Design for Production of Propylene

79

Propylene Top Producers

Faculty of Chemical Engineering(FChE)Faculty of Chemical Engineering(FChE)
Faculty of Chemical Engineering
(FChE)
Faculty of Chemical Engineering
(FChE)

SKKK 4153 PLANT DESIGN
2014/2015-SEM 1

FINAL REPORT
PROPYLENE PRODUCTION PLANT

LECTURER
ASSOC. PROF. IR. DR. SHARIFAH RAFIDAH WAN ALWI

DESIGN TEAM
EQUINOX

NO.
TEAM MEMBERS
MATRIC NO
1.
EMAD MOHAMMED NOMAN AL-DHUBHANI
AA103001
2.
MUHAMMAD FAIRIS BIN HADIPORNAMA
A11KK0022
3.
KELVIN RAWING SEBASTIAN
A11KK0065
4.
NUR FADZLYANA BINTI HAMDAN
A11KK0035
5.
MIMI KHAIRIAH BINTI AWANG
A11KK0169

TABLE OF CONTENTS

Page
TABLE OF CONTENTS 2

CHAPTER 1 INTRODUCTION
Background of Propylene 5
Uses of Propylene 6
Propylene Manufacturing 8
Market Survey 9
1.4.1 Introduction 9
1.4.2 Production of Propylene 9
1.4.3 Propylene Consumption 11
1.4.4 Outlook for Production of Propylene in Malaysia 12
1.4.5 Market Prices of Polypropylene 12
1.5 Raw Materials 14
1.5.1 Source of Raw Materials 14
1.5.2 Raw Material Specifications 14

CHAPTER 2 PROCESS SYNTHESIS STEP
2.1 Step 1-Eliminate Differences in Molecular Type 15
2.2 Step 2- Distribute the Chemicals 24
2.2.1 Basic Material Balances 24
2.3 Step 3- Eliminate Differences in Compositions 26
2.4 Step 4 –Eliminate Differences in Temperature, Pressure and Phase 28
2.5 Step 5- Task Integration 30

CHAPTER 3 MATERIAL,ENERGY BALANCES AND PROCESS SIMULATION
3.1 Material Balance 33
3.1.1 Overall Mass Balances 33
3.1.2 Mass Balances for Separation Unit 1 35
3.1.3 Mass Balance for Mixer 36
3.1.4 Mass Balance for Reactor 37
3.1.5 Mass Balance for Separation Unit 2 38
3.1.6 Mass Balance For Separation Unit 3 39
3.1.7 Mass Balance For Separation Unit 4 40
3.2 Energy Balances 40
3.3 Simulation Result from ASPEN HYSYS 60
3.3.1 Material Balance 60
3.3.2 Energy Balances 60
3.4 Percentage Differences between Manual Calculation and HYSYS 61
3.4.1 Mass Balances 61
3.4.2 Energy Balances 61

CHAPTER 4: HEAT INTEGRATION
4.1 Process Energy Integration 62
4.2 Algorithm Table 63
4.3 Heat Exchanger Network 64
4.4 Process Flow Diagram Heat Exchanger Network 64

CHAPTER 5: PROCESS OPTIMIZATION 66

CHAPTER 6: EQUIPMENT SIZING AND COSTING
6.1 Introduction 69
6.2 Reactor 69
6.2.1 Sizing of reactor 69
6.2.2 Costing of reactor 70
6.3 Pump 70
6.3.1 Sizing of pump 70
6.3.2 Costing of pump 70
6.4 Distillation columns (S1) 70
6.4.1 Sizing and costing of the main vessel 70
6.4.2 Sizing and costing of the reflux drum 71
6.4.3 Sizing and costing of the condenser 72
6.4.4 Sizing and costing of the re-boiler 73
6.5 Compressor 74
6.5.1 Sizing of compressor 74
6.5.2 Costing of compressor 74
6.6 Heat exchanger (HE2) 74
6.6.1 Sizing of Heat exchanger (HE 2) 74
6.6.2 Costing of Heat Exchanger (HE2) 76

CHAPTER 7: TOTAL CAPITAL INVESTMENTS 77

CONCLUSIONS 79

APPENDICES A
APPENDICES B
APPENDICES C
APPENDICES D
APPENDICES E

CHAPTER 1

INTRODUCTION

1.1 Background of Propylene

Propylene, also called propene is generally described as a volatile and a colorless gas at room temperature. It has same empirical formula with cyclopropane but different ways of atom connected. Propylene is categorized as a alkene hydrocarbon compound with a molecular formula of C3H6. The presences of the double bond make it slightly lower boiling point than propane and thus more volatile. The existences of natural propylene are in the environment from sources such as vegetation and combustion such as fires, motor vehicle exhaust, and tobacco smoke. Propylene is not expected to persist in the environment. Since propylene is a gas, the exposure of propylene into the air is expected to be lower amount when released into the environment. Because of its relatively short half-life in the atmosphere and typically low environmental concentrations, propylene's contribution to potential global warming is considered minor and its ozone depletion potential is negligible.

Figure 1: Structural formula of Propylene.

Propylene reacts violently with oxide of nitrogen and also a number of other substances and condition. Essentially all of the propylene produced for chemical purposes is consumed as a chemical intermediate in other chemical manufacturing processes. This hydrocarbon is widely used in the manufacture of cumene, resins, fibres, elastomers and other chemicals which enable the manufacture of many chemicals and plastics. In addition to its use as a chemical intermediate, propylene is produced and consumed in refinery operations for the production of gasoline components.

1.2 Uses of Propylene
Propylene is a major product of the petrochemical industry. It is one of the highest volume chemicals produced globally. Propylene is primarily used as an intermediate for the production of other chemical raw materials that are subsequently used to manufacture a large variety of substances and products. Manufacture of polypropylene, a widely used plastic, consumes more than half of the world's production of propylene. Propylene is also used in the manufacture of acrylonitrile, oxo process chemicals, cumene, isopropanol, polygas chemicals, and propylene oxide. Table 1.1 below highlights several of the main applications of propylene and its derivatives.

Table 1.1: Selected Propylene Application

Product Application

Application Description

Polypropylene

Polypropylene is used to make many well-known plastic products.
Polypropylene resins can be injection molded and extruded (into fibers, film, and sheets) to make a variety of products.
Polypropylene may also be blow-molded or thermoformed, but these processes are less often used.
Polypropylene is extremely corrosion resistant, lightweight, flexible, and formed or welded.

Propylene Oxide

Used mainly as a chemical intermediate in the production of polyurethane polyols and propylene glycols.
Used in the manufacture of propylene glycol, which helps to make antifreeze, resins for reinforced plastics, pharmaceuticals, packaging materials, dyes, and hydraulic fluids, and humectants for foods, drugs, cosmetics, and pet foods.
Derivatives of propylene oxide include polyether polyols; propylene glycol; di- and tripropylene glycol; poly (propylene glycol)s; surfactants; glycol ethers; and isopropanolamines.

Isopropanol

A variety of solvent applications, such as in printing inks, surface coatings, and as a solvent for resins, shellacs, and gums.
As a component of personal care products, such as after-shaves; and as an antiseptic and disinfectant, such as rubbing alcohol.
Used in the production of acetone, methyl isobutyl ketone (MIBK), iso-propylamines and isopropyl acetate.

Cumene

Alkylation of benzene with propylene yields cumene.
Consuming in phenol production for the manufacture phenolic resins, caprolactam and bisphenol A.

Ethylene-Propylene Elastomers

About half of the worldwide production of EP elastomer goes into the manufacture of automobile body and chassis parts, hoses, weatherstripping, and tires.
Also used to make thermoplastic polyolefin elastomers, polymer modifiers, and other products used in automobile components beside used in single-ply roofing.

Oxo Process Chemicals

Propylene is used to manufacture Isobutyraldehyde, which is converted to isobutanol solvent for surface coatings.
Propylene is also used to make n-Butyraldehyde, which is converted to n-butanol or 2-EH. n-Butanol is a solvent for lacquers and coatings, and is an intermediate for several chemicals.

Polygas Chemicals (nonene, dodecene, heptenes)

Refinery production of polymer gasoline also yields nonene, dodecene and heptene and propylene is consumed to yield these polygas chemicals.
Nonene is used in nonylphenol and isodecyl alcohol that act as an intermediate for surfactants, lubricating oil additives, and phosphite antioxidants.
Heptenes are consumed to make isooctyl alcohol, which is used in the manufacture of another phthalate ester.

1.3 Propylene Manufacturing

Lotte Chemical Titan Holding Sdn. Bhd. is one of the manufacturer and supplier of propylene, located at Pasir Gudang, Johor Bahru. This company will be the benchmark of Equinox Team to design a plant that can produce 100,000 lb/hr of propylene. The team will propose a variety of production reactions of propylene and there are several production processes such as catalytic dehydrogenation of propane, reformation of olefins reaction (metathesis reaction), and the conversion of methanol to propylene. The most sustainable and economically reaction processes will be chosen for the plant design. This includes the comparison between the cost of raw materials, safety, environmental impacts, percentage yield of conversion, energy consumption, and other factors that might affect the reaction process.

1.4 Market Survey

1.4.1 Introduction
Market survey or market outlook will cover a review on the production and consumption of propylene in addition to that there will be another section to discuss the prices of propylene and its raw materials.
1.4.2 Production of Propylene
Propene production increased in (Europe and North America only) from 2000 to 2008, it has been increasing also in East Asia, most notably Singapore and China. Total world production of propene is currently about half that of ethylene. About 56% of the worldwide production of propylene is obtained as a co-product of ethylene manufacture, and about 33% is produced as a by-product of petroleum refining. About 7% of propylene produced worldwide is on-purpose product from the dehydrogenation of propane and metathesis of ethylene and butylenes; the remainder is from selected gas streams from coal-to-oil processes and from deep catalytic cracking of vacuum gas oil (VGO). The supply of propylene remains highly dependent on the health of the ethylene industry as well as on refinery plant economics.
In 2010, production of polypropylene represented 65% of total world propylene consumption, ranging from 53% in North America to more than 90% in Africa and the Middle East.
Table 1.2: Annual Production of Propene (Propylene)
World
80.0 million tones
Europe
14.3 million tons
US
14.3 million tones

Figure 1.1 below shows how the production of propylene increased from 11 million tons in 1994 to 16 million tons in 2007 but it had dropped since that time to 14.3 million tons in 2013.

Figure 1.1: Wastern European Propylene Capacity, Production and Consumption 1994-2013
Top world companies are leading the production of propylene with LyondellBasell, Netherlands on top of propylene producing companies by 2009. The top propylene-producing companies are listed as bellow:

Figure 1.2: Propylene top producers
1.4.3 Propylene Consumption
After experiencing zero growth or declines in 2008 and 2009, global propylene consumption grew at a rate of almost 7.5% in 2010, led by Asia at 11% year-on-year. The economic recession of 2008/2009 reflected both a reduction in pull-through demand for polypropylene, as well as a supply-chain inventory rundown, reminiscent of the early 1980s downturn. World petrochemical industries have historically witnessed very few upheavals that combined the effects of both energy volatility and depressed downstream demand.
The fifteen largest worldwide producers of propylene accounted for almost 51% of world capacity as of 2010, representing about the same level of concentration as five years ago. The most significant changes in the last two years have been Sinopec taking over the top spot, a position long occupied by ExxonMobil, and PetroChina jumping from the seventh spot to number four.
World consumption of propylene is forecast to grow slightly better than global gross domestic product (GDP) rates over the next five years. Average growth will be 5% per year, higher than GDP in general and higher than ethylene specifically, with growth for polypropylene being much better than that for polyethylene. Growth will be led by the Middle East, Asia, Central and Eastern Europe, and South America at 12.5%, 6.5%, 5%, and 4.5% per year, respectively. Asia is a mixed bag of growth rates with China and India at 8–10% annually and the mature economies of Japan, the Republic of Korea, and Taiwan at 1–2% per year. Near-term growth will be relatively slow in the mature economies of North America and Western Europe.

Figure 1.3: World Consumption of Propylene in 2010

1.4.4 Outlook for Production of Propylene in Malaysia
Malaysia's petrochemical sector has contributed significantly to the development of local downstream plastic processing activities. Malaysia is one of the largest plastics producers in Asia, providing a steady supply of feedstock materials for the plastic processing industry such as propylene.
Table 1.3: Production, Import, Export and Consumption of PP in Malaysia
Product
Unit: KTPA
2007
2008
2009
2010
change
Propylene
Production
839
870
867
808
0.3%

Import
33
40
25
8
0%

Export
78
97
95
50
-2%

Consumption
765
811
797
744
-5%

1.4.5 Market Prices of Polypropylene
Polypropylene prices are on the rise since the last decade and it is expected to continue rising as the demand increases for the chemical material, Capacity and Prices for Polypropylene - End-Use Sectors in Asia-Pacific to Drive Growth" 2014 market research report says worldwide polypropylene capacity increased at a Compound Annual Growth Rate (CAGR) of 5.2% from 2003, reaching 65 million tons per year (MMTY) in 2013, and is expected to continue rising to 86 MMTY by 2018, at a slightly higher CAGR of 5.8%. It forecasts China and Russia to be the leading contributors to future polypropylene capacity increases, and will account for a combined 45% of global additions over the next five years. As Malaysia is part of the global market it is normal for prices in Malaysia to be affected by the global prices, following are prices of PP and its raw material (propane, ethylene, butene and methanol) as achieved from ICIS.com.

Table 1.4: Propylene and raw products prices
Product
Price

RM/Ib
RM/Kg
Propylene
1.962
4.326
Propane
0.79
1.742
Ethylene
2.158
4.758
Butene
1.118
2.465
Methanol
1.635
3.609

1.5 Raw Materials
1.5.1 Source of Raw Materials
The raw material that is utilized in this process is liquefied petroleum gas (LPG) propane. The term LPG actually encompasses more than one variety of gaseous fuel. There are a number of hydrocarbon gases that fall into the classification of "LPG". Their common distinguishing characteristic is that they can be compressed into liquid at relatively low pressures. LPG is stored under pressure, as a liquid, in a gas bottle. It turns back into gas vapor when you release some of the pressure in the gas bottle by turning on your appliance. Almost all of the uses for LPG involve the use of the gas vapor, not the liquefied gas.
The gases that fall under the "LPG" label, including Propane, Butane, Propylene, Butadiene, Butylene and Isobutylene, as well as mixtures of these gases. The two most common are Propane and Butane.
The main supplier of LPG used in this process is Kleenheat Gas which is part of Wesfarmers Chemicals, Energy and Fertilizers, one of eight divisions of Wesfarmers Limited, with origins dating back to 1914.
They have a long history in the Australian gas industry with over 55 years of experience retailing and distributing Liquefied Petroleum Gas (LPG), over a decade of experience distributing Liquefied Natural Gas (LNG) and advancing technology through their brand EVOL LNG, and most recently retailing natural gas in Western Australia.
1.5.2 Raw Material Specifications
Table 1.5: LPG propane supplied by Kleenheat Gas Australia
LPG specification
CAS Number 74-98-6
Component
Mole percentage (%)
Propane
80
Butane
18
Butanes ,pentanes ,butadiene and heavier
2

CHAPTER 2

PROCESS SYNTHESIS STEP

2.0 SYNTHESIS STEPS
Process synthesis involves the selection of processing operations to convert raw materials to products, given that the states of the raw material and product streams are specified. The most widely accepted approach for process synthesis is introduced by Rudd, Powers, and Siirola (1973) in a book entitled Process Synthesis. There are 5 key synthesis steps which are:
Eliminate differences in molecular types
Distribute the chemicals by matching sources and sinks
Eliminate differences in composition
Eliminate differences in temperature, pressure, and phase
Task integration; combination of operations into unit processes and decide between continuous and batch processing

2.1 Step 1 – Eliminate Differences in Molecular Type
Propylene from Propane via Dehydrogenation

Dehydrogenation is an endothermic equilibrium reaction; it is carried out in the presence of heavy-metal catalyst (chromium). The following equation shows the propane dehydrogenation reaction:

Propane Dehydrogenation Reaction
About 86 wt% of propane is converted to propylene. To mitigate cracking reactions, dehydrogenation reaction in this technology occurs in conditions such as temperature ranges between 580 and 650 °C, and pressures slightly below atmospheric. For further information, Table 2.1 shows the thermophysical property data for this process.

Figure 2.1: Commercial process flow diagram (Dehydrogenation)

Figure 2.2: Detailed process flow diagram (Dehydrogenation)

Process Description of Propylene Dehydrogenation

The propane dehydrogenation process is used to supply polymer-grade propylene from propane to meet the growing propylene market, independent of a steam cracker or Fluid Catalytic Cracking (FCC) unit. It provides a dedicated, reliable source of propylene to give more control over propylene feedstock costs.

From Figure 2.2, the process flow diagram consists of a reactor section, product recovery section and catalyst regeneration section. Hydrocarbon feed is mixed with hydrogen-rich recycle gas and is introduced into the heater to be heated into the desired temperature (over 540 °C) and then enter the reactors to be converted at high mono-olefin selectivity. Several interstage heaters are used to maintain the conversion through supplying heat continuously since the reaction is endothermic. Catalyst activity is maintained by continuous catalyst regenerator (CCR) or shutting down reactors one by one and regenerating the reactor by the regeneration air, the continuous catalyst regenerator is where the catalyst is continuously withdrawn from the reactor, then regenerated, and fed back to the reactor bed.

Reactor effluent is compressed, dried and sent to a cryogenic separator where net hydrogen is recovered. The olefin product is sent to a selective hydrogenation process where dienes and acetylenes are removed. The propylene stream goes to a de-ethanizer where light-ends are removed prior to the propane-propylene splitter. Unconverted feedstock is recycled back to the depropanizer where it combines with fresh feed before being sent back to the reactor section.

Table 2.1: Physical And Chemical Properties Of Reactant And Product For Dehydrogenation Reaction
REACTION
PROPANE
PROPYLENE
HYDROGEN
Properties
Molecular formula
C3H8
C3H6
H2
Molar mass
44.10 g mol 1
42.08 g mol 1
2.016 g mol-1
Appearance
Colourless gas
Colourless gas
Colourless gas
Odor
Odourless
Gassy/ aromatic

Density
2.0098 mg mL 1 (at 0 °C, 101.3 kPa)
1.81 kg/m3, gas (1.013 bar, 15 °C)
613.9 kg/m3, liquid
0.08988 g/L (at 0 °C, 101.325 kPa)
Melting point
187.7 °C; 305.8 °F; 85.5 K
,185.2 °C ( 301.4 °F; 88.0 K)
13.99 K ( 259.16 °C, 434.49 °F)
Boiling point
42.25 to 42.04 °C; 44.05 to 43.67 °F; 230.90 to 231.11 K
47.6 °C ( 54 °F; 226 K)
20.271 K ( 252.879 °C, 423.182 °F)
Solubility in water
40 mg L 1 (at 0 °C)
0.61 g/m3

Vapor pressure
853.16 kPa (at 21.1 °C)
144.06 psia
100kPa (at 20 ºC)
Thermochemistry
Std enthalpy of
formation ΔfHo298
105.2– 104.2 kJ mol 1
+20.41 kJ/mol
0
Std enthalpy of
combustion ΔcHo298
2.2197– 2.2187 MJ mol 1
-2058.4 kJ/mol
-285.84 kJ/mol

To screen out whether this reaction will bring profit or not, the gross profit is calculated as shown below:
C3H8 C3H6 + H2

C3H8
C3H6
H2
lbmol
1
1
1
Molecular weight
44.09
42.08
2.016
lb
44.09
42.08
2.016
lb/lb of propylene
1.0478
1
0.048
RM/lb
0.79
1.96
11.30

Gross profit for reaction path 1 = 1.96(1) + 11.30(0.048) – 1.0478(0.79) = RM 1.67 /lb propylene
Propylene from Ethylene and Butenes via Metathesis

Metathesis is a general term for a reversible reaction between two olefins, in which the double bonds are broken and then reformed to form new olefin products. In order to produce propylene by metathesis, a molecule of 2-butene and a molecule of ethylene are combined to form two molecules of propylene. Some of the thermophysical property data is shown on table 2.

Metathesis Reaction

Figure 2.3: Commercial process flow diagram (Metathesis)

Figure 2.4: Detailed process flow diagram (Metathesis)
Process Description of Metathesis of Ethylene and Butene
Propylene is formed by the metathesis of ethylene and butene-2, and butene-1 is isomerised to butene-2 as butene-2 is consumed in the metathesis reaction. In addition to the main reactions, numerous side reactions between olefins also occur. Ethylene feed can be polymer grade ethylene or a dilute ethylene stream. Any saturated hydrocarbons, such as ethane and methane, do not react.
From Figure 2.4, fresh C4s (plus C4 recycle) are mixed with ethylene feed (plus recycle ethylene) and sent through a guard bed to remove trace impurities from the mixed feed. The feed is heated prior to entering the vapour phase fixed-bed metathesis reactor where the equilibrium reaction takes place. The reactor is regenerated in-situ on a regular basis. The catalyst promotes the reaction of ethylene and butene-2 to form propylene and simultaneously isomerises butene- 1 to butene-2. The per-pass conversion of butylene is greater than 60 per cent, with overall selectivity to propylene exceeding 90 per cent.
The product from the metathesis reactor is primarily propylene and unreacted feed. Reactor effluent is sent to the ethylene recovery tower where the unreacted ethylene is recovered and recycled to the reactor. The C2 tower bottom is processed in the C3 tower to produce propylene product and a C4 recycle stream. Purge streams containing non-reactive light material, C4s and heavier are also produced. Ultra-high purity propylene exceeding polymer grade specification is produced without a propylene fractionation system, since the only source of propane is that contained in the C4 and ethylene feeds.
Table 2.2: Physical And Chemical Properties Of Reactant And Product For Metathesis Reaction
REACTION
ETHYLENE
BUTENE
PROPYLENE
Properties
Molecular formula
C2H4
C4H8
C3H6
Molar mass
28.05 g/mol
56.10 g/mol
42.08 g mol 1
Appearance
Colorless gas
colorless
Colorless gas
Odor
Odorless
odorless
Gassy/ aromatic
Density
1.178 kg/m3 at 15 °C, gas
0.62 g/cm3
1.81 kg/m3, gas (1.013 bar, 15 °C)
613.9 kg/m3, liquid
Melting point
169.2 °C (104.0 K, -272.6 °F)
185.3 °C ( 301.5 °F; 87.8 K)
,185.2 °C ( 301.4 °F; 88.0 K)
Boiling point
103.7 °C (169.5 K, -154.7 °F)
6.47 °C (20.35 °F; 266.68
47.6 °C ( 54 °F; 226 K)
Solubility in water
3.5 mg/100 mL (17 °C)[

0.61 g/m3
Thermochemistry
Std enthalpy of
formation ΔfHo298
52.28 kJ mol 1
1.17 kJ/mol
+20.41 kJ/mol
Std enthalpy of
combustion ΔcHo298
-1410.99 kJ mol 1
-2718.6 kJ/mol
-2058.4 kJ/mol

To screen out whether this reaction will bring profit or not and whether it is better from reaction A, the gross profit is calculated as shown below:
C2H4 + C4H8 2C3H6

C2H4
C4H8
C3H6
lbmol
1
1
2
Molecular weight
28.05
56.10
42.08
lb
28.05
56.10
84.16
lb/lb of propylene
0.33
0.667
1
RM/lb
2.16
1.18
1.96

Gross profit for reaction path 2 = 1.96(1) – 2.16(0.33) – 1.18(0.667) = RM 0.46 /lb propylene

Table 2.3: Summary of Review and Screening of Alternative Processes

Dehydrogenation of propane
C3H8 C3H6 + H2
Metathesis of from Ethylene & Butenes
C2H4 + C4H8 2C3H6
Gross Profit
(Appendix 1)
RM 1.67 / lb propylene
RM 0.46/lb propylene
Type of process
Continuous process
Continuous process
Safety
Propane is flammable.
Butane and ethylene is flammable, and ethylene also may cause dizziness
By-product
Hydrogen
No by-product
Operating condition
Temperature: 560 – 650 ºC
Pressure : slightly below atmospheric pressure
Temperature: 90-100ºC
Pressure: 100 – 110 bar
Conversion
86% percent of conversion
90% percent of conversion
Flammability
Flammable
Flammable

From the table above, it shows that the dehydrogenation of propane reaction is a better process compared to the metathesis reaction.

2.2 Step 2 – Distribute the Chemicals

2.2.1 Basic Material Balance
ReactorT = 500 OCP = 1 bar gfrgfr gdm2C3H6m3H2R lb/hr C3H8m1C3H8F lb/hr C3H8ReactorT = 500 OCP = 1 bar gfrgfr gdm2C3H6m3H2R lb/hr C3H8m1C3H8F lb/hr C3H8
Reactor
T = 500 OC
P = 1 bar
gfrgfr

gd

m2C3H6
m3H2
R lb/hr C3H8
m1C3H8
F lb/hr C3H8
Reactor
T = 500 OC
P = 1 bar
gfrgfr

gd

m2C3H6
m3H2
R lb/hr C3H8
m1C3H8
F lb/hr C3H8

Overall Reaction Equation :
C3H8 C3H6 + H2

Basis : 100000 lb/hr of propylene (C3H3)
86% of conversion

C3H8

C3H6
H2
stoichiometry
1

1
1
Mass flowrate (lb/hr)
m1

m2=100,000.00
m3
MW (lb/lbmol)
44.10

42.08
2.01
n, (lbmole/hr)
2376.43

2376.43
2376.43

Number of moles of propylene formed
= (100,000 lb/hr)/42.08
= 2376.43 lbmole/hr C3H6

Assume 100% conversion, the mass flow rate of feed, m1 = 2376.43 x 44.1
= 104800.56 lb/hr
for 86% conversion, the mass flow rate of recycle, R = (1-0.86)/0.86 x 104800.56
R = 17060.53 lb/hr
Mass flowrate of H2 , m3
= (no. of mole) X (molecular weight)
= (2376.43) x (2.01)
= 4776.62 lb/hr

Mass flowrate feed to the reactor, F = m1 + R
= 104800.56 + 17060.53 = 121861.09lb/h
2.3 Step 3 – Eliminate Differences in Composition
Pt-SnPt-SnC3H6C3H840˚CC3H6C3H840˚CC3H633˚CC3H633˚CH2-137.1˚cH2-137.1˚cH2C3H8C3H6C4H8C4H10C5H12H2C3H8C3H6C4H8C4H10C5H120.99 C3H80.009 C4H100.001 C5H1247°C0.99 C3H80.009 C4H100.001 C5H1247°C
Pt-Sn
Pt-Sn
C3H6
C3H8
40˚C

C3H6
C3H8
40˚C

C3H6

33˚C
C3H6

33˚C
H2
-137.1˚c

H2
-137.1˚c

H2
C3H8
C3H6
C4H8
C4H10
C5H12

H2
C3H8
C3H6
C4H8
C4H10
C5H12

0.99 C3H8
0.009 C4H10
0.001 C5H12
47°C
0.99 C3H8
0.009 C4H10
0.001 C5H12
47°C
Reactor600˚C, 1.0 barReactor600˚C, 1.0 bar
Reactor
600˚C, 1.0 bar
Reactor
600˚C, 1.0 bar

17.5 bar17.5 bar10 bar10 bar
17.5 bar
17.5 bar
10 bar
10 bar
S4S415 bar15 barS3S3S2S215 bar15 barS1S152.°C52.°C
S4
S4
15 bar
15 bar
S3
S3
S2
S2
15 bar
15 bar
S1
S1
52.°C
52.°C
-137.1°C-137.1°CLPG:C3H8C4H10C5H12LPG:C3H8C4H10C5H1242˚C42˚C
-137.1°C

-137.1°C

LPG:
C3H8
C4H10
C5H12
LPG:
C3H8
C4H10
C5H12
42˚C
42˚C
-137.1˚CC3H8C3H6C4H8C4H10C5H12-137.1˚CC3H8C3H6C4H8C4H10C5H12
-137.1˚C
C3H8
C3H6
C4H8
C4H10
C5H12

-137.1˚C
C3H8
C3H6
C4H8
C4H10
C5H12

108.4˚CC4H8C4H10C5H12108.4˚CC4H8C4H10C5H1242˚CC3H842˚CC3H8110.7°CC4H10C5H12110.7°CC4H10C5H12
108.4˚C
C4H8
C4H10
C5H12
108.4˚C
C4H8
C4H10
C5H12
42˚C
C3H8
42˚C
C3H8
110.7°C
C4H10
C5H12

110.7°C
C4H10
C5H12

Figure 2.5: Flowsheet with separation units of propylene production process

In order to enable all chemicals involved to be supplied to their sinks, separation operations are needed. Figure 2.5 shows the separation units that are needed in a propylene production process. Since the raw material using in this process is from LPG that consists 80% propane, 18% butane and 2% pentane, so S1 as a separating unit is needed to separate propane from butane and pentane. However, the separation is not perfect. There will still have some butane and pentane that will be distillate but in a small proportion. As referred to table 2.4, S1 will be operated at 15bar. The bubble point at distillate product is 47˚C and the dew point of mixtures at the bottom product is 110.7˚C.
When the separation between propane, butane and pentane is done, propane as a reactant will enter the reactor which will be operate at 600˚C and 1 bar. These pressure and temperature is selected because the dehydrogenation process of propylene only will occur at these conditions.
After the reaction occurs, there have a lot of products produced. In order to separate the products, 3 separation units will be used. The first product that will be separated is hydrogen gas. The reason is, hydrogen gas has a low value of critical pressure and it will be difficult to separate the other products if the hydrogen maintain in the product mixtures. S2 will be used as separation unit that will be operated at pressure 10 bar and temperature -137.1 at dew point of vapor of the product mixture.
Next, after separate hydrogen gas, we will separate propane and propene from the side product. From Table 2.4 at 1 atm, the boiling point of C3 is very low, - 48˚C, and hence if C3 were recovered at 1 atm as the distillate of the S3, very costly refrigeration would be necessary to condense the reflux stream. At 18 bar , the bubble point of propane and propylene mixture is at 40˚C and much less cost refrigeration could be used. The bottom products which are consists butane, butene and pentane has a dew point 108.4˚C at 17.5 bar.
After separation unit S3 is inserted into the process design, S4 follows naturally. The distillate from S3 is separated into nearly pure species in the S4, which is specified at 15 bar. Under these conditions, the distillate (nearly pure propylene) boils at 33˚C and can be condensed with inexpensive cooling water, which is available at 25˚C. However, S4 need special separation unit due to small difference of boiling point between propane and propylene.

Table 2.4: Boiling points and critical constant
Chemical
Normal boiling point
(1atm, ˚C)
Boiling point (˚C)
Critical constant

15 bar
17.5 bar
20 bar
Tc (˚C)
PC (bar)
H2
-252.78
-
-
-
-240.01
12.96
C3H8
-42.11
41.00
45.00
53.55
96.74
42.51
C3H6
-47.62
33.00
35.85
42.65
91.06
45.55

2.4 Step 4 – Eliminate Differences in Temperature, Pressure, and Phase

Figure 2.6: Flowsheet with temperature-, pressure-, and phase-change operations in the propylene production process.
Figure 2.6 shows the changes of the state of chemicals. Since the original state of the raw material (LPG) is at 20°C and 18 bar, its temperature is raised to 52°C at 15 bar. The LPG is then introduced into a separation column (S1) at 15 bar with 99% conversion that separates the propane gas from other LPG products. Here, only 99% of LPG is converted to propane gas where another 1% is butane gas and pentane gas.
The process begins by mixing the upper products from S1 (propane gas, butane gas and pentane gas) with a stream of recycle propane gas at 47°C and 15 bar. The mixing of upper products from S1 and recycle propane undergoes the following operations:
The product mixture is preheated before it is introduced to the reactor. The reaction occurs at around 600oC and 1 bar.
The products mixture is then cooled to its dew point -137.1oC at 10 bar.
Then, the product mixture is introduced into a condenser (S2) that separates the hydrogen gas from other liquid products.
In addition, the liquid mixture that condensed at -137.1oC at 10 bar from the condenser is operated upon as follows:
Its pressure is increased to 17.5 bar.
The temperature is then raised to a liquid at its bubble point, 42oC at 17.5 bar.
Then, the liquid mixture is introduced into a separation column (S3) that separates the propane gas and propylene gas from other liquid products.
Next, the upper products (propane gas and propylene gas) from separation column (S3) are then entered into separation column (S4) at 40oC. The propylene gas with a boiling point of 33oC at 15 bar is come out as an upper product from separation column (S4).
Finally, the propane liquid from the recycle stream (at 42oC and 15 bar) undergoes the operation where its temperature is raised to the mixing temperature at 47 oC at 15 bar.

2.5 Step 5 – Task Integration
Figure 2.7 shows task integration for the propylene production process. At this stage in process synthesis, it is common to make the most obvious combinations of operations, leaving many possibilities to be considered when the flowsheet is sufficiently promising to undertake the preparation of a base case design. Below are the descriptions of unit process shown in Figure 2.7:
Heat exchanger
Heat exchanger is needed to increase or decrease the temperature of the stream. A heat exchanger is a piece of equipment built for efficient heat transfer from one medium to another. The media may be separated by a solid wall to prevent mixing or they may be in direct contact.

Depropanizer
A propane rich liquefied petroleum gas (LPG) feedstock is sent to a depropanizer to reject butanes and heavier hydrocarbons.

Furnace
Since the outlet temperature from the mixer is 47˚C and we need to increase the temperature to 600˚C, the furnace is used to heat up the stream. This follows heuristics 26 which explained near-optimal minimum temperature approaches in heat exchangers depend on the temperature level. For 250 to 350˚F, the stream must be heat up in a furnace for flue gas temperature above inlet process fluid temperature. An industrial furnace or direct fired heater is equipment used to provide heat for a process or can serve as reactor which provides heats of reaction. Furnace designs vary as to its function, heating duty, type of fuel and method of introducing combustion air.

Oleflex Reactor
The UOP Oleflex process is a catalytic dehydrogenation technology for the production of light olefins from their corresponding paraffins. One specific application of this technology produces propylene from propane. The Olexflex process uses a platinum catalyst to promote the dehydrogenation reaction
Pump
Since the pressure change operation involves a liquid, it is accomplished by a pump, which requires only 66 Bhp, assuming an 80% efficiency. The enthalpy change in the pump is very small and the temperature does not change by more than 1˚C

Distillation Column
To separate the mixture of C3 and butane, butane and pentane, distillation column is selected as the best separation unit. Distillation is based on the fact that the vapour of a boiling mixture will be richer in the components that have lower boiling points. Therefore, when this vapour is cooled and condensed, the condensate will contain more volatile components. At the same time, the original mixture will contain more of the less volatile material.

Propane-Propylene Splitter
C3 splitters are frequently designed with vapor-recompression heat pumps when sufficient low-energy heat sources are not available. The heat of vaporization of propylene and propane at 100psia are nearly identical. The only energy needed for a C3 splitter heat pump is the compressor duty, which is typically only 11-12% of the total reboiler duty. Therefore, the energy savings are significant. In addition, C3 splitter heat pump system operates at much lower pressure than conventional columns without heat pumping. The high-pressure compressor discharge stream is the same as the conventional tower's top pressure.

P-P SplitterP-P SplitterFigure 2.7: Flowsheet task integration for the propylene production process
P-P Splitter
P-P Splitter

CHAPTER 3

MATERIAL AND ENERGY BALANCES AND PROCESS SIMULATION

3.1 MATERIAL BALANCES
3.1.1 Overall Mass Balance
ReactorT = 500 OCP = 1 bar gfrgfr gdm2C3H6m3H2R lb/hr C3H8m1C3H8F lb/hr C3H8ReactorT = 500 OCP = 1 bar gfrgfr gdm2C3H6m3H2R lb/hr C3H8m1C3H8F lb/hr C3H8
Reactor
T = 500 OC
P = 1 bar
gfrgfr

gd

m2C3H6
m3H2
R lb/hr C3H8
m1C3H8
F lb/hr C3H8
Reactor
T = 500 OC
P = 1 bar
gfrgfr

gd

m2C3H6
m3H2
R lb/hr C3H8
m1C3H8
F lb/hr C3H8

Overall Reaction Equation :
C3H8 C3H6 + H2

Basis : 100000 lb/hr of propylene (C3H3)
86% of conversion

C3H8

C3H6
H2
stoichiometry
1

1
1
Mass flowrate (lb/hr)
m1

m2=100,000.00
m3
MW (lb/lbmol)
44.10

42.08
2.01
n, (lbmole/hr)
2376.43

2376.43
2376.43

S-10.990 C3H80.009 C4H100.001 C5H120.90 C4H100.10 C5H120.80 C3H80.18 C4H100.02 C5H12F1 lb/hrD1 lb/hrB1 lb/hr123S-10.990 C3H80.009 C4H100.001 C5H120.90 C4H100.10 C5H120.80 C3H80.18 C4H100.02 C5H12F1 lb/hrD1 lb/hrB1 lb/hr1233.1.2 Mass Balance for Separation Unit 1
S-1
0.990 C3H8
0.009 C4H10
0.001 C5H12

0.90 C4H10
0.10 C5H12

0.80 C3H8
0.18 C4H10
0.02 C5H12

F1 lb/hr
D1 lb/hr
B1 lb/hr
1

2

3

S-1
0.990 C3H8
0.009 C4H10
0.001 C5H12

0.90 C4H10
0.10 C5H12

0.80 C3H8
0.18 C4H10
0.02 C5H12

F1 lb/hr
D1 lb/hr
B1 lb/hr
1

2

3

No.
Component
Stream 1
Stream 2
Stream 3

Mole Fraction
Mass Flowrate, (lb/hr)
Mole Fraction
Mole Fraction
1
Propane
0.80
104800.56
0.990
104800.56
0
0
2
Butane
0.18
31076.58
0.009
1116.11
0.9
29960.47
3
Pentane
0.02
4286.49
0.001
346.38
0.1
3940.11

3.1.3 Mass Balance for Mixer
M-1a1 C3H8a2 C4H10a3 C5H120.990 C3H80.009 C4H100.001 C5H12D1 = 106263.1 lb/hrF = 123323.6 lb/hrR = 17060.53 lb/hr1.0 C3H83410M-1a1 C3H8a2 C4H10a3 C5H120.990 C3H80.009 C4H100.001 C5H12D1 = 106263.1 lb/hrF = 123323.6 lb/hrR = 17060.53 lb/hr1.0 C3H83410
M-1
a1 C3H8
a2 C4H10
a3 C5H12

0.990 C3H8
0.009 C4H10
0.001 C5H12

D1 = 106263.1 lb/hr
F = 123323.6 lb/hr
R = 17060.53 lb/hr
1.0 C3H8
3

4

10

M-1
a1 C3H8
a2 C4H10
a3 C5H12

0.990 C3H8
0.009 C4H10
0.001 C5H12

D1 = 106263.1 lb/hr
F = 123323.6 lb/hr
R = 17060.53 lb/hr
1.0 C3H8
3

4

10

No.
Component
Stream 3
Stream 4
Stream 10

Mole Fraction
Mole Fraction
Mole Fraction
1
Propane
0.990
104800.56
0.991
121861.09
1.0
17060.53
2
Butane
0.009
1116.11
0.007
1116.11
0
0
3
Pentane
0.001
346.38
0.002
346.38
0
0

3.1.4 Mass Balance for Reactor

5544R-1 121861.09 lb/hr C3H8 +1116.11 lb/hr C4H10346.38 lb/hr C5H120.991 C3H80.007 C4H100.002 C5H12C3H8C4H10C5H12C3H6C4H8H2123299.61 lb/hrR-1 121861.09 lb/hr C3H8 +1116.11 lb/hr C4H10346.38 lb/hr C5H120.991 C3H80.007 C4H100.002 C5H12C3H8C4H10C5H12C3H6C4H8H2123299.61 lb/hr
5

5

4

4

R-1
121861.09 lb/hr C3H8
+
1116.11 lb/hr C4H10
346.38 lb/hr C5H12
0.991 C3H8
0.007 C4H10
0.002 C5H12

C3H8
C4H10
C5H12
C3H6
C4H8
H2

123299.61 lb/hr
R-1
121861.09 lb/hr C3H8
+
1116.11 lb/hr C4H10
346.38 lb/hr C5H12
0.991 C3H8
0.007 C4H10
0.002 C5H12

C3H8
C4H10
C5H12
C3H6
C4H8
H2

123299.61 lb/hr

The percentage of conversion for propane and butane are 86% and 90% respectively and since the weight percent of pentane is too small, we assume that pentane is remain unreacted.

No.
Component
Stream 4
Stream 5

Mole Fraction
Mole Fraction
1
Propane
0.991
121861.09
0.0747
17060.53
2
Butane
0.007
1116.11
0.0004
111.61
3
Pentane
0.002
346.38
0.0009
346.38
4
Propene
0
0
0.4587
100000
5
Butene
0
0
0.0033
969.58
6
Hydrogen
0
0
0.4620
4811.36

3.1.5 Mass Balance for Separation Unit 2
567S-2C3H8C4H10C5H12C3H6C4H8H2123299.61 lb/hrH2C3H8C4H10C5H12C3H6C4H84811.36 lb/hr118488.25 lb/hr567S-2C3H8C4H10C5H12C3H6C4H8H2123299.61 lb/hrH2C3H8C4H10C5H12C3H6C4H84811.36 lb/hr118488.25 lb/hr
5

6

7

S-2
C3H8
C4H10
C5H12
C3H6
C4H8
H2

123299.61 lb/hr
H2

C3H8
C4H10
C5H12
C3H6
C4H8

4811.36 lb/hr
118488.25 lb/hr
5

6

7

S-2
C3H8
C4H10
C5H12
C3H6
C4H8
H2

123299.61 lb/hr
H2

C3H8
C4H10
C5H12
C3H6
C4H8

4811.36 lb/hr
118488.25 lb/hr

No.
Component
Stream 5
Stream 6
Stream 7

Mole Fraction
Mole Fraction
Mole Fraction
1
Propane
0.0747
17060.53
0
0
0.1388
17060.53
2
Butane
0.0004
111.61
0
0
0.0007
111.61
3
Pentane
0.0009
346.38
0
0
0.0017
346.38
4
Propene
0.4587
100000
0
0
0.8526
100000
5
Butene
0.0033
969.58
0
0
0.0062
969.58
6
Hydrogen
0.4620
4811.36
1
4811.36
-
-

S-3C3H8C4H10C5H12C3H6C4H8118488.25 lb/hrC3H8C3H6C4H10C5H12C4H81427.57 lb/hr117060.68 lb/hr789S-3C3H8C4H10C5H12C3H6C4H8118488.25 lb/hrC3H8C3H6C4H10C5H12C4H81427.57 lb/hr117060.68 lb/hr789
S-3
C3H8
C4H10
C5H12
C3H6
C4H8

118488.25 lb/hr
C3H8
C3H6

C4H10
C5H12
C4H8

1427.57 lb/hr
117060.68 lb/hr
7

8

9

S-3
C3H8
C4H10
C5H12
C3H6
C4H8

118488.25 lb/hr
C3H8
C3H6

C4H10
C5H12
C4H8

1427.57 lb/hr
117060.68 lb/hr
7

8

9

No.
Component
Stream 7
Stream 8
Stream 9

Mole Fraction
Mole Fraction
Mole Fraction
1
Propane
0.1388
17060.53
0.1400
17060.53
0
0
2
Butane
0.0007
111.61
0
0
0.08
111.61
3
Pentane
0.0017
346.38
0
0
0.20
346.38
4
Propene
0.8526
100000
0.8600
100000
0
0
5
Butene
0.0062
969.58
0
0
0.720
969.58


S-4117060.68 lb/hrC3H8C3H6C3H6C3H817060.53 lb/hr100000 lb/hr81110S-4117060.68 lb/hrC3H8C3H6C3H6C3H817060.53 lb/hr100000 lb/hr81110
S-4
117060.68 lb/hr
C3H8
C3H6

C3H6

C3H8

17060.53 lb/hr
100000 lb/hr
8

11

10

S-4
117060.68 lb/hr
C3H8
C3H6

C3H6

C3H8

17060.53 lb/hr
100000 lb/hr
8

11

10

No.
Component
Stream 8
Stream 1
Stream 11

Mole Fraction
Mole Fraction
Mole Fraction
1
Propane
0.140
17060.53
1
17060.56
0
0
2
Propene
0.860
100000
0
0
1
100000

ENERGY BALANCES

Table 3.1: Table of Data for Heat of Capacities
Cp=A+B*T+C*T^-2+D*T^-3
Compound
Molecular Weight
Hf
Hv
Kj/mol
Cp=A+B*T+C*T^-2+D*T^-3

Average Cp for liquid
KJ/(mol.K)

A*10^3
B*10^5
C*10^8
D*10^12

Propane
44.09
-103.8
18.77
68.023
22.59
-13.11
31.71
0.10584
Propene
42.08
20.41
18.42
59.58
17.71
-10.17
24.6
0.1199
Butane
58.12
-124.7
22.306
92.3
27.88
-15.47
34.98
0.13367
Butene
56.1
1.17
21.916
82.88
25.64
-17.27
50.50
0.09396
Pentane
72.15
-146.4
25.77
114.8
34.09
-18.99
42.26
0.167
Hydrogen
2.016
0
0.904
28.84
0.00765
0.3288
-0.8698
-

Energy Balance
We use heat of vaporization instead of liquid heat capacities to calculate the stream enthalpy and the value stated in Table 3.1. For a mixed stream, both equations are applied based on the vapor/liquid fraction involved. If there is no reaction occur in a unit (i.e. initial component = final component), enthalpy change for the unit is express as below:
(Vapor)(Vapor)
(Vapor)
(Vapor)

(Liquid)(Liquid)
(Liquid)
(Liquid)
Where, n is the total molar flow rate of that specific stream
For streams with composition or component change (i.e. reactor), heat of formation must be included.

3.2.1 Separation Unit 1
S-10.990 C3H80.009 C4H100.001 C5H120.90 C4H100.10 C5H120.80 C3H80.18 C4H100.02 C5H12F1 lb/hrD1 lb/hrB1 lb/hr123S-10.990 C3H80.009 C4H100.001 C5H120.90 C4H100.10 C5H120.80 C3H80.18 C4H100.02 C5H12F1 lb/hrD1 lb/hrB1 lb/hr123
S-1
0.990 C3H8
0.009 C4H10
0.001 C5H12

0.90 C4H10
0.10 C5H12

0.80 C3H8
0.18 C4H10
0.02 C5H12

F1 lb/hr
D1 lb/hr
B1 lb/hr
1

2

3

S-1
0.990 C3H8
0.009 C4H10
0.001 C5H12

0.90 C4H10
0.10 C5H12

0.80 C3H8
0.18 C4H10
0.02 C5H12

F1 lb/hr
D1 lb/hr
B1 lb/hr
1

2

3

For Stream 1
Liquid phase
Stream temperature, T = 325.13 K and consider datum at 298.13 K

Component
Flow rate
Ibmole/hr
Flow rate
mole/hr
Cp
H
KJ/hr
Propane
2376.43
1077948.64
0.10584
3080432.27
Butane
534.69
242535.38
0.13367
875332.014
Pentane
59.41
26948.37
0.167
121510.2

4077274.484 KJ/hr

For Stream 2
Liquid phase

Component
Flow rate
Ibmole/hr
Flow rate
mole/hr
Cp
liquid
H
Propane
2376.43
1077924.884
0.10584
2509926.534
Butane
19.203
8710.288
0.13367
25614.69
Pentane
4.8
2177.23
0.167
7999.14

2543540.36 KJ/hr

For Stream 3
/Stream temperature, T = 377.13 K and consider datum at 298.13 K
Component
Flow rate
Ibmole/hr
Flow rate
mole/hr
Cp
liquid
H
Butane
8.86
4018.8
0.13367
42438.24
Pentane
0.756
342.91
0.167
4524.011

46962.25 KJ/hr

H=46962.25 +2543540.36 -4077274.484 =-1486771.86 KJ/hr

3.2.2 Heat exchanger 1
Stream inlet 20C, liq.Stream inlet 52 C, liq.Stream inlet 20C, liq.Stream inlet 52 C, liq.
Stream inlet 20C, liq.
Stream inlet 52 C, liq.
Stream inlet 20C, liq.
Stream inlet 52 C, liq.

Component
Flow rate
Ibmole/hr
Flow rate
mole/hr
Cp
H
Propane
2376.96
1077948.64
0.10584
3650882.69
Butane
534.7
242535.38
0.13367
1037430.53
Pentane
59.41
26948.37
0.167
144012.08

4832325.3 KJ/hr

3.2.3 Furnace

Stream temperature, T = 600 C and consider datum at 47 ˚C

Component
Flow rate
Ibmole/hr
Flow rate
mole/hr
CpdT
H
KJ/hr
Propane
2376.43
1077948.64
69.62
75046784.32
Butane
534.69
242535.38
90.91
22048891.4
Pentane
59.41
26948.37
112.17
3022798.66

100118475 KJ/hr

3.2.4 Energy balance for Heat Exchanger 2 (HE2)

Stream inlet600˚C(Gas phase)Stream inlet600˚C(Gas phase)Stream outlet1˚C(Mixture phase)Stream outlet1˚C(Mixture phase)
Stream inlet
600˚C
(Gas phase)
Stream inlet
600˚C
(Gas phase)
Stream outlet
1˚C
(Mixture phase)
Stream outlet
1˚C
(Mixture phase)

Stream inlet at 600˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H

H2
2393.7
16.8053
40226.84661
C3H8
386.86
71.2935
27580.60341
C3H6
2376.4
59.5563
141529.5913
C4H10
1.9203
93.1652
178.9051336
C4H8
17.283
82.9306
1433.28956
C5H12
4.8009
114.9629
551.9253866

5180.9642
438.7138
211501.1614

Stream outlet at 1˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H
Hv

H2
2393.7
-0.6922
-1656.91914
-
C3H8
386.86
-2.5392
-982.314912
18.77
C3H6
2376.4
-2.8776
-6838.32864
18.42
C4H10
1.9203
-3.2082
-6.16070646
22.306
C4H8
17.283
-2.255
-38.973165
21.916
C5H12
4.8009
-4.008
-19.2420072
25.77

5180.9642
-15.5802
-9541.938571
107.182

Q = n( Hout - Hin – (- Hv))
=5180.9642(-15.5802–438.7138-(-107.182))
= -1798374.845 kJ/hr

Stream 61˚CStream 61˚C3.2.5 Energy Balance for Flash Separator (S2)
Stream 6
1˚C
Stream 6
1˚C
Flash separatorFlash separatorStream 51˚CStream 51˚CStream 71˚CStream 71˚C
Flash separator
Flash separator
Stream 5
1˚C
Stream 5
1˚C
Stream 7
1˚C
Stream 7
1˚C

Stream 5 (Feed Stream) at 1˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H

H2
2393.7
-0.6922
-1656.91914
C3H8
386.86
-2.5392
-982.314912
C3H6
2376.4
-2.8776
-6838.32864
C4H10
1.9203
-3.2082
-6.16070646
C4H8
17.283
-2.255
-38.973165
C5H12
4.8009
-4.008
-19.2420072

5180.9642
-15.5802
-9541.938571

Stream 6 (Distillate stream) at 1˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H

H2
2393.7
-0.6922
-1656.91914

Stream 7 (Bottom stream) at 1˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H

C3H8
386.86
-2.5392
-982.314912
C3H6
2376.4
-2.8776
-6838.32864
C4H10
1.9203
-3.2082
-6.16070646
C4H8
17.283
-2.255
-38.973165
C5H12
4.8009
-4.008
-19.2420072

2787.2642
-14.888
-7885.019431

There's no heat transfer from the flash column:
Q= ΔH=0
Q = n Hout - n Hin
=(-7885.019431+ (-1656.91914)) – (-9541.938571)
= 0 kJ/hr

3.2.6 Energy balance for heat exchanger 3

Stream inlet1˚CStream inlet1˚CStream outlet42˚CStream outlet42˚C
Stream inlet
1˚C
Stream inlet
1˚C
Stream outlet
42˚C
Stream outlet
42˚C

Stream inlet at 1˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H

C3H8
386.86
-2.5392
-982.314912
C3H6
2376.4
-2.8776
-6838.32864
C4H10
1.9203
-3.2082
-6.16070646
C4H8
17.283
-2.255
-38.973165
C5H12
4.8009
-4.008
-19.2420072

2787.2642
-14.888
-7885.019431

Stream outlet at 42˚C datum at 25˚C
Compound
Flowrate
(mol/hr)
CpdT
(kJ/mol)
n H

C3H8
386.86
1.2825
496.14795
C3H6
2376.4
1.1118
2642.08152
C4H10
1.9203
1.7249
3.31232547
C4H8
17.283
1.5516
26.8163028
C5H12
4.8009
2.1421
10.28400789

2787.2642
7.8129
3178.642106

Q = n Hout - n Hin
= 3178.642106– (-7885.019431)
= 63273.40588 kJ/hr

3.2.7 Energy Balance for Compressor (C1)
STREAM OUTLET10 bar600˚CSTREAM OUTLET10 bar600˚CSTREAM INLET1 bar600˚CSTREAM INLET1 bar600˚C
STREAM OUTLET
10 bar
600˚C
STREAM OUTLET
10 bar
600˚C
STREAM INLET
1 bar
600˚C
STREAM INLET
1 bar
600˚C

Inlet (stream 2)
Phase
Vapor
Component
Mixture
Pressure (bar)
1
Temperature (oC)
600
Total Molar Flow Rate (kmol/hr)
2349.64

Outlet (stream 3)
Phase
Vapor
Component
Mixture
Pressure (bar)
10
Temperature (oC)
600
Total Molar Flow Rate (kmol/hr)
2349.64

The outlet temperature of a stream by assuming the process is an ideal system.
T2=T1P2P1γ-1γ
For ideal system, γ = 1.3.
Hence,
T2=6001011.3-11.3
=1020.75˚C
For energy balance,
Q= H=outniHi-inniHi
=outni298.15TCpdT-inni298.15TCpdT

Since there is no change component flow rate,
Q= H=ni298.151293.968.023×10-3+22.59×10-5T-13.11×10-8T2+31.71×10-12T3dT+298.151293.959.58×10-3+17.71×10-5T-10.17×10-8T2+24.6×10-12T3dT+298.151293.992.3×10-3+27.88×10-5T-15.47×10-8T2+34.98×10-12T3dT+298.151293.982.88×10-3+25.64×10-5T-17.27×10-8T2+50.50×10-12T3dT+298.151293.9114.8×10-3+34.09×10-5T-18.99×10-8T2+42.26×10-12T3dT+298.151293.928.84×10-3+0.00765×10-5T+0.3288×10-8T2-0.8698×10-12T3dT
=2349.64 (175.44+144.36+227+197.87+278.61+30.52)
= 2476050.63 kJ/hr

3.2.8 Separation Unit 3 (S3)
S-3C3H8C4H10C5H12C3H6C4H8118488.25 lb/hrC3H8C3H6C4H10C5H12C4H81427.58 lb/hr117060.68 lb/hr789S-3C3H8C4H10C5H12C3H6C4H8118488.25 lb/hrC3H8C3H6C4H10C5H12C4H81427.58 lb/hr117060.68 lb/hr789
S-3
C3H8
C4H10
C5H12
C3H6
C4H8

118488.25 lb/hr
C3H8
C3H6

C4H10
C5H12
C4H8

1427.58 lb/hr
117060.68 lb/hr
7

8

9

S-3
C3H8
C4H10
C5H12
C3H6
C4H8

118488.25 lb/hr
C3H8
C3H6

C4H10
C5H12
C4H8

1427.58 lb/hr
117060.68 lb/hr
7

8

9

For Stream 7
Liquid stream
Component
Flow rate
(lbmol/hr)
Flow rate
(mol/hr)
Cp
CpdT
(kJ/mol)
n H
kJ/hr
Propane
386.86
1.75447 х 105
0.10584
1.7993
315681.79
Butane
1.92
870.748
0.13367
2.2724
1978.69
Pentane
4.80
2176.871
0.167
2.8390
6180.14
Propene
2376.43
10.777 x 105
0.1199
2.0383
2196675.91
Butene
17.28
7836.735
0.09396
1.5973
12517.62

2533034.15

For Stream 8
Gas stream
Component
Flow rate
(lbmol/hr)
Flow rate
(mol/hr)
Cp
CpdT
(kJ/mol)
n H
kJ/hr
Propane
386.86
1.75447 х 105
0.10584
2.0110
352823.92
Propene
2376.4
1077732.426
0.1199
2.2781
2.45518 х 106

2.8080 x 106

For Stream 9
Liquid stream
Component
Flow rate
(lbmol/hr)
Flow rate
(mol/hr)
Cp
CpdT
(kJ/mol)
n H
kJ/hr
Butane
1.9203
870.884
0.13367
11.1481
9708.70
Butene
17.283
7838.095
0.09396
7.8363
61421.66
Pentane
4.80
2176.871
0.1670
13.9278
30319.02

101449.38

H = 101449.38+ 2.8080 x 106 - 2533034.15 = 376415.23kJ/hr

3.2.9 Separation Unit 4 (S4)
S4C3H6C3H8C3H8C3H6F1 lb/hrD1 lb/hrB1 lb/hr81112S4C3H6C3H8C3H8C3H6F1 lb/hrD1 lb/hrB1 lb/hr81112
S4
C3H6

C3H8

C3H8
C3H6

F1 lb/hr
D1 lb/hr
B1 lb/hr
8

11

12

S4
C3H6

C3H8

C3H8
C3H6

F1 lb/hr
D1 lb/hr
B1 lb/hr
8

11

12

For Stream 8
Liquid stream

Component
Flow rate
(lbmol/hr)
Flow rate
(mol/hr)
Cp
CpdT
(kJ/mol)
n H
kJ/hr
Propane
386.95
1.75517 х 105
0.10584
4.7628
8.35952 х 105
Propene
2376.43
10.77929 х 105
0.1199
5.3955
58.15965 х 105

66.51917 х 105

For Stream 11
Gas stream

Component
Flow rate
lbmol/hr
Flow rate
(mol/hr)
CpdT
(kJ/mol)
n H
(kJ/hr)
Propene
2376.43
10.77929 х 105
2.258
24.33963 х 105

24.33963 х 105

For Stream 12
Liquid stream
Component
Flow rate
(lbmole/hr)
Flow rate
(mol/hr)
Cp
CpdT
(kJ/mol)
n H
(kJ/hr)
Propane
386.95
1.75517 х 105
0.10584
4.97448
8.73105 х 105

8.73105 х 105

Q = H = 8.73105 х 105 + 24.33963 х 105- 66.51917 х 105= -33.44849 х 105 kJ/hr

3.2.10 Pump (P1)

outletoutlet
outlet
outlet
inletinlet
inlet
inlet

Component
Molar Flow Rate (lbmol/hr)
Molar Flow Rate, F (kmol/hr)
Molar volume,v(kmol/m3)
Fv
C3H8
C3H6
C4H10
C4H8
C5H12
386.86
2376
1.92
17.28
4.801
1.75476 х 102
10.77734 х 102
0.00870 х 102
0.07838 х 102
0.02178 х 102
21.9375
23.2486
0.09667
0.09048
0.11452
3849.50
25055.81
0.08410
0.70918
0.24942
Fv = 28906.35

Inlet Stream
Phase
Liquid
Pressure (bar)
1
Temperature (oC)
-47

Outlet Stream
Phase
Liquid
Pressure (bar)
18
Temperature (oC)
-47

Q = Fv (P)
Q = 28906.35 х (18 -1) = 4.91408 х 105 kJ/hr

3.2.11 Energy Balance for Reactor

873K (600˚C), 1 barR-1 121861.09 lb/hr C3H8 +1116.11 lb/hr C4H10346.38 lb/hr C5H120.991 C3H80.007 C4H100.002 C5H12C3H8C4H10C5H12C3H6C4H8H245873K (600˚C), 1 bar873K (600˚C), 1 barR-1 121861.09 lb/hr C3H8 +1116.11 lb/hr C4H10346.38 lb/hr C5H120.991 C3H80.007 C4H100.002 C5H12C3H8C4H10C5H12C3H6C4H8H245873K (600˚C), 1 bar
873K (600˚C), 1 bar
R-1
121861.09 lb/hr C3H8
+
1116.11 lb/hr C4H10
346.38 lb/hr C5H12
0.991 C3H8
0.007 C4H10
0.002 C5H12

C3H8
C4H10
C5H12
C3H6
C4H8
H2

4

5

873K (600˚C), 1 bar
873K (600˚C), 1 bar
R-1
121861.09 lb/hr C3H8
+
1116.11 lb/hr C4H10
346.38 lb/hr C5H12
0.991 C3H8
0.007 C4H10
0.002 C5H12

C3H8
C4H10
C5H12
C3H6
C4H8
H2

4

5

873K (600˚C), 1 bar

1. C3H8 C3H6 + H2 Hr1
2. C4H10 C4H8 + H2 Hr2
C3H8 (g), 873K (600˚C), 1 barC3H8 (g), 298K (25˚C), 1 barC3H6 (g), 298K (25˚C), 1 barC3H6 (g), 873K (600˚C), 1 bar>>>HRHP Hr1C3H8 (g), 873K (600˚C), 1 barC3H8 (g), 298K (25˚C), 1 barC3H6 (g), 298K (25˚C), 1 barC3H6 (g), 873K (600˚C), 1 bar>>>HRHP Hr1
C3H8 (g), 873K (600˚C), 1 bar
C3H8 (g), 298K (25˚C), 1 bar
C3H6 (g), 298K (25˚C), 1 bar
C3H6 (g), 873K (600˚C), 1 bar
>
>
>
HR
HP
Hr1
C3H8 (g), 873K (600˚C), 1 bar
C3H8 (g), 298K (25˚C), 1 bar
C3H6 (g), 298K (25˚C), 1 bar
C3H6 (g), 873K (600˚C), 1 bar
>
>
>
HR
HP
Hr1

Reaction path for PropaneReaction path for Propane
Reaction path for Propane
Reaction path for Propane
Q1= nΔH=nHR1+nHp1+ Hr1 (for Propane)
Q2= nΔH=n nHR2+nHp2+ Hr2 (for Butane)
Q= Q1+ Q2

1. Energy balance for Propane
Component
HR1
Hp1
Hr1
Flow rate
(mol/hr)
1253192.74
2330911.57
-
Specific Enthalpy, H (kJ/mol)
-91.8
184.72
-
H
(kJ/hr)
-115043093.5
430565985.2
140.21

315523031.9

2. Energy balance for Butane
Component
HR2
Hp2
Hr2
Flow rate
(mol/hr)
8708.84
16547.07
-
Specific Enthalpy, H (kJ/mol)
-119.0
240.61
-
H
(kJ/hr)
-1036351.96
3981390.5
125.87

1945164.41

Q= Q1+ Q2
Q= 315523031.9kJhr+ 1945164.41kJhr
Q = 317.468 x 106 kJ/hr

3.3 Simulation Result from ASPEN HYSYS
3.3.1 Material Balance
Stream no.
Mass
(lbmole/hr)
(Hysys)
stream 1
3179
stream 2
3179
stream 3
769.4
stream 4
2410
stream 5
2523
stream 6
2523
stream 8
4692
stream 9
4692
stream 10
4692
stream 11
4692
stream 12
2533
stream 13
2159
stream 14
2533
stream 15
2533
stream 16
156.3
stream 17
2377
stream 23
2263
stream 24
113.4
stream 25
113.2

3.3.2 Energy Balance
Equipment
Energy, kJ/h (HYSYS)
Q-HE00
5.76E+06
Q-HE01
1.79E+08
Q-HE02
1.65E+07
Q-Furnace
9.23E+07
Q-Compressor
6.78E+07
Q-Pump
6.90E+04

3.4 Percentage Difference between Manual Calculation and HYSYS Calculation
3.4.1 Mass Balance
Stream no.
Mass (lbmole/hr)
(manual)
Mass
(lbmole/hr)
(Hysys)
% Diff
stream 1
2971
3179
6.54
stream 2
2971
3179
6.54
stream 3
721.5
769.4
6.23
stream 4
2400.4
2410
0.40
stream 5
2787.3
2523
10.48
stream 6
2787.3
2523
10.48
stream 8
5181
4692
10.42
stream 9
5181
4692
10.42
stream 10
5181
4692
10.42
stream 11
5181
4692
10.42
stream 12
2787
2533
10.03
stream 13
2393
2159
10.84
stream 14
2787
2533
10.03
stream 15
2787
2533
10.03
stream 16
165.4
156.3
5.82
stream 17
2663.5
2377
12.05
stream 23
2376.4
2263
5.01
stream 24
124.5
113.4
9.79
stream 25
124.5
113.2
9.98

3.4.2 Energy Balance
Equipment
Energy, kJ/h (manual)
Energy, kJ/h (HYSYS)
% Diff
Q-HE00
4.83E+06
5.76E+06
16.09
Q-HE01
1.80E+06
1.79E+08
99.00
Q-HE02
6.33E+04
1.65E+07
99.62
Q-Furnace
1.00E+08
9.23E+07
8.46
Q-Compressor
2.48E+06
6.78E+07
96.35
Q-Pump
4.91E+05
6.90E+04
612.70

CHAPTER 4

HEAT INTEGRATION

4.1 PROCESS ENERGY INTEGRATION

Tmin = 10˚C

Table 4.1: Steam Table Data
Stream
Type
Tsupply (˚C)
Ttarget (˚C)
FCp (MW/K)
C1
Cold
20
50.08
0.053
C2
Cold
43.85
576.30
0.048
H1
Hot
870.7
-137.1
0.049
C3
Cold
-136.8
30
0.028

4.2 Algorithm Table4.2 Algorithm TableT (˚C)
4.2 Algorithm Table
4.2 Algorithm Table

0.049

T (˚C)
FCpC - FCpH
(MW/K)
Hi (MW)
1st Cascade
865.70

H1

PinchhPinchh0
Pinchh
Pinchh

284.4
-0.049
-13.9356

581.30

13.94

526.22
-0.001
-0.52622

55.08

14.46

6.23
0.052
0.32396

48.85

C2

14.14

0.048

13.85
0.004
0.0554

35.00

14.08

10
0.032
0.3200

25.00
C1

13.76

0.053

156.8
-0.021
-3.2928

-131.80

C3

17.06

0.028
10.3
-0.049
-0.5047

-142.10

Qc minQc min17.56
Qc min
Qc min

Figure 4.1: Algorithm Table
4.3 Heat Exchanger Network
T pinch (870.7˚C)T pinch (870.7˚C)
T pinch
(870.7˚C)
T pinch
(870.7˚C)
E11E22E3CE11E22E3CC1C1 H (MW) FCp(MW/K)
E11
E22
E3
C
E11
E22
E3
C
C1
C1
1.5942 0.053
C2C2 50.08 20
C2
C2
25.5576 0.048
C3C3 576.3 43.85
C3
C3
4.6704 0.028
H1H1 30 -136.8
H1
H1
-137.1-137.149.3822 0.049
-137.1
-137.1
870.7 25.5576 1.5942 4.6704 17.56

Figure 4.2: Heat ExchangerNetwork

Table 4.2: Summary of Temperature of Heat Exchanger

TH,in (˚C)
TH,out(˚C)
TC,in (˚C)
TC,out (˚C)
E1
870.7
349.12
43.85
576.3
E2
349.12
316.59
20
50.08
E3
316.59
221.28
-136.8
30
C
221.28
-137.1
-
-

4.4 Process Flow Diagram Heat Exchanger Network4.4 Process Flow Diagram Heat Exchanger NetworkFigure 4.3: Process Flow Diagram Heat Exchanger Network
4.4 Process Flow Diagram Heat Exchanger Network
4.4 Process Flow Diagram Heat Exchanger Network
CHAPTER 5

OPTIMIZATION

Optimization is the tool to maximize our profit by minimizing the supply of raw material and maximizing the product. In this case, our target that we want to maximize it the production of propene (100000 Ib/hr) and our supply that we want to minimize it is the propane which is initially set to 104800 Ib/hr depending on the stoichiometric coefficient of (propane/propene =1.048) and (Hydrogen/propene= 0.0457).

Propane

Propene
+
Hydrogen
Ibmol
1

1

1
MW
44.1

42.08

2.16
Ib
44.1

42.08

2.16
Ib/Ib propene
1.048

1

0.0457
USD RM/Ib
0.79

1.96

11.3

Step 1:
Define decision variables:
P1= amount of product (Propene)
P2=amount of byproduct (Hydrogen)
R=amount of reactant (Propane)
Z=maximum profit

Step 2:
Define objective function
Maximum profit (Z) = (1.96*P1 +11.3*P2)-(0.79*R)

Step 3:
Defining equality and inequality constraints:
Inequality constraints
Propane supply R << 104800 Ib/hr
Propene production P1>> 100000 Ib/hr

Equality constraints
R= 1.048*P1
P2=0.0457*P1

Non-negativity constraint
R, P1, P2 0

Step 4
Optimization technique
We used solver add-in in Microsoft excel:
P1=100000 Ib/hr
P2=4789.36 Ib/hr
R=104800 Ib/hr
Z=167327.768 RM/hr
After optimization the maximum profit is close to the manually calculated one= RM 167327.768/hr.

CHAPTER 6

EQUIPMENT SIZING AND COSTING

6.1 Introduction

In this chapter, the equipment sizing is done to all equipment that is involved in the proposed propylene production plant. Equipment sizing is a very important aspect of process design as it enables the subsequent analysis that is involved in process design such as mechanical design and economy analysis. The sizing involves the reactors, distillation column, compressor, pump, and heat exchangers.

6.2 Reactor
6.2.1 Sizing of Reactor
Parameter
SI
Volumetric Flowrate , Q
3517.02 ft3/hr
Retention time (half-full), t
5 min
Reactor Volume, V
586.17 ft3
Vessel Inside Diameter, Di
7.20 ft
Vessel Length, L
14.4 ft
Design Type
Vertical
Material of Contruction
Low- Alloy Steel SA-387B

6.2.2 Costing of Reactor
Cost of vessel, Cv = $ 40, 279
Cost of ladders and nozzles, CPL = $ 10, 264
Cost of purchase CP = $ 58, 599
Total cost with bare-module = 4.16 (58, 599) = $ 243, 772

6.3 Pump
6.3.1 Sizing of Pump
Pressure inlet, P1 = 1000kPa = 145.04psi
Pressure outlet, P2 = 1750kPa = 253.82psi
Pressure drop, ΔP = 750kPa = 108.78psi
Volumetric flow rate, Q = 93.57 m3/hr = 413.09 gpm
Pump head, H = ΔP (2.31)SG = ΔPρ = 356.82 ft

6.3.2 Costing of Pump
Cost of pump, CP = $ 6577.78
Cost of motor, CP = $ 4689.90
Total cost with bare-module = (6577.78 + 4689.90) (3.30)
= $ 37,183.34

6.4 Distillation Column
6.4.1 Sizing and costing of the main vessel:

Parameters
S1
Domed head wall thickness, a
13.7 mm
Tray spacing, b
2ft
Column diameter, c
6.05 ft
Column wall thickness, d
0.5 in
Design type
Vertical
Material of Construction
Carbon steel
Material of insulation
Mineral wool,60mm
Column type
Plate column
Plate type
Sieve
Domed head type
Torispherical
Costing in $:

Cost of vessel. Cv= $ 67436
Cost of ladders and nozzles, CPL= $ 21642
Cost of plates, CT= $ 27462

Total cost with bare-module
=4.16 (67436+21642+27462) = $ 484809

6.4.2 Sizing and costing of the reflux drum:

Parameters
S1
Domed head wall thickness,
13.7 mm
Vessel length,L
8.924 ft
vessel diameter, D
17.85 ft
Column wall thickness, d
0.562 in
Design type
Vertical
Carbon steel
Material of insulation
Mineral wool,60mm
After bare-model:
Cost= $ 223290

6.4.3 Sizing and costing of the condenser:

Parameters
S1
Length of tube
20 ft
Area of transfer,Ac
273.1 ft2
Carbon steel
Type of HE
fixed head, shell tube exchanger
Cp= $ 21721
After bare-module,
Cost= $ 68857

6.4.4 Sizing and costing of the re-boiler:

Parameters
S1
Length of tube

20 ft
Area of transfer,AR
3.146 ft2
Carbon steel
Type of HE
kettle reboiler
CB= $ 65325
With bare-module
$ 3.17(65325) = $ 207080

Total cost for S1:

Vessel or Equipment
Cost in $
main vessel
484809
the reflux drum
223290
the condenser
68857
the re-boiler
207080
Total
984036

6.5 Compresssor
6.5.1 Main Sizing Parameters

Parameters
Compressor
Compressor Type
Centrifugal
Drive Type
Steam turbine
Stainless steel
Inlet Volumetric Flow Rate, QI
83283.83 ft3/min
Inlet Pressure, PI
14.5 psi
Outlet Pressure, PO
72.52 psi
Specific Heat Ratio, k
1.10

6.5.2 Costing in $:

Purchase cost of compressor = $ 7,328,904

6.6 Heat Exchanger
6.6.1 Sizing of Heat exchanger (HE 2)

Heat exchanger type

2 shell and 4 tubes
Design type

Fixed Head
Heat exchanger orientation
Horizontal
Tube inlet direction

Horizontal
Heat duty (kJ/s)

1594.2
Heat duty (Btu/hr)

5.44x10^6

Hot
Cold
Tin (˚C)
870.7
43.85
Tout (˚C)
349.12
576.3

From Figure 18.15 (a), FT = 0.85 and 2-4 exchanger is used.

Ui = 235.5 Btu/oF.ft2.hr

Velocity of tube-side;

Cross section are/pass;

By using 0.75 in. O.D. 16 BWG tubing with I.D. of 0.62 in.;

Inside area/tube =
= 2.097x10-3 ft2/tube

Area per tube;

=
= 0.288 ft2/tube

L = 5.58 ft

6.6.2 Costing of Heat Exchanger (HE2)

FBM = 3.17

FM=1.08+(86.34100)0.5
= 2.01

FL = 1.25 (Tube length = 5.58 ft2)

FP=0.9803+0.018 145.04100+0.0017 (145.04100)2

= 1.01

Fixed head:

CB=exp11.0545-0.9228ln86.34+0.09861ln86.342
= $7,334.88

CP=2.011.251.01(7,334.88)
= $18,613.18

Bare-module cost = 3.17 ( 18,613.18)
= $59,003

CHAPTER 7

TOTAL CAPITAL INVESTMENT AND PAYBACK PERIOD

Total Capital Investment

By using method 3, which is based on the individual factors method of Guthrie, 1969, 1974 there are few steps to find the total capital investments, CTCI.

Firstly, we need to prepare an equipment list, giving the equipment tittle, label, size, material of construction, design temperature, and design pressure.

Equipment Tittle
Label
Size
Design Temp.
(˚C)
Design Pressure (bar)
Bare-module Cost, CBM
Reactor
R1
V=586.17 ft3
Di = 7.20 ft
L= 14.4 ft
Low- Alloy Steel SA-387B
576
1
$ 243, 772
Pump
P1
H =356.82 ft
Cast Steel
-137.1
Pinlet = 10
Poutlet = 17.5
$ 37,183
Distillation Column
S1
D = 6.05 ft
t = 0.5 in

Carbon steel
50
15
$ 984,036
Compressor
C1
Q = 83283.83 ft3/min
Carbon Steel
526
Pinlet = 1
Poutlet = 10
$ 15,757,144
Heat Exchanger
HE2
A = 86.34 ft2
Carbon Steel
870.7
10
$ 59,003
CTBM
$17,081,138

After we get the value of total bare module cost, CTBM, we need to find the site development cost, Csite, building cost, Cbuildings, and offsite facilities cost, Coffsite facilities by assuming some factor. The calculation of total capital investment cost is shown below:

Assume it is grass-roots plant, the value fo CSITE is 10-20% of CTBM. Assume we take 15% of CTBM.
CSITE = 0.15 (17,081,138)
CSITE = $ 2,562,170.75

Assume it is process buildings, the value of CBUILDINGS is 10% of CTBM
CBUILDINGS = 0.10 (17,081,138)
CBUILDINGS = $ 1,708,113.80

The value of COFFSITE FACILITIES is 5% of CTBM
COFFSITE FACILITIES = 0.05 (17,081,138)
COFFSITE FACILITIES = $ 854,056.90

Use factor of 1.18 to cover a contingency and a contractor fee
CTPI = 1.18 ( CTBM + CSITE + CBUILDINGS + COFFSITE FACILITIES)
CTPI = 1.18 (17,081,138+ 2,562,170.75 + 1,708,113.80 + 854,056.90)
CTPI = $ 39,969863.01

The value of CWC can be estimated 17.6% of CTPI
CWC = 0.176 (39,969863.01)
CWC = $ 7,034,695.89

Thus,
CTCI = CTPI + CWC
CTCI = $ 39,969,863.01+ $ 7,034,695.89
CTCI = $ 47,004,558.90

Payback Period

Payback period is the time in which the initial cash outflow of an investment is expected to be recovered from the cash inflows generated by the investment. It is one of the simplest investment appraisal techniques.

The formula to calculate payback period of a project depends on whether the cash flow per period from the project is even or uneven. In case they are even, the formula to calculate payback period is:

Payback Period =
Initial Investment

Cash Inflow per Period
= RM 152,764,816.40RM 167,327.77/hr
=912.967 hr ×1 day24 hr ×1 month30 days
= 1 month 9 days

CONCLUSION

Propylene is one of the highest volume of chemicals produced globally and primarily used as an intermediate for the production of other chemical raw materials. These chemical raw materials are then subsequently used to manufacture a large variety of substances and products. Example of such product is propylene, a widely used plastic where the manufacturing process consumes more than half of the world's production of polypropylene. There are other uses as well, such as manufacture of acrylonitrile, oxo process chemicals, cumene, isopropanol, polygas chemicals, and propylene oxide. This shows that the production of propylene has its demand in the global industry, hence a good marketability, especially in recent years where the price of propylene in the market is expected to continue rising as the demand increases for the chemical material. Market research report says worldwide polypropylene capacity increased at a Compound Annual Growth Rate (CAGR) of 5.2% from 2003, reaching 65 million tons per year (MMTY) in 2013, and is expected to continue rising to 86 MMTY by 2018, at a slightly higher CAGR of 5.8%. As Malaysia is a part of the global market, it can be expected that prices in Malaysia to be affected by the global prices.

In terms of reaction pathways for this particular project, a screening process was done based on gross profit, economic potential as well as other factors related such as energy consumption, toxicity, safety and environmental impacts. There are two reaction pathways suggested for the production of propylene, which are dehydrogenation of propane, and metathesis reaction of ethylene and butene. From the screening process, it was shown that dehydrogenation of propane reaction is a better process compared to the metathesis reaction. Based on the gross profit calculation, a dehydrogenation process would bring in a gross profit of RM 1.67/lb propylene with 86% conversion compared to only RM 0.46/lb propylene for metathesis reaction with a 90% conversion yield. Since the calculation was based on gross profit, further analysis need to be done in order to optimize the production process of propylene via the dehydrogenation of propane process for a sustainable plant design.

In addition to the reaction pathways and process screening, a process synthesis for the production of propylene from dehydrogenation of propane was done by following the steps that was introduced by Rudd, Powers, and Siirola. From these steps, a general overview of the whole process, starting from the raw materials into products is translated into a process flow diagram, as well as the operating parameters were obtained. This is an important step in designing the production process of our desired product before performing a further optimization of the processes and unit operations involved.
In a nutshell, after we had done a simulation, optimization and process integration, our total capital investment is $ 47,004,558.90.

APPENDICES A

CALCULATION OF MATERIAL BALANCES
Sample Calculation for Mass Balance

1. Overall mass balance
Number of moles of propylene formed
= (100,000 lb/hr)/42.08
= 2376.43 lbmole/hr C3H6

Assume 100% conversion, the mass flowrate of feed, m1 = 2376.43 x 44.1
= 104800.56 lb/hr

for 86% conversion, the mass flowrate of recycle, R = (1-0.86)/0.86 x 104800.56
R = 17060.53 lb/hr
Mass flowrate of H2 , m3
= (no. of mole) X (molecular weight)
= (2376.43) x (2.01)
= 4776.62 lb/hr

Mass flowrate feed to the reactor, F = m1 + R
= 104800.56 + 17060.53 = 121861.09lb/hr

2. Separation Unit 1
Overall mass balance :
F1 = D1 + B1
From the overall mass balance, we know that the mass flowrate of propane at D1 is 104800.56 lb/hr and the composition is assume 0.990 of C3H8, so
0.990 D1 = mass flow rate of C3H8
D1 = mass flow rate of C3H8 / 0.990
= 104800.56 / 0.990
= 106263.1 lb/hr distillate

Propane balance :
0.8 F1 = 0.990 D1
F1 = (0.990 x 106263.1)/0.8
F1 = 140163.6 lb/hr feed

F1 = D1 + B1
B1 = 140163.6 – 106263.1
B1 = 33900.6 lb/hr of bottom product

3. Mixer
Overall mass balance :
F = D1 + R where R is the recycle of propane

from previous calculation, the value of D1 = 106263.1 lb/hr and R = 17060.53 lb/hr. Hence,
F = 106263.1 lb/hr + 17060.53 lb/hr = 123323.6 lb/hr
Propane balance : 0.990 x (106263.1) + 17060.53 x (1.0) = 123323.6 x (a1)
a1 = 0.991
Butane balance : 0.009 x (106263.1) = 123323.6 x (a2)
a2 = 0.007
and the weight percent of pentane
a3 = 1 – 0.991 – 0.007 = 0.002

4. Reactor
For dehydrogenation of propane, 0.86% of propane is converted. The unreacted propane recycled.
(a) Mass flowrate of propane recycle
R = (1-0.86)/0.86 x 104800.56
R = 17060.53 lb/hr
(b) no of mol of propane recycle
= 17060.53/ 44.1 = 386.9 lbmole/hr
(c) no of mole of propene produce
= (121861.09 / 44.1) x 0.86 = 2376.64 lbmole/hr
(d) no of mole of butene produce
= 1116.106 / 58.12 x 0.9 = 17.28 lbmole/hr
(e) no of mole of hydrogen produce
= no of mole of hydrogen from propane + no of mole of hydrogen from propane
= 2376.4 + 17.28 = 2393.7 lbmole/hr
(f) mass flowrate of hydrogen produce
= 2393.7 x 2.01 = 4811.36 lb/hr

APPENDICES B

CALCULATION OF ENERGY BALANCE
Sample Calculation for Energy Balance
Molar Flow Rate for Propane
n=2763.3lbmolhr x1kgmol2.205lbmolx1000mol1kgmol = 1253197.28mol/hr

For reaction of propane
HR1=873K298KCp(reactant)dt
HR1=873K298K68.023×10-3+22.59×10-5T-13.11×10-8T2+31.71×10-12T3 dt
HR1=-91.81kJ/mol
HR1=873K298KCp(product)dt
Hp1=298K873K68.023×10-3+22.59×10-5T-13.11×10-8T2+31.71×10-12T3 dt+59.58×10-3+17.71×10-5T-10.17×10-8T2+24.6×10-12T3 dt+ 28.84×10-3+0.00765×10-5T+0.3288×10-8T2-0.8698×10-12T3 dt
Hp1=91.79kJ/mol+75.75kJ/mol +17.18kJ/mol
Hp1= 184.72kJ/mol
Hr1=vi Hf=1 HfC3H6+1 HfH2- 1 Hf C3H8
Hr1=20.41 kJmol+0 - -119.8kJmol=140.21 kJ/mol

APPENDICES C

ASPEN HYSYS

APPENDICES D

CALCULATION OF HEAT INTEGRATION

Calculation for temperature of heat exchanger.
E1
Q=FCpH T
25.5576 = 0.049 (870.7-T)
T = 349.12˚C

Q=FCpC T
25.5576 = 0.048 (T – 43.85)
T = 576.3˚

E2
Q=FCpH T
1.5942 = 0.049 (349.12-T)
T = 316.59˚C

Q=FCpC T
1.5942 = 0.053 (T-20)
T = 50.08˚C

E3
Q=FCpH T
4.6704 = 0.049 (316.59-T)
T = 221.28˚C

Q=FCpC T
4.6704 = 0.028 (T—136.8)
T = 30˚C

APPENDICES E

CALCULATION OF SIZING AND COSTING

REACTOR
Q = 3517.02 ft3/hr
Retention time =5 min at half full :
Volume, V = (3517.02 ft3/hr) × (5min×1 hr60 min ×2) = 586.17 ft3
Assume L/ D = 2
V = π (D/2)2L = (πD3)/2
D = (2V/ π)1/3 = [2(586.17)/ π] 1/3 = 7.20 ft
L= 2D = 14.4 ft

Operating Pressure = 1 bar = 14.5 psig :
Pd = exp { 0.60608 = 0.91615 [ln(14.5)] + 0.0015655 [ln(14.5)]2} = 21.48 psig (eqn. 22.61)
S = 10993.86 psi (low – alloy)
E = 1.0
tP = 21.48 × 7.2 ×12 2 10993.861.0 - 1.2 (21.48) = 0.085 in
Minimum wall thickness, tP = 0.375 in
tS = tP + tC = 0.375 + 0.125 = 0.5 in
W = 3.14 [ 7.2 + 0.0417) (14.4 + 0.8 (7.2)] 0.0417 (490) = 9366.83 lb
Cv = exp { 7.0132 + 0.18255[ ln (9366.83) ] + 0.02297 [ ln (9366.83)]2} = $ 40, 279
CPL = 361.8 ( 7.2 ) 0.73960 (14.4) 0.70684 = $ 10, 264
Cp = FMCv + CPL = 1.2 (40, 279) + 10, 264 = $ 58, 599
Bare-Module cost
= 4.16 ( 58, 599 ) = $ 243, 772

PUMP
Pressure inlet, P1 = 1000kPa = 145.04psi
Pressure outlet, P2 = 1750kPa = 253.82psi
Pressure drop, ΔP = 750kPa = 108.78psi
Q = 93.57 m3/hr = 413.09 gpm
H = ΔP (2.31)SG = ΔPρ = 108.78 psi x 1lb/in21 psi x ft343.9 lb x 144 in21 ft2
H = 356.82 ft
S = Q (H)0.5 = 413.09(356.82)0.5 = 7803.14 gallon.ft0.5/min
ln S = 8.962
CB = exp [9.7171 - 0.6019(8.962) + 0.0519(8.962)2] = $ 4872.43
FT = 1, FM = 1.35 (Assume cast steel)
CP = FTFMCB = (1)(1.35)(4872.43) = $ 6577.78 for pump

PT = Q Hρ33000 = 413.09 galmin x 356.82 ft x 43.9 lbft3 x 0.1334ft31 gal x 133000
= 26.16 lb.ftmin
ln Q = 6.024
ηp = -0.316 +0.24015 (6.024) – 0.01199(6.024)2
= 0.6956
PB = PTηp = 26.160.6956 = 37.61 lb.ftmin
ln PB = 3.627
ηm = 0.80 + 0.0319(3.627) – 0.00182(3.627)2
= 0.892
PC = PTηpηm = 26.16(0.6956)(0.892) = 42.16 lb.ftmin
ln Pc = 3.741
CB = exp [5.8259+0.13141(3.741)+ 0.053255 (3.741)2 + 0.028628 (3.741)3 – 0.0035549(3.741)4]
= $ 2605.50
FT = 1.8 (assume explosion-prof enclosure)
CP = FTCB = 1.8(2605.50) = $ 4689.9 for motor
FBM = 3.30
CPTotal (Pump + Motor) = (6577.78 + 4689.9) (3.30)
= $ 37,183.34

DISTILLATION COLUMN
Distillation column, S1
Main vessel sizing

Diameter, DT
FLG= 0.1345
CSB=0.34
FST=0.757
Assume:
FF=1,FHA=1
C=0.2574
Uf=2.758 ft/s
Ad/AT=0.10378
Assume 80% flooding
DT=4(4251003600)0.8(2.758)(3.14)(1-0.10378)(2.09)=6.05 ft=1.844 m

Purchase costs of the vessel

P0=1500 Kpa=217.55 psig
Pd= 5.582 psig
Di=6.05 ft
L= 50 ft
tp=(5.582)(6.05)(12)2150000.85-(1.2*5.582)=0.0158 < tabulated data thus tp=0.375 in
ts=0.375+0.125=0.5 in
W= π (Di+ts) (L+0.8Di) ts ρ= 21458 Ib
Cv= $ 67436
CPL=300.9*6.050.63316500.8016= $21642
CT= $ 27462

Total cost after bare-module= 4.16(27462+21642+67436)= $ 484806

Cost of the reflux drum

Dvolumetric =3350
Volume flow = (1+3) *(3350) = 13400 ft3/hr
Assume residence time of 5 mins at full capacity and L/D=2
V=13400 ft3/hr * (5 min/ (60 min/hr))
V=116 ft3
D=8.924 ft3
L=17.85 ft3

tp =7/16=0.4375 in
ts =0.4375+0.125=0.562 in
W=π (Di + ts)*(L+0.8*Di)*(ts)*ρ
W=16150 Ib
Cv= $ 733210
After bare-model:
Cost= $ 223290

Condenser

Qc= -0.126*105
Uf=4.402 Btu/ (ft2*hr*F)
TLM= 10.48 F
Ac=273.1 ft2

Assume fixed head, shell tube exchanger and carbon steel, 20 feet long:
FL=1
FP=1
TM=2.732
CB= 7950

Cp= $ 21721
After bare-module,
Cost= $ 68857

Re-boiler costing

Heat flux 5000 Btu/hr.ft3
Q=16600 Kj/hr=15733.76 Btu/hr
AR=QR/ Flux
AR=3.146 ft2
Choose kettle reboiler with carbon steel, 20 ft long
FL=FM=FP=1
CB= $ 65325
With bare-module
$ 3.17(65325) = $ 207080

COMPRESSOR
(a) Preliminary estimate of brake horsepower, PB

Inlet volumetric flow rate, QI = 83,283.83 ft3/min
Inlet pressure, PI = 14.5 psi
Outlet pressure, PO = 72.52 psi
Specific heat ratio, k = 1.10
Mechanical efficiency, ηB = 0.4952

PB=0.00436×1.101.10-1×83283.83×14.50.4952×72.5214.51.10-11.10-1 =18430.94 BHp
(b) Purchase cost of compressor

Assumption:
Drive efficiency, ηC = 0.75
Material factor, FM = 1.00 (carbon steel)
Drive type factor, FD = 1.15 (steam turbine)

PC=18430.940.75=24574.59 Hp

Base purchase cost,
CB=exp7.5800+0.80lnPC
CB=exp7.5800+0.80ln24574.59
CB=$ 6,372,960

Total purchase cost,
CP=FDFMCB
CP=1.15×1.00×$ 6,372,960
CP=$ 7,328,904

HEAT EXCHANGER (HE2)
Sizing of Heat exchanger (HE 2)

Heat exchanger type

2 shell and 4 tubes
Design type

Fixed Head
Heat exchanger orientation
Horizontal
Tube inlet direction

Horizontal
Heat duty (kJ/s)

1594.2
Heat duty (Btu/hr)

5.44x10^6

Hot
Cold
Tin (˚C)
870.7
43.85
Tout (˚C)
349.12
576.3

From Figure 18.15 (a), FT = 0.85 and 2-4 exchanger is used.

Ui = 235.5 Btu/oF.ft2.hr

Velocity of tube-side;

Cross section are/pass;

By using 0.75 in. O.D. 16 BWG tubing with I.D. of 0.62 in.;

Inside area/tube =
= 2.097x10-3 ft2/tube

Area per tube;

=
= 0.288 ft2/tube

L = 5.58 ft

Costing of Heat Exchanger (HE2)

FBM = 3.17

FM=1.08+(86.34100)0.5
= 2.01

FL = 1.25 (Tube length = 5.58 ft2)

FP=0.9803+0.018 145.04100+0.0017 (145.04100)2

= 1.01

Fixed head:

CB=exp11.0545-0.9228ln86.34+0.09861ln86.342
= $7,334.88

CP=2.011.251.01(7,334.88)
= $18,613.18

Bare-module cost = 3.17 ( 18,613.18)
= $59,003

PROCESS FLOW DIAGRAM(HEAT EXCHANGER NETWORK)PROCESS FLOW DIAGRAM(HEAT EXCHANGER NETWORK)
PROCESS FLOW DIAGRAM
(HEAT EXCHANGER NETWORK)
PROCESS FLOW DIAGRAM
(HEAT EXCHANGER NETWORK)

Design for Production of Propylene

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