPG RECOVERY UNIT ( LRU ) CAPACITY / STREAM FACTOR i)Gas and unstabilised naphtha from ACU
99000 Tonnes / Year
ii)Gas from existing Coker Unit
93696
iii)Stream factor
320 days
iv)Unit Turndown
25%
v)Year of Commissioning
1986
,,
,,
INTRODUCTION Gases from Coker-A, Coker-B & Stabiliser off-gas from AVU-I / II / III are compressed in a two stage steam turbine driven compressor. Compressed gases at a pressure of 14.0Kg/cm²g along with unstabilised naphtha are cooled to 40°C in air and water cooler successively and fed to a discharge knock-out pot where gas and condensate (mainly-LPG) are separated. Gases from the knock-out pot are passed through an absorber column and flow counter to the naphtha and kerosene streams in two separate sections respectively. Naphtha absorbs any C3,C4 fractions present in the gas. Kerosene further minimises the loss of naphtha entrained by the gases. Kerosene is taken from cikers and rich kerosene from this absorber is fed back to the fractionating column of cokers. Rich naphtha from the lower zone of the absorber along with the condensate obtained from the compressor discharge knock-out drum is preheated by Debutaniser bottom stream and pumped to stripper column where light ends ( C1 and C2 ) are stripped off by reboiler vapour and fed back to the inlet of compressor discharge KO drum ( to recover C3 ,C4 , if any). Fuel gas from the absorber top goes to a knock-out drum and fed to the refinery gas network. Stripper bottom containing mainly LPG and Naphtha
are fed to the Debutaniser column for separation of LPG and Naphtha. LPG is withdrawn from the top reflux drum and stabilised naphtha from the bottom of the debutaniser column. A required part of this stabilised naphtha is recycled back to the absorber as absorbing medium and rest of stabilised naphtha goes as product. Both LPG and stabilised naphtha products are further washed in caustic soda wash section separately for removal of any H2S. Products are further passed through sand filters and then sent to the product storage tanks FEED - STOCK a)
Gas from ACU
COMPONENT
WT. %
Methane
26.28
Ethane
15.28
Ethylene
4.17
Propylene
9.29
i-Butane
24.00
n-Butane
7.68
C is-Trans Butane
8.56
C5
11.12
c)
Gases from existing Coker COMPONENT
WT. %
Methane
25.7
Ethane/Ethylene
18.9
Propane
17.7
Propylene
8.6
i-Butane
1.
b) Unstabilised naphtha from ACU Sp. Gravity at 15°C Vap. Pressure
- 0.710 22.0 Kg/cm²g
PRODUCT CHARACTERISTICS LPG Copper strip corrosion for 1 hr at 38°C
1b Max.
Dryness enrained water
No free
H2S
Absent.
Odor (Min)
Level 2
Total Volatile Solution (Max.)
0.02
W.T (95% Vaporised at 760 mm Hg (Max)
+ 2°C
Vapour pressure at 65°C Kg/cm²g
15 Max.
STABILISED NAPHTHA TBP
C5 + 140°C
15°C
0.7109
HYDROCARBON % WT - Saturates
64.5
- Olefins
23.5
- Aromatics
12.0
Mercaptans % wt.
0.016
Cu Strip corrosion But (3hrs. at 50°C)
2a
RVP psig (max.)
10.0
FUEL GAS
WT. %
Methane
45.23
Ethylene
7.20
Ethane
25.50
Propylene
7.60
Propane
11.70
C5 +
2.77
BATTERY LIMIT CONDITIONS FEED / PRODUCTS
Kg/cm²g
0°C
DESTINATION / SOURCE
Gas
2.2
45
ACU/existing
Unstabilised
16.5
45
ACU
Fuel Gas
2.0
50
Refinery FG system
LPG
20.3
40
LPG Storage
Stabilised Naphtha
5.0
40
Naphtha/MS Tanks.
Naphtha
MATERIAL BALANCE S.No.
COMPONENT
1.
Gas
2.
LPG
3.
Naphtha
FEED
PRODUCT
T/YR
T/YR.
192696
DESTINATION
92660
Existing FG System
52016
LPG Storage
57220
Naphtha/MS Tank.
UTILITIES CONSUMPTION UTILITY
PRESSURE Kg/cm²g
TEMP. 0°C
CONSUMPTION
High Pressure Steam(SH)
34
415
25000 Kg/hr
Medium pressure Steam
19.0
211
9594 Kg/hr
Low pressure steam (SL)
2.5
290
Intermittent Lost Station.
** Cooling water (WC)
2.5
693 M³/hr
Inst. Air
3.5
40
60 NM³/hr
Plant Air
6.5
40
340 NM³/hr. *
Fresh water
7.0
Ambient
6 *
* Intermittent. * If ACU is down consumption shall be 347M³/hr CHEMICALS Caustic soda Corrosion inhibitor
12.0 tones/Year. 0.75 tones/Year.
C CRU(C CATA ALYTI IC RE EFOR RMIN NG UN NIT) ytic reforming g is a ch hemicall processs used d to con nvert pe etroleum m Cataly refine ery naphthas, typicallly havin ng low octane o rratings,, into hiigh-octane liqu uid produ ucts called reformate es which h are co ompone ents of high-oc ctane ga asoline (also known k a as petro ol). Bassically, the t process re e-arrang ges or rre-stru uctures the hy ydrocarrbon mo oleculess in the e naphth ha feed dstocks as well as bre eaking some of the molecu ules into o smalle er molec cules. T The ove erall eff fect is that t th he uct refo ormate contain ns hydrocarbons with h more complex c x molec cular produ shape es havin ng highe er octan ne value es than the hy ydrocarbons in the naphtha feedsstock. I In so do oing, the e proce ess sepa arates h hydroge en atom ms from m the hydro ocarbon n molecu ules and d produ uces verry significant amount a s of by yproduct hydro ogen gass for usse in a n numberr of the e other processses invo olved in na moderrn petroleum refinery r y. Othe er bypro oducts are sma all amou unts of f metha ane, eth hane, prropane and buttanes.
The reaction r n chem mistry All th he reacttions oc ccur in tthe pre esence of o a cattalyst and a a high parttial pressure of hydrogen. Dep pending upon the type e or verrsion of f catalytic reforming ussed as well w as tthe dessired re eaction severitty, the reactio on condittions ra ange fro om tem mperatures of about a 4 495 to 525 5 °C and fro om pressures of f about 5 to 45 5 atm. The common c ly used catalyttic refo orming catalyssts contain nob ble mettals suc ch as pla atinum a and/or rhenium m, which h are ve ery sussceptible to po oisoning g by sulfurr and nitrogen compou unds. Therefo T re, the naphth ha feed dstock to t a cataly ytic ref former is alway ys pre-processsed in a hydrodesulfu urizatio on unit which h removes both h the su ulfur an nd the nitrogen n n compo ounds. The four f majjor cata alytic rreforming reac ctions a are:1: The e dehyd drogena ation of f naphtthenes to convvert the em into aromattics as exemp plified in the convers c sion metthylcyc clohexan ne (a na aphthen ne) to toluene: t :-
2: The e isome erizatio on of no ormal paraffin p ns to issoparaf ffins ass exemp plified in i the co onversion of normal o octane to t 2,5-D Dimethy ylhexan ne (an issoparaf ffin), ass shown n below:
3: The e dehyd drogena ation and arom matizattion of paraff fins to a aromatics (comm monly ca alled de ehydroc cyclizattion) as exemp plified in n the co onversion of norma al hepta ane to toluene t , as sho own belo ow:
4: The e hydro ocrackin ng of p paraffin ns into smallerr molecules as exemplified by b the crracking of normal hep ptane in nto isop pentane e and etthane, a as shown below w:
The hydrocr h racking of para affins iss the on nly one of the above f four ma ajor reforming re eactionss that c consume es hydrrogen. T The isom merizattion of normal paraffins does not consum c me or prroduce hydroge h en. How wever, b both the dehyd drogena ation of f naphth henes and a the dehydrrocycliz zation o of paraf ffins produ uce hydrrogen. The T ove erall ne et produ uction o of hydro ogen in the catalytic reforming of f petrolleum na aphthass rangess from a about 50 5 to 20 00 cubiic mete ers of hyd drogen gas (att 0 °C and 1 atm m) per cubic m meter of f liquid naphth ha feedsstock.
Typical naph htha fe eedstoc cks o ad liquid d distilla ate from atmo ospheric c distilllation c column is i called d The overhea naphttha and will bec come a major compon nent of the ref finery'ss gasoliine (petro ol) prod duct aftter it iss furthe er proc cessed tthrough h a cata alytic ic hydro odesulfu urizer to t remo ove sulf fur-conttaining hydroc carbons and a catalyt c reformer to reform m its hy ydrocarb bon molecules into mo ore com mplex molecule m es with a higherr octane rating g value.. The na aphtha is a mix xture o of very many differrent hy ydrocarbon com mpounds. It ha as an initial boiling po oint of about a 35 °C and d a fina al boiling g point of about 200 °C, and d it conttains pa araffin,
naphthene (cyclic paraffins) and aromatic hydrocarbons ranging from those containing 4 carbon atoms to those containing about 10 or 11 carbon atoms. "light" naphtha containing most (but not all) of the hydrocarbons with 6 or less carbon atoms and a "heavy" naphtha containing most (but not all) of the hydrocarbons with more than 6 carbon atoms. The heavy naphtha has an initial boiling point of about 140 to 150 °C and a final boiling point of about 190 to 205 °C. The naphthas derived from the distillation of crude oils are referred to as "straight-run" naphthas. It is the straight-run heavy naphtha that is usually processed in a catalytic reformer because the light naphtha has molecules with 6 or less carbon atoms which, when reformed, tend to crack into butane and lower molecular weight hydrocarbons which are not useful as high-octane gasoline blending components. Also, the molecules with 6 carbon atoms tend to form aromatics which is undesirable because governmental environmental regulations in a number of countries limit the amount of aromatics (most particularly benzene) that gasoline may contain. Key specifications of Petrol : BS-II
BS-III/ Euro-III equivalent
Euro-IV
Regular
Premium
Regular
Premium
Sulphur, (max)
ppm
w
500
150
150
50
50
Benzene, (max)
Vol
%
3
1
1
1
1
88
91
95
91
95
No Spec.
42
42
35
35
No Spec.
21
18
21
18
RON (min) Aromatics, (max)
Vol
%
Olifins, Vol % (max)
TEMP PERAT TURE & PRES SSURE COND DITIO N : SL
EQUIPMENT E (KG/C CM2G)
TEM MPERAT TURE (0C)
NO.
EO OR
S SOR
EOR
PRESSURE SOR R
1.
FEED B F BEFORE E 0 02-EE-0 001
6 65.0 (T TI-1112))65.0
20.9 9
20.9 9
2.
F FURNA ACE I/L L
28 85.3 (T TI-1114))325.0
17.6
17.6 6
3.
R REACTO OR I/L L
33 30.0 (T TI-1122))370.0
16.5 5 (PI-110 02) 16.5 5
4.
R REACTO OR O/L L
33 30.0
15.0 0 (PI-110 03) 15.0 0
5.
HYDRO H O-TREA ATER S SEPARA ATOR 4 45.0 (T TI-1203)45.0
6.
HYDRO H O-TREA ATER PURGE
7.
STRIPP S PER FEE ED A AFTER EXCHA ANGER186.0 (TI-13111)186.0 0 1 14.5
8.
STRIPP S PER OV VER-HEA AD136.0 1 127.0 14.4 (PI I-1303))14.4
9. -
S STRIPP PER BOTTOM2 225.0 (TI-1316 6)225.0 0 -------
10.
R REFLUX X
37 70.0
5 58.5 (T TI-1313))57.3
13.0 0
14.5 5
13.0 0
(P PI-1302 2) (T TI-1312 2) -----
18.2 (PI-1304 18.2 2 & PI-13 305)
BRIEF PROCESS DESCRIPTION: (I)NAPHTHA SPLITTER UNIT (NSU) : IBP-140 0C cut naphtha from storage (TK 250, 251, 252) is fed to splitter column 01-CC-001 under flow control by off site pump 41-PA-1A/B at tray No. 14. The feed is heated up to 95 0C in splitter feed/bottom exchanger 01-EE-001 A/B against splitter bottom stream before it enters the column. The over head vapours are totally condensed in air condensers 01-EA-001. The liquid collected is pumped by splitter reflux pump 01-PA-001 A/B and one part sent as top reflux back to the column under flow control 02-FC1102 to maintain the top temperature. The balance, which constitutes the IBP-70 0C cut naphtha is sent to storage under reflux drum level control 01-LC-1101 after cooling in a water cooler 01-EE-002. Reflux drum boot water is drained in OWS manually. The pressure of splitter is controlled at reflux drum by passing a part of hot column overhead vapours around the condenser or releasing the reflux vapours to flare through a split range controller (01-PC-1101). The splitter bottom product which constitutes 70-140 0C cut naphtha is pumped to spliter feed/bottom exchanger 01-EE-001 A/B by hydro treater feed pumps 01-PA-003 A/B. The bottom product after exchanging heat with feed is split into two streams. One is fed to the hydro treater unit at a temp. of 65 0C and the other is sent to storage under column level control 01-LC-1102 after being colled in splitter bottom column 01-EE-003. The heat necessary for splitter reboiling is supplied by splitter reboiler furnace 01-FF-001 and desired temperature maintained by controlling the fuel firing. The circulation through reboiler is provided by splitter reboiler pumps 01-PA-002 A/B. 01-FF-001 is double pass vertical
cylinderical furnace having four burners fired from the bottom. It has soot blowing facility for convection section.
(II)HYDROTREATER UNIT (HTU) : REACTION AND SEPARATION SECTION : The naphtha from NSU is fed to HTU by a pump 01-PA-003 A/B. The feed flow is controlled by flow control valve 02-FC-1101. The feed then mixed with Rich Hydrogen Gas from HP separator of reformer. The Rich Hydrogen gas flow is controlled by 02-FC-1202. Both the liquid naphtha and rich hydrogen gas are pre-heated in a series of exchangers 02-EE001 A/B/C/D/E/F which are feed/reactor effluent heat exchangers. Then mixture is heated upto reaction temperature in a furnace 02-FF-001 and fed to the reactor 02-RB-001. The furnace 02-FF-001 is four pass having three burners fired from bottom. The furnace is having facility of soot blowing. The reactor inlet temperature is maintained by 02-TC-1101 cascaded with either fuel oil or fuel gas PC's. The furnace is provided with all safety shut down inter locks. It has also provision of decoking. The desulfurisation and hydro treating reaction takes place in 02-RB-001 at almost constant temperature since heat of reaction is quite negligible. The reactor is provided with facility of steam and air for regeneration of catalyst. The catalyst for reactor is HR-306. The reactor catalyst bed has been provided with five number of thermo couple points at various location to get the bed temperature during regeneration of the catalyst. The reactor effluent after having heat exchanged in 02-EE-001 series with feed goes to air cooler 02-EA-001. The air cooler fans pitch is variable i.e. cooling load can be adjusted as per situation requirement. After air cooler the effluent is cooled in a trim cooler 02-EE-002. The product is collected in a separator vessel 02-VV-001. Sour water is
drained from the separator drum boot manually. The separator drum pressure is maintained by 02-PC-1201, releasing the excess gas in FG system. In event of emergency the separator excess pressure can be released to flare through an on-off c/v HV-1201. A line has been provided to feed the naphtha to stripper, during start up, bypassing the reaction/separation section. STRIPPER SECTION : The separator liquid is pumped by 02-PA-001A/B under flow control 02FC-1201 cascaded with 02-LC-1201 to stripper feed/bottom exchanger 02-EE-003 A/B/C when it gets heat exchanged by hot stripper bottom stream. The stripper column consists of 28 Nos. of valve trays one to eight number of trays are single pass and the rest double pass. Feed coming from 02-EE-003 A/B/C enters at 9th tray from two sides. The over head vapours are cooled down in 02-EA-002 air condenser and collected in 02VV-002 stripper refflux drum. The fan load can be adjusted. The condensed hydro-carbons are returned to column top by pump 02-PA002A/B under flow control 02-FC-1301 cascaded with 02-LC-1302 as reflux to maintain the top temp. The water accumulated in the boot is sent for disposal as sour water. The reflux drum pressure is maintained by 02-PC-1301 releasing excess gas in the FG system. The facility is there to inject corrosion inhibitor by pump 02-PA-005A.
Stripper bottom product exchanged heat with stripper feed in 02-EE003A/B/C and then sent to reformer as hot feed. The excess or required hydro-treated naphtha is sent to storage after being cooled in 02-EE-004 A/B under level control 02-LC-1301.
The necessary heat for stripper reboiling is supplied by 02-FF-002 reboiler heater. 02-CC-001 product is circulated through 02-FF-002 single pass cylinders vertical furnace by 02-PA-003 A/B. Partial
vap pourizattion occ curs in 02-FF-002. Reboilin R ng is controlled d by 02 2-TC-13 301 at 3rd pla ate from the bottom m of 02--CC-001. Furnace is p provide ed with all saf fety intter lock ks.
Hydro tre eated naphtha n a from hydro treate er unit is pum mped to o requirred pre essure b by
03-P PA-001 A/B un nder flo ow conttrol 03--FC-1101 A/B and a
mix xed witth recy ycle gass from m the recycle gas co ompresssor (03 3-KA-00 01). The mixed d feed is pre heated d in the feed-e effluent excha anger 03-EE-0 0 001 followed by fee ed/efflu uent ex xchange er 03-E EE-002 2. Then n the mixture m e is bro ought u upto the reacttion tem mperature (48 80 0C) by hea ating in n the prehea ater 03 3-FF-00 01 and tthen fed to 1stt reactor 03-R RB-001.. As the re eaction n is end do-therrmic, th he tem mperature drops, so the first rea actor ef ffluentt is heatted in the t firsst inter heaterr 03-FF F-002 prior p to be sen nt to th he secon nd reac ctor 03-RB-00 02. In the sam me way 03-RB--002 ef ffluent is heatted in the seco ond inte er heater 03-FF-003 prior to be f fed to the t thirrd reac ctor 03--RB-003. The efflu uent fro om the last re eactor 03-RB--003 is split into two o strea ams and d send for hea at recovery pa arallely to feed/efflu uent ex xchange er (03-E EE002) and d stablizer re eboiler (03-EE-003)). The outlet from the two t exc changerrs is co ombined d by a these wa ay valve e 03-TI IC-11011 and th hen cooled dow wn suc ccesivelly in tthe Ze eemann Secatthen ex xchanger (03 3-EE-00 01),
reformer effluent cooler 03-EA-001 and effluent trim cooler 03-EE-004. The cooled reactor effluent is flashed in the reformer separator 03-VV001. Vapour and liquid phase are separated in separator 03-VV-001. Part of the gas phase constitutes the hydrogen recycle gas to the reactor circulated by recycle gas compressor 03-KA-001. Remaining amount, corresponding to the amount of gas produced, is compressed by the hydrogen rich gas compressor 03-KA-002 A/B. The pressure control in separator is achieved by a kick back gas flow from HP Absorber (03-VV003) to separator. Should the gas be produced in excess to 03-KA-002 A/B capacity, degassing in split range to fuel gas is performed, through 03-PC-1401 A and 03-PC-1402. The separator liquid is sent by reformer separator bottom pumps (03-PA002 A/B) under level control 03-LC-1401 for recontacting with the gas compressed by 03-KA-002 A/B. The hot flue gases from all the three reformer furnaces are combined and sent to stream generation system forwaste heat recovery to produce MP steam. Provision is there to dry the recycle gas into a dryer (03-RB004). The dryer can later be regenerated. The unit has also been provided with facilities for continuous chloriding, water injection, DMDS/Ccl4 injection and caustic soda circulation. The separator (03-VV-001) vapour after passing through KO drum (03VV-002) is compressed in the H2 Rich Gas Compressor (03-KA-002 A/B) and recontacted with separator liquid. The recontacted vapour and liquid is cooled in a cooler (03-EE-005) and then fed to HP absorber (03-VV003). The aim of this device is to allow for high recovery of the C5 contained in the gas phase of separator and improve the quality (H2 concentration) of the produced gas.
A part of hydrogen rich vapour goes to HTU as a make up hydrogen and balance goes to the fuel gas system under pressure control 03-PC-1601. The liquid from the 03-VV-003 is drawn off under level control 03-LC1601 and mixed with stabilizer vapour distillate. The combined stream is cooled in LPG absorber feed cooler 03-EE-006 and flashed in LPG absorber. Off-gas is sent under pressure control to fuel gas system. The liquid from 03-VV-004 is pumped by stabilizer feed pumps 03-PA-003 A/B. After pre heating in stabilizer feed/bottom exchanger 03-EE-007 the mixture is fed to the stabilizer 03-CC-001 at tray No. 13. Stabilizer over head vapours are partialy condensed in stablizer condenser 03-EE-008 and flashed in stabilizer reflux drum 03-VV-005. The vapour phase is sent to LPG absorber for C3 and C4 recovery. A part of condensed liquid is pumped as reflux to the column by stabilizer reflux pump 03-PA-004 A/B under the flow control and the balance is sent to LPG Recovery Unit under level control of reflux drum. The heat of reboiling to the stabilizer is provided by the hot reactor effluent
in
the
stabilizer
reboiler
03-EE-003
and
the
desired
temperature maintained by controlling the flow of reactor effluent by the three way valve. The bottom product, stabilized reformate, is cooled in the feed/bottom exchanger 03-EE-007followed by reformate cooler 03-EA-002 and reformate trim cooler 03-EE-009 before being routed to storage Tk 77 to 84.
FLUIDISED CATALYTIC CRACKING Fluid catalytic cracking (FCC) is the most important conversion process used in petroleum refineries. It is widely used to convert the high-boiling, high-molecular weight hydrocarbon fractions of petroleum crude oils to more valuable gasoline, olefinic gases and other products. Cracking of petroleum hydrocarbons was originally done by thermal cracking which has been almost completely replaced by catalytic cracking because it produces more gasoline with a higher octane rating. It also produces byproduct gases that are more olefinic, and hence more valuable, than those produced by thermal cracking.
The feedstock to an FCC is usually that portion of the crude oil that has an initial boiling point of 340 °C or higher at atmospheric pressure and an average molecular weight ranging from about 200 to 600 or higher. This portion of crude oil is often referred to as heavy gas oil. The FCC process vaporizes and breaks the long-chain molecules of the high-boiling hydrocarbon liquids into much shorter molecules by contacting the feedstock, at high temperature and moderate pressure, with a fluidized powdered catalyst.
In effect, refineries use fluid catalytic cracking to correct the imbalance between the market demand for gasoline and the excess of heavy, high boiling range products resulting from the distillation of crude oil.
The first commercial use of catalytic cracking occurred in 1915 when Almer M. McAfee of the Gulf Refining Company developed a batch process using aluminum chloride (a Friedel Crafts catalyst known since 1877) to catalytically crack heavy petroleum oils. However, the prohibitive cost of the catalyst prevented the widespread use of McAfee's process at that time.
The mode ern FCC units a are all continuo c ous processes which o operate e 24 hou urs a da ay for as a much h as 2 to 3 yea ars betw ween sh hutdown ns for routine r ma aintenan nce. There are e a numb ber of differe ent prop prietary y designs thatt have been b devveloped d for mo odern F FCC unitts. Each h design n is ava ailable u under a license e tha at mustt be purrchased d from the t dessign devveloper by any petrole eum ref fining company y desiring to co onstruc ct and o operate e an FCC C of a given g dessign. Bassically, there are a two o differrent con nfigurattions fo or an FC CC unit: the "sttacked" type where w th he reac ctor and d the ca atalyst regene erator are a con ntained in a sin ngle vesssel witth the reactor r above the cattalyst reg generattor and the "siide-by-sside" ty ype whe ere the e reacto or and catalyst c t reg generattor are in two separatte vessels. The ese are e the ma ajor FC CC dessigners and lic censors: Sid de-by-sside connfigurattion:
CB&I Lummus ExxonMobil Research and Engineering (EMRE) Shell Global Solutions International Stone & Webster Engineering Corporation (SWECO) / Institut Francais Petrole (IFP) Universal Oil Products (UOP) - currently fully owned subsidiary of Honeywell
Reactor and Regenerator The schematic flow diagram of a typical modern FCC unit in Figure 1 below is based upon the "side-by-side" configuration. The preheated high-boiling petroleum feedstock (at about 315 to 430 °C) consisting of long-chain hydrocarbon molecules is combined with recycle slurry oil from the bottom of the distillation column and injected into the catalyst riser where it is vaporized and cracked into smaller molecules of vapor by contact and mixing with the very hot powdered catalyst from the regenerator. All of the cracking reactions take place in the catalyst riser. The hydrocarbon vapors "fluidize" the powdered catalyst and the mixture of hydrocarbon vapors and catalyst flows upward to enter the reactor at a temperature of about 535 °C and a pressure of about 1.72 barg. The reactor is in fact merely a vessel in which the cracked product vapors are: (a) separated from the so-called spent catalyst by flowing through a set of two-stage cyclones within the reactor and (b) the spent catalyst flows downward through a steam stripping section to remove any hydrocarbon vapors before the spent catalyst returns to the catalyst regenerator. The flow of spent catalyst to the regenerator is regulated by a slide valve in the spent catalyst line. Since the cracking reactions produce some carbonaceous material (referred to as coke) that deposits on the catalyst and very quickly reduces the catalyst reactivity, the catalyst is regenerated by burning off the deposited coke with air blown into the regenerator. The regenerator operates at a temperature of about 715 °C and a pressure of about 2.41 barg. The combustion of the coke is exothermic and it produces a large amount of heat that is partially absorbed by the regenerated catalyst and provides the heat required for the vaporization of the feedstock and the endothermic cracking reactions that take place in the catalyst riser. For that reason, FCC units are often referred to as being heat balanced. The hot catalyst (at about 715 °C) leaving the regenerator flows into a catalyst withdrawal well where any entrained combustion flue gases are allowed to escape and flow back into the upper part to the regenerator. The flow of regenerated catalyst to the feedstock injection point below the catalyst riser is regulated by a slide valve in the regenerated catalyst
line. The hot flue gas exits the regenerator after passing through multiple sets of two-stage cylones that remove entrained catalyst from the flue gas, The amount of catalyst circulating between the regenerator and the reactor amounts to about 5 kg per kg of feedstock which is equivalent to about 4.66 kg per litre of feedstock. Thus, an FCC unit processing 75,000 barrels/day (12,000,000 litres/day) will circulate about 55,900 metric tons per day of catalyst. Distillation column The reaction product vapors (at 535 °C and a pressure of 1.72 barg) flow from the top of the reactor to the bottom section of the distillation column (commonly referred to as the main fractionator) where they are distilled into the FCC end products of cracked naphtha, fuel oil and offgas. After further processing for removal of sulfur compounds, the cracked naphtha becomes a high-octane component of the refinery's blended gasolines. The main fractionator offgas is sent to what is called a gas recovery unit where it is separated into butanes and butylenes, propane and propylene, and lower molecular weight gases (hydrogen, methane, ethylene and ethane). Some FCC gas recovery units may also separate out some of the ethane and ethylene. Although the schematic flow diagram above depicts the main fractionator as having only one sidecut stripper and one fuel oil product, many FCC main fractionators have two sidecut strippers and produce a light fuel oil and a heavy fuel oil. Likewise, many FCC main fractionators produce a light cracked naphtha and a heavy cracked naphtha. The terminology light and heavy in this context refers to the product boiling ranges, with light products having a lower boiling range than heavy products. The bottom product oil from the main fractionator contains residual catalyst particles which were not completely removed by the cyclones in the top of the reactor. For that reason, the bottom product oil is referred to as a slurry oil. Part of that slurry oil is recycled back into the main fractionator above the entry point of the hot reaction product vapors so as to cool and partially condense the reaction product vapors as they enter the main fractionator. The remainder of the slurry oil is pumped through a slurry settler. The bottom oil from the slurry settler contains most of the slurry oil catalyst particles and is recycled back into the catalyst riser by combining it with the FCC feedstock oil. The socalled clarified slurry oil or decant oil is withdrawn from the top of slurry settler for use elsewhere in the refinery or as a heavy fuel oil blending component.
Regenerator flue gas Depending on the choice of FCC design, the combustion in the regenerator of the coke on the spent catalyst may or may not be complete combustion to carbon dioxide (CO2). The combustion air flow is controlled so as to provide the desired ratio of carbon monoxide (CO) to carbon dioxide for each specific FCC design.[1][4] In the design shown in Figure 1, the coke has only been partially combusted to CO2. The combustion flue gas (containing CO and CO2) at 715 °C and at a pressure of 2.41 barg is routed through a secondary catalyst separator containing swirl tubes designed to remove 70 to 90 percent of the particulates in the flue gas leaving the regenerator.[8] This is required to prevent erosion damage to the blades in the turboexpander that the flue gas is next routed through. The expansion of flue gas through a turbo-expander provides sufficient power to drive the regenerator's combustion air compressor. The electrical motor-generator can consume or produce electrical power. If the expansion of the flue gas does not provide enough power to drive the air compressor, the electric motor/generator provides the needed additional power. If the flue gas expansion provides more power than needed to drive the air compressor, than the electric motor/generator converts the excess power into electric power and exports it to the refinery's electrical system.[3] The expanded flue gas is then routed through a steam-generating boiler (referred to as a CO boiler) where the carbon monoxide in the flue gas is burned as fuel to provide steam for use in the refinery as well as to comply with any applicable environmental regulatory limits on carbon monoxide emissions.[3] The flue gas is finally processed through an electrostatic precipitator (ESP) to remove residual particulate matter to comply with any applicable environmental regulations regarding particulate emissions. The ESP removes particulates in the size range of 2 to 20 microns from the flue gas.[3] The steam turbine in the flue gas processing system (shown in the above diagram) is used to drive the regenerator's combustion air compressor during start-ups of the FCC unit until there is sufficient combustion flue gas to take over that task.
COKER A GENERAL DATA 1. Capacity
0.6 MMTPA
2. On-stream hours per year.
7200
3. Original Technology 4. Year of Commissioning 5. Turn Down
Russian 1964 95 %
The unit is designed for the following three cases:CASE-I
Feed corresponding to future refinery configuration having Resid Desulphurisation unit, while processing 6.0 MMTPA high sulphur crude 50:50 wt Arab mix)
CASE-II
Feed corresponding to future refinery configuration without Resid Desulphurisation, while processing 6.0 MMTPA low sulphur crude (Bonny light).
CASE-III
Feed corresponding to future refinery configuration without Resid Desulphurisation, while processing 4.2 MMTPA low sulphur crude (Bonny light).
PROCESS DESCRIPTION Delayed coking’ process is an effective conversion process for upgradation of the heavy residuals from the refinery distillation unit into valuable distillates and premium quality petroleum coke. In this process the heavy residual feed stocks are heated up to coking temperature and the mixture is allowed to stand for prolong period in large insulated vessels called coke drums. During this time, the heavy stock undergoes thermal cracking at large high b.pt. H/C molecules are decomposed into smaller lower boiling point molecules and at the same time some reactive molecules undergoes pyrolytic polymerization forming fuel oil
and coke. Coke is formed by two different mechanism. In one the colloidal suspension of the asphltene and resin components is re-arranged resulting in the precipitation of the compounds to form highly-cross-linked structure of amorphous coke. The compounds are also subjected to a cleavage of this sulphatic groups. Coke formed from these resinasphatene compounds is undesirable for making premium grade coke. The other reaction mechanism involves the polymerisation and condensation of aromatics, grouping a large number of these compounds to such a degree that eventually coke is formed. The coke produced from these aeromatic compounds is the most suitable premium grade needle coke. The other products of pyrolysis are separated into distillate fuels and recovered separately in a fractionator column. Reduced crude is received in the Feed Surge Drum in Coker-A through offsite Pumps. The feed RCO is then preheated to 240°C by heat exchanger against Coker products like Coker Kero, Light Diesel Oil (LDO) Product and LDO CR. The preheated RCO is fed to fractionator column at two levels, one below and the other above the vapour inlet nozzle. This facility is provided to control fractionator bottom temperature. The feed material along with the recycle stock is pumped to the reaction coils of the coker furnace at a temperature of 340 – 360 °C and the material is heated to the temperature of 500°C which resulted in partial vapourization and mild cracking of the stock. The vapour liquid mixture then enters the coke chamber which is in coking service, where the vapour experience further cracking as it passes through the coke chamber and the liquid experience successive cracking and polymerization until it is converted to vapour and coke. The unit has two blocks and each block has two coke
chambers, one in coking service while the other is being decoked with high pressure water jets. Each block has got individual heater. The coke chamber overhead vapours enter the fractionator via a quench column at a temperature of about 425°C. In the fractionator column, gas and naphtha are obtained as overhead products and kerosene, LDO and CFO as side draw off products. Kerosene, LDO and CFO are steam stripped in the stripper columns, cooled prior to their being routed to their destinations. A LDO circulating reflux stream is drawn and is utilised for HP steam generation. LDO Product, CFO Product and FO IR are utilized for generating MP steam. A cold LDO stream is used as quench in the quench column. Besides refinery slop and gas oil from offsites tank can be used as quench in the vapour line of the coke chambers. The residue from the bottom of quench column is sent to storage after further cooling. The vapour from the fractionator overhead are cooled in air cooler and water condensers and then led to reflux drum where gas and liquid separate out. The gases from the fractionator reflux drum sent to LPG Recovery unit of the refinery. Condensed naphtha from reflux drum is also routed to the LPG recovering unit for stablilization. A part of condensed naphtha sent back to fractionator column as top reflux. Coke from the cooled drained chamber is cut and cleared by hydraulic jets operating at a pressure of about 200 Kg/cm². Coke along with water falls to the ground. The coke from the drop-out area is sent to storage using Coke Sizer and conveyors.
COKER B PROCESS DESCRIPTION
Delayed
coking’
process
is
an
effective
conversion
process
for
upgradation of the heavy residuals from the refinery distillation unit into valuable distillates and premium quality petroleum coke. In this process the heavy residual feed stocks are heated up to coking temperature and the mixture is allowed to stand for prolong period in large insulated vessels called coke drums. During this time, the heavy stock undergoes thermal cracking at large high b.pt. H/C molecules are decomposed into smaller lower boiling point molecules and at the same time some reactive molecules undergoes pyrolytic polymerization forming fuel oil and coke. Coke is formed by two different mechanism. In one the colloidal suspension of the asphltene and resin components is re-arranged resulting in the precipitation of the compounds to form highly-crosslinked structure of amorphous coke. The compounds are also subjected to a
cleavage
of
this
sulphatic
groups.
Coke
formed
from
these
resinasphatene compounds is undesirable for making premium grade coke. The
other
reaction
mechanism
involves
the
polymerisation
and
condensation of aromatics, grouping a large number of these compounds to such a degree that eventually coke is formed. The coke produced from these aeromatic compounds is the most suitable premium grade needle coke. The other products of pyrolysis are separated into distillate fuels and recovered separately in a fractionator column.
Reduced crude is received in Coker-B from offsite storage tanks by a 18” dia pipeline. The feed stock RCO from storage is preheated to 200°C by
heat exchanger against Coker products like Coker Kero, heavy gas oil (HGO), coker fuel oil (CFO) and residue. The preheated RCO is further heated to 240°C in the pre-heat section of the coker furnace and fed to fractionator column. This feed goes to the fractionator is at two levels, one below and the other above the vapour inlet nozzle. This facility is provided to control fractionator bottom temperature. The feed material along with the recycle stock is pumped to the reaction coils of the coker furnace at a temperature of 380°C. The material is heated to a temperature of 500°C which resulted in partial vapourization and mild cracking of the stock. The vapour liquid mixture then enters the coke chamber which is in coking service, where the vapour experience further cracking as it passes through the coke chamber and the liquid experience successive cracking and polymerization until it is converted to vapour and coke. The unit has two coke chambers, one in coking service while the other is being decoked with high pressure water jets. The coke chamber overhead vapours enter the fractionator via a quench column at a temperature of about 425°C. In the fractionator column, gas and naphtha are obtained as overhead products and kerosene, HGO and CFO as side draw off products. Kerosene, and HGO are steam stripped in the stripper columns, cooled prior to their being routed to their destinations.
HYDROGEN GENERATION UNIT GENERAL DATA Design Capacity Stream Factor Turn down ratio Original Technology Date of Commissioning
34 TMT H2 production 8,000 hours per year 30 % Haldor Topsoe 24 TH Apri’2002
BRIEF PROCESS DESCRIPTION To meet the make up requirement of Hydrogen for DHDT Unit, naphtha steam reforming type Hydrogen unit has been considered where Hydrogen is produced by steam reforming of Naphtha. Naphtha is
first desulphurised over a desulphurisation catalyst where, in presence of hydrogen, non-reactive sulphur compounds are hydrogenated to hydrogen sulphide which is then absorbed on Zince Oxide beds. The desulphurised feed is mixed with preheated steam and then heated to the desired temperature before entering steam reforming furnace tubes containing a nickel based catalyst. The reformed gases leave the tubes and after exchanging heat to generate steam, pass through a CO shift convertor where most of the carbon monoxide is reacted with excess steam to produce additional hydrogen and carbon dioxide. The converted gases leave the reactor and preheat the incoming Naphtha, Boiler Feed water and Demineralised water. The impurities like carbon monoxide, carbon dioxide, methane, nitrogen and water vapour are removed by high pressure adsorption on molecular sieves. Activated carbon and alumina gel in PSA (Pressure Swing Adsorption) system. All adsorbed gases are removed during deadsorption & regeneration of the beds and used as fuel in reformer furnace, and Hydrogen with 99.5% (vol) purity is fed to bullet / DHDT unit. Feed Stock
DESIGN
CASE 1: 30% (WT%) MIXTURE OF RFCCU OFF GAS AND 70% (WT%) SRN CASE 2: 100% CAPACITY ON SRN Type of Feed Feed composition
GAS / LIQUID For Liq. Feed
For gas feed
(SRN –C5-90OC
(RFFCU OFF
CUT)
GASES)
ATACHED AS ANNEXURE-1
Gas MW
Feed Characteristics
-
For Liq. Feed
For gas feed
(SRN –C5-90OC
(RFFCU OFF
CUT)
GASES)
Liq. Sp. Gravity @ 150C
0.692
Feed composition /TBP
ATACHED AS
ATACHED AS
ANNEXURE-1
ANNEXURE-1
C5-90OC
-
0.025 / NIL
-
ATACHED AS
-
Cut range, deg.C Total Sulfur / Nitrogen, WT% PONA, Vol%
ANNEXURE-1
Distillation ASTM D86, OC
IBP
: 41
10 VOL%
: 44
50 VOL%
: 59
70 VOL%
: 71
FBP
: 98
-
Calorific Value, Kcal/Kg
10,492
-
C/H Ratio, wt / wt
5.7
-
Any other additional feed to PSA unit Feed Type
-
Flow rate (Kg/hr)
-
Fuel Type Liquid Naphtha Purge gas from PSA Syn. Gas (in
Flow rate (Kg/hr) -
case PSA shut down)
Internal fuel to Furnace (off Gas from PSA), 3
NM /Hr
CASE-1 CASE-2
PSA –I OFF GAS : 25845
25250
Steam/carbon ratio for the feed for Steam
-
Reforming Blow down 2nd Demister vent quantities Kg/hr
-
Inlet temperature to the tabular Reformer, deg,
640
C Reformer exit temperature Deg, C
930
Pressure around the reformer inlet of reformer, Kg/cm2 a / out let of reformer,
24.2
Kg/cm2 a Efficiency of PSA, %
-
Whether pre-reformer is used (Yes/No)
YES
Operating condition for pre-reformer Pressure,
28.4 –27.9 / 490 -419
Kg/cm2 a (Inlet/ Outlet) / Temperature, deg.C (Inlet/ Outlet) Pre-reformer location (After/before feed
BEFORE FEED PREHEAT COIL
preheat coil) No. of Stages of Reforming
SINGLE
Feed temperature/pressure @ unit b/l (deg.C/
SRN (C5 –90)
Kg/cm2 a)
FROM RFCCU 40 / 5.0
Product temperature/pressure @ unit b/l (deg.C/
45 (MAX.) / 21 (MIN.)
Kg/cm2 a) Product Quality (Percent purity of Hydrogen)
99.9
Product yield, (NM3/hr)
47252
Extent of Air cooling, deg.C
20
Product run down temperature cooler inlet, deg.C
131
DM water heater heating (by syngas) if any Temperature of syn gas to the Exchanger, deg.C
RICH GASES
270 / 331 / 224
51 / 13.5
Heat recovered, (MM Kcal/hr)
2.29 / 4.22 / 3.30
Flue gas heat recovery Air pre-Heaters for the furnace Type of air Pre-
9.15 + 4.85
Heater (MM Kcal/hr) Steam generation duty in furnace (MM Kcal/hr)
9.04
Steam superheating duty in furnace (MM Kcal/hr)
7.84
Feed super heating duty in furnace (MMKcal/hr)
4.26 + 4.12
Final flue gas temperature, deg.C
159
Cooling Water
Design Data
Type (Once thru/circulating Sea/Water)
Circulating Water
Flow Rate (m3/hr)
120
Supply / return temperature (oC)
33 / 42
Flow rate, (T/hr)
Generation Consumption
Net
in the unit
import / export
HP Steam
68.6
36.0
- 32.6
MP Steam
-
-
-
LP Steam
-
2.0
+ 2.0
Others, if any
-
-
-
Operating Pressure & Temperature (Kg/cm2a)/ oC) HP Steam
36.0 / 400
MP Steam
11 / 275
LP Steam
4.5 / 180
Total Power Consumption in HGU (KW)
2000
D DIES EL HYDRO H OTRE EATI ING UNIT T ( D DHDT T )
GEN NERAL L DATA A
Design Capacity : D 2.2 MMTPA M S Stream m Factor : 8,000 0 hourss per ye ear :40% wn T Turndo % of de esign ca apacity y l Techn O Origina nology : UOP P, USA T D Date of f Commissionin ng : 20TH Oct’2 2002
EDSTO OCK DE EFINIT TION FEE
e design feed is a blend con ntaining g Straiight run Gasoil from low The sulphur im mported crude e (SRGO-LS), Straig ght run Gasoil from high h mported crude e from middle e east ((SRGO--HS), T Total Cy ycle Oil sulphur im om FCC CU (TCO O), Ligh ht Coke er Gaso oil from m Cokerr unit (L LCGO) with fro bellow mentioned d prope erties. (Table II-1)
Capacity MTPA (BPSD)
Case I
Case II
Case III
2.2 (47460)
2.2 (47855)
2.2 (47433)
SRGO-LS
46.0
SRGO-LS
6.0
SRGO-LS
26.0
SRGO-HS
4.6
SRGO-HS
48.0
SRGO-HS
12.7
TCO-FCC
31.6
TCO-FCC
31.0
TCO-FCC
39.2
LCGO
17.8
LCGO
15.0
LCGO
22.1
Composition Wt %
Case-1
Case-2
Case-3
API
30.26
31.62
30.18
Sulphur Wt %
0.726
1.92
1.37
Nitrogen wppm
664
610
758
17
16
21
Flash Point Ԩ
50
50
58
Pour Point Ԩ
3
3
3
0.86
0.91
1.07
0.3
0.3
0.3
Cetane Number
42.4
43.7
39.0
Silicon wppm
0.53
0.45
0.66
5.1
4.8
6.2
Case I
Case II
Case III
IBP
141
142
139
10 %
192
192
182
30 %
263
259
248
Bromine Number g/100g
Metal Ni+V wppm Iron wppm
Chloride wppm
ASTM Distillate
50 %
288
284
280
70 %
309
309
305
90 %
344
348
343
95 %
359
361
358
FBP
399
402
399
MAKE UP HYDROGEN The make-up hydrogen for the Hydrotreater will be supplied from the Hydrogen Unit having the following characteristics:99.5 Vol. % min Hydrogen Purity Chloride 1 vppm max Balance will comprise mainly methane & trace of CO, CO2 and N2. PRODUCT SPECIFICATION The Hydrotreated products shall be routed to the storage and the properties of the Diesel will meet the following specification. Properties
Diesel
Flash Point Ԩ
Not less than 40
Cetane Number
48.5
Sulphur Wt ppm
2000
Pour Point Ԩ
Equal to or Not higher than the Feed
Stability UOP 413, mg/100ml
< 1.6
Distillation Temperature for Recovery
Not more than the Feed
@ 85.5 Vol % and 95 Vol % Water Content
ATTRIBUTES
CAPACITY
0.05 Vol % max
UNITS
DESIGN
MMTPA
2.20
ATTRIBUTES
T’PUT
UNITS
DESIGN
M3/Hr
314.0
CATALYSTS TK-10
MT
Quantity
Years
5
Life
M3/Kg of Cat
44.5
Feed Processed TK-711
Quantity
MT
Life
Years
2
Feed Processed
M3/Kg of Cat
17.8
RF200
MT
Quantity
Life
Years
2
Feed Processed
M3/Kg of Cat
17.8
HC-K
MT
Quantity
Life
Years
5
Cycle Length
Years
2
Feed Processed
M3/Kg of Cat
44.5
H2 consumption (SOR/EOR)
NM3/M3 OF fresh
182.3/191.2
Feed SWEET DIESEL YIELD(SOR/EOR)
Wt% of Fresh Feed
96.9/95
GAS TO OIL RATIO
NM3/M3 OF FEED
250-500
MMKCal/Hr.
20.08
Efficiency
%
91.94
H2 Partial Pressure
Kg/Cm2
59 Min.
Recycle Gas Purity
%H2
70 Min
Kg/Cm2
102.9/102.4
OPERATING PARAMETERS FURNACE
Heat Duty
REACTOR-1 Inlet Pressure(SOR/EOR)
ATTRIBUTES
UNITS
DESIGN
Inlet Temprature(SOR/EOR)
0
323/370
Outlet Temprature(SOR/EOR)
0
368/406
Delta P. Across reactor(EOR)
Kg/Cm2
3.1
Delta T Across Reactor(EOR)
0
35
WABT Ist Bed
0
WABT IInd bed
0
Quench Flow
Kg/Hr.
20000
Inlet Temprature(SOR/EOR)
Kg/Cm2
99.6/99.1
Outlet Temprature(SOR/EOR)
0
323/372
Delta P Across Reactor(EOR)
0
368/406
Delta T Across Reactor(EOR)
Kg/Cm2
8.8
WABT Ist Bed
0
34
WABT IInd bed
0
Ist Quench Flow
Kg/Hr.
21600
Iind Quench Flow
Kg/Hr.
21600
H.P.SEPARATOR Pr.
Kg/Cm2
84.5
C C
C C C
REACTOR-2 Inlet Pressure(SOR/EOR)
C C
C C
ATTRIBUTES
UNITS
DESIGN
STRIPPER Pressure
Kg/Cm2
8.5
Top Temprature
0
165
Stripping Steam
MT/Hr.
6.8
C
Utility Summary for DHDT A.
Cooling Water
: Design Data
Type (Once thru/circulating
: Circulating Water
Sea/Water)
B.
Flow Rate (m3/hr)
: 2480
Supply / return temperature (oC)
: 33 / 42
Steam
:
Flow rate, (T/hr)
: Generation Consumption in the unit
Net import / export
-
HP Steam
: -
98.7
98.7
MP Steam
: -
10.6
10.6
LP Steam
: 87.0
1.0
- 86.0
Operating Pressure & Temperature (Kg/cm2a)/ oC)
:
HP Steam
: 36.0 / 400
MP Steam
: 11 / 275
LP Steam
: 4.5 / 180
Others, if any
: -
C.
Fuel
:
i) Heater absorbed duty, (MM Kcal/hr)
: -
ii) Heater efficiency
: -
iii) Heater fired duty (MM Kcal/hr)
: 2.7
D.
Or fuel consumed (T/hr) Total Power Consumption in HGU
: 9183
(KW)
PROCESS DESCRIPTION DHDT is installed for upgradation of Coker Gas Oil as well as quality improvement of few diesel components. The feed is mixed with Hydrogen-rich recycle gas & make up Hydrogen after being compressed in respective compressor and reheated by exchanging heat with hot reactor effluent. The mixture is further heated to the desired reactor temperature in a fired heater and is fed to the Hydrotreater reactor. Hydrotreating reactions are exothermic in nature and hence recycle gas is introduced as quench between the beds of the reactor to cool reaction fluid and redistribute vapour and liquid. The reactor effluent is cooled by heat exchange with feed and recycle gas before it is finally cooled in the air cooler and then flashed in the separator. The hydrogen rich separator gas is scrubbed with Lean Amine in Recycle Gas Amine scrubber to remove H2S, recompressed and combined with make up hydrogen coming from the Hydrogen plant and then returned to the reactor. Sour water is coalesced and removed from the bottom of the separator and sent to Sour Water Stripper Unit. The liquid from the separator is sent to a stripper via heat exchangers part of the condensed stripper overhead is pumped to stripper as reflux and rest is taken out as naphtha product. The uncondensed vapour ex stripper is sent to Amine Absorption Unit. The bottom product from stripper gas to storage as hydrotreated gas oil component after cooling.
OIL MOVEMENT & STORAGE Oil Movement & Storage (OM & S) is an important function of the Production Department. In Barauni Refinery, OM & S section consists of the following sub sections:
OM & S - Receipt.
OM & S - Despatch
LPG
Utilities
Coke Handling
OM & S PUMPS A. Crude and Intermediate Product are pumped through centrifugal pumps. Besides there are booster pumps in the pipelines and transfer pumps to the marketing. B.Finished Product
Pumps of centrifugal type are present to pump the
finished products like SRN, MS, MRN, SKO, HSD, LDO, LSHS, caustic transfer, slack and paraffin wax, Phenol Extract / CBFS / FO, LPG Bottling, LPG bulk loading, LPG Intank pump for bulk loading from Mounded bullets.
TANK WAGON LOADING GANTRY A. White Oil Loading Gantry Maximum 38 number of BG T/Ws can be placed in one line for the products (SRN/MS/SKO/HSD). 24 numbers of BTPN T/Ws can be placed in one line. B.Black Oil Loading Gantry It has two rail lines, both lines have BG and MG tracks. In this gantry loading of LDO/Phenol Extract/LSHS can be done in BG/MG Tank wagons. FO/LDO points are multiple and LSHS/Phenol Extract points are common.
C.Lube Oil Loading Gantry It has two rail lines known as line no. 5 and 6. only B.G. T/Ws can be loaded in this gantry with CBFS-500, Rubber Extender Oil. However, none of above products are loaded now in this gantry. LPG Storage in Horizontal tanks(bullets), Horton spheres(presently there are four) Horton spheres service. These are: Two of capacity 300 M3 & Two of capacity 1500 M3 each and 6 nos. Mounded bullets capacity 1500 M3 each.
UTIITIES PROCESS COOLING WATER SYSTEM 1.0
Cooling Water
2.0
Type
(Once
through
/ Both
Circulating fresh water or sea water) 2.1
No. of cells
5
Min.
Nor.
Max.
5.0
-
-
2.2
Supply Press. , Kg/Cm2
2.3
Design wet bulb temp., oC
29.0
-
-
2.4
Supply temp., oC
33.0
-
-
2.5
Return temp., oC
45.0
-
-
2.6
Return Press., Kg/Cm2 (a)
0.3
At the top of cooling tower 2.9
If any treatment done
Yes
If yes, Detailes
2.11
Cooling
water
Non -oxidising biocides used.
circulating
7 motors
pumps 2.11.1
Type
Centrifugal
2.11.2
No. (Working+ standby)
2.11.3
Capacity, M3 / Hr ( Rated /
4 working +3 standby
3825
Normal)
2.11.4
Head, Kg/Cm2
2.11.5
Power
5.0
Consumption,
KW
720 kw (each)
(Rated / Normal /
CRU COOLING WATER SYSTEM
1.0
Cooling Water
2.0
Type
(Once
through
/ Circulating fresh water
Circulating fresh water or sea water)
2.1
No. of cells
2
Min.
Nor.
2.2
Supply Press. , Kg/Cm2 (a)
5.0-6.0
2.3
Design wet bulb temp., oC
29.0
2.4
Supply temp., oC
33.0
2.5
Return temp., oC
45.0
2.6
Return Press., Kg/Cm2 (a)
3-5
Max.
At the top of cooling tower
2.9
If any treatment done
If yes, Detailes
Yes
Non -oxidising biocides used.
2.10
Cooling water Make up, M3 / Hr
50
2.11
Cooling water circulating pumps
2 nos.
2.11.1
Type
2.11.2
No. (Working+ standby)
Centrifugal
1 working +1 standby
2.11.3
Capacity, M3 / Hr ( Rated /
1200
Normal)
2.11.4
Head, Kg/Cm2
5.6
2.11.5
Power Consumption, KW (Rated
160
/ Normal /
BXP COOLING WATER SYSTEM
1.0
Cooling Water
2.0
Type
(Once
through
/ Circulating fresh water
Circulating fresh water or sea water)
2.1
No. of cells
5
Min.
2.2
Supply Press. , Kg/Cm2 (a)
5.0-6.0
2.3
Design wet bulb temp., oC
29.0
2.4
Supply temp., oC
32.0
2.5
Return temp., oC
45.0
2.6
Return Press., Kg/Cm2 (a)
3-5
Nor.
Max.
At
the
top
of
cooling
tower 2.9
If any treatment done
Yes
If yes, Detailes
2.10
Non -oxidising biocides used.
Cooling water Make up, M3
-
/ Hr 2.11
Cooling water circulating
5 nos.
pumps 2.11.1
Type
Centrifugal
2.11.2
No. (Working+ standby)
2.11.3
Capacity, M3 / Hr ( Rated
3 working +2 standby
4450
/ Normal)
2.11.4
Head, Kg/Cm2
2.11.5
Power
5.1
Consumption,
KW
6.6 (each)
(Rated / Normal /
TPS COOLING WATER SYSTEM 1.0
Cooling Water
2.0
Type
(Once
through
/
fresh water or sea water)
Circulating Circulating fresh water
2.1
No. of cells
3
Min.
2.2
Supply Press. , Kg/Cm2 (a)
3.5
2.3
Design wet bulb temp., oC
29.0
2.4
Supply temp., oC
32.0
2.5
Return temp., oC
40
2.6
Return Press., Kg/Cm2 (a)
2.5
Nor.
Max.
At the top of cooling tower
2.7
Blow
Down
Quantity,
M3
/
Hr
75-100
(continuous) 2.8
Cooling water balance, M3 / Hr
2.9
If any treatment done
If yes, Detailes
Total to TPS
Yes
Non -oxidising biocides used.
2.10
Cooling water Make up, M3 / Hr
75
2.11
Cooling water circulating pumps
3 nos.
2.11.1
Type
Centrifugal
2.11.2
No. (Working+ standby)
2.11.3
Capacity, M3 / Hr ( Rated / Normal)
2.11.4
Head, Kg/Cm2
2.11.5
Power
Consumption,
2 working +1 standby
3600
5.1
KW
(Rated
/
450
Normal /
UTILITIES A gas flare, alternatively known as a flare stack, is an elevated vertical conveyance found accompanying the presence of oil and gas wells, rigs, refineries, chemical plants, natural gas plants, and landfills. They are used to eliminate waste gas which is otherwise not feasible to use or transport. They also act as safety systems for non-waste gas and is released via pressure relief valve when needed to ease the strain on equipment. They protect gas processing equipments from being overpressured. Also in case of an emergency situation, the flare system helps burn out the total reserve gas. The size and brightness of the resulting flame depends upon how much flammable material was released. Steam can be injected into the flame to reduce the formation of black smoke. The injected steam does however make the burning of gas sound louder, which can cause complaints from nearby residents. Compared to the emission of black smoke, it can be seen as a valid trade off. In more advanced flare tip designs, if the steam used is too wet it can freeze just below the tip, disrupting operations and causing the formation of large icicles. In order to keep the flare system functional, a small amount of gas is continuously burned, like a pilot light, so that the system is always ready for its primary purpose as an over-pressure safety system. The continuous gas source also helps diluted mixtures achieve complete combustion.
Flare gas recovery system There are two compressors for recovery of waste flare gas, compressing the same and put it back to refinery fuel gas header for consumption in furnace/ boilers. Capacity of each compressor is 450 NM3/Hr. Water treatment Many industries have a need to treat water to obtain very high quality water for demanding purposes. Water treatment produces organic and mineral sludges from filtration and sedimentation. Ion exchange using natural or synthetic resins removes calcium, magnesium and carbonate ions from water, replacing them with hydrogen and hydroxyl ions. Regeneration of ion exchange columns with strong acids and alkalis produces a wastewater rich in hardness ions which are readily precipitated out, especially when in admixture with other wastewaters.
Treattment o of indusstrial w wastewa ater The differen d nt typess of con ntamina ation of f waste ewater require r a varie ety of strate egies to o removve the c contamination.. Solidss removval Most solids c can be remove ed using g simple e sedime entation techn niques with w the e solidss recove ered as slurry or slud dge. Verry fine solids and a solids with h densitties clo ose to the denssity of water pose p special prroblemss. In such case e filtrattion or ultrafiltration n may be b requiired. Allthough h, floccu ulation may be e used, using a alum salts or th he addition of polyele ectrolyttes. Oils and a gre ease re emoval: API oiil-wate er separator
A separator is a gra avity se eparatio on devic ce desig gned by y using Stokes The API Law to o define the rise r velo ocity of f oil dro oplets b based on their density y and
size. The T dessign is based b o on the specific s c gravitty diffe erence b between the oil and th he wasttewaterr because thatt differrence iss much smaller s r than the t specif fic gravvity diff ference e betwe een the e suspen nded so olids and d waterr. The suspe ended so olids se ettles to o the bottom b o the sseparattor as a sedime of ent laye er, the oiil rises to top of the separattor and d the cle eansed wastew water iss the [3] middle e layer betwee en the o oil layer and th he solid ds. Typica ally, the oil lay yer is skimmed d off an nd subsequently re-prrocesse ed or dispossed of, and the e botto om sediment la ayer is rremoved by a c chain an nd fligh ht scraper (or ssimilar device)) and a sludge s p pump. T The watter laye er is sen nt to atment consistting usu ually of a Electtro-flottation m module for f further trea emoval of o any rresidual oil and d then tto some e type o of biolo ogical additiional re treatm ment un nit for remova al of und desirab ble disso olved chemical compo ounds.
A typical parrallel plate sep parator Parallel plate e separa ators are similar to API A sep paratorss but th hey incllude tilted d paralle el plate assemb blies (to enhan nce the e degree e of oil--water separation). The parallel plates prrovide more m su urface for susspended d oil drople ets to c coalesce e into la arger globules g s. However, the e paralllel plate es. The e resultt is thatt a para allel pla ate sepa arator require r es signif ficantly y less sp pace th han a convventiona al API separat s tor to achieve a the sam me degree of separattion.
e organ nics Removval of biodegrradable
Biodegradable organic matterial of f plant or anim mal origin is usu ually po ossible to t extende ed conventiona al waste ewater treatm ment pro ocessess such as a treat using e activa ated slu udge or tricklin ng filte er. Prob blems ca an arise e if the wastew water iss excesssively d diluted with wa ashing water or o is hig ghly con ncentra ated suc ch as neat blood b orr milk. The T pre esence of clea aning ag gents, disinfectants, pestic cides, o or antibiotics c can have e detrim mental impacts on tre eatmen nt proce esses.
Activa vated slludge process p s
Activa ated slu udge is a bioch hemical processs for ttreating g sewag ge and in ndustriial waste ewater tthat usses air ((or oxyg gen) and micro oorganissms to biologic cally oxidiz ze organ nic pollutants, produc cing a waste w sludge (o or floc) contain ning the e oxidiz zed matterial. Trick kling filter pro ocess
A schematic cross-ssection n of the e contac ct face of the bed me edia in a trickling filter
A typical com mplete tricklin t ng filterr system m A tric ckling f filter co onsists of a be ed of ro ocks, grravel, sllag, pea at moss, or plastic media a over which w w wastewa ater flo ows dow wnward and con ntacts a layer (or fillm) of m microbial slime e coveriing the bed me edia. Ae erobic c conditio ons are maintained b by force ed air f flowing through h the b bed or by b naturral convvection of cess invvolves a adsorption of organic o compou unds in the wa astewater air. The proc e micro obial slim me laye er, diffu usion of f air intto the slime s la ayer to provide e by the the ox xygen rrequired d for th he biochemical oxidattion of the org ganic co ompound ds. The end e prod ducts in nclude c carbon dioxide e gas, w water an nd othe er produ ucts of the ox xidation n. As th he slime e layer thicken ns, it be ecomes difficu ult for the t air to penettrate th he layerr and an n inner anaerob bic laye er is forrmed. Treattment o of othe er organ nics ls includ Synth hetic orrganic materia m ding solvents, paints, pharmaceutic cals, pestic cides, c coking product p s and so forth h can be e very difficult d t to tre eat. Treattment m methodss are of ften spe ecific to t the m materia al being treated. Metho ods inc clude Advanced Oxid dation Process P ing, disstillatio on, adssorption n, vitrification, incine eration,, chemical imm mobilisa ation or landf fill disp posal. be capable of b biologic cal Some materiials such as some dettergentss may b degra adation and in such s ca ases, a modifie m ed form m of wasstewate er treattment can c be use ed. Treattment o of acidss and a alkalis Acidss and alk kalis ca an usually be ne eutralissed und der conttrolled conditions. Neutrralisatio on freq quently produces a pre ecipitatte that will req quire trreatment as a solid ressidue th hat may y also be b toxic c. In som me case es, gassses may y be ed requ uiring trreatmen nt for the t gas stream m. Some e other forms of evolve treatm ment arre usually requ uired fo ollowing g neutra alisation n.
Waste streams rich in hardness ions as from de-ionisation processes can readily lose the hardness ions in a buildup of precipitated calcium and magnesium salts. This precipitation process can cause severe furring of pipes and can, in extreme cases, cause the blockage of disposal pipes.Treatment is by concentration of de-ionisation waste waters and disposal to landfill or by careful pH management of the released wastewater. Treatment of toxic materials Toxic materials including many organic materials, metals (such as zinc, silver, cadmium, thallium, etc.) acids, alkalis, non-metallic elements (such as arsenic or selenium) are generally resistant to biological processes unless very dilute. Metals can often be precipitated out by changing the pH or by treatment with other chemicals. Many, however, are resistant to treatment or mitigation and may require concentration followed by landfilling or recycling. Dissolved organics can be incinerated within the wastewater by Advanced Oxidation Processes.
Quality Control Modernization and Infrastructure Development Modernization and Renovation of Quality Control Laboratory is under progress. A new laboratory building is under construction and the old laboratory building is being renovated phase wise. For smooth commissioning and operation of MSQ project and other test facilities, a new laboratory is being set up. Carbon, Nitrogen, Sulphur & Chloride Analyser: This instrument is capable to check these elements in sub ppm level in different petroleum products. This is a microprocessor based automatic instrument. GC Oxygenates: This special type of Gas Chromatograph is used to check oxygenates content in gasoline. This parameter is required to be tested to certify BSIII MS. Potentiometric Titration System: It is used for evaluation of Diene content (MAV), Bromine number, Bromine index etc. in petroleum products. RON Engine: An advance model RON engine is under final stage of procurement/arrival.
Developmental Studies Viscosity grade Bitumen: BIS introduced new Specifications for Bitumen (IS 73:2006) implemented where some new specification parameters like viscosity at different temperature and vacuum were incorporated. Test facilities for these new parameters were developed. Antioxidant Dose Optimization Study: To optimize the dosing rate of anti oxidant in MS blending component ex-FCCU, a study was conducted. Effectiveness, optimum dosing quantity, reaction time of antioxidant was determined and informed to production for implementation. BS III HSD Certification: Till date HSD produced in BR is certified under BSII specification. Test facilities have been developed to certify HSD under BSIII specification. Few batches of BSIII HSD have already been certified. BS III MS Certification: Test facility has been developed for carrying out additional test required for certification of BSIII MS. BSIII MS Production: Blend study for production of BSIII MS before MSQ commissioning with imported isomerate & MTBE. CLO up-gradation: CLO is a low demand / low value stream ex-FCCU. It has certain disadvantages like high viscosity, high density, high ash etc. Blend study was conducted to find out the possibility of using CLO as FO blending component. IFO up-gradation & Certification of FO: To find the possibility of selling IFO as Fuel Oil, samples from all IFO tanks were collected and tested for all Fuel Oil parameters. Certification of F.O., a new product is done w.e.f. April 2009. Trial runs conducted to assess FO/IFO quality .Further blend study and optimization of PPD doses were done to meet certification of winter grade F.O. ATF Production: ATF produced from different crude mix was found to be failing in Acid number and JEFTOT test. Study was conducted to reduce acid number by caustic wash followed by water wash. Commissioning of DHDT 3rd Reactor: Laboratory support was given by way of continuous product quality evaluation for commissioning of DHDT rd
3 reactor. DHDT PGTR: For performance evaluation, Quality of all rundown streams was evaluated during DHDT PGTR.
Loss Control (I)Quality Give Away (QGA) Prevention: To generate awareness among all concerned, QGA value in terms of money was calculated for all certified tanks and circulated. Significant improvement has taken place in this field. (II)Portable Flue Gas Analyser: Two portable Flue Gas Analysers were procured and put in service for in-situ analysis of flue gas to improve furnace efficiency. (III)LPG Loss Control: Fuel gas was checked on regular basis for LPG slippage. Environmental Management Effluent Quality Monitoring under new MINAS Monitoring of Raw water and drinking Quality Cost Reduction in Utility Consumption Use of hydrogen & nitrogen generator in place of cylinders Use of Orsat apparatus in place of dragger tube for checking H2S content Use of low cost Argon in place of Helium for operation of low level Sulphur and chloride analyser.
