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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 . 1. 2. 3. 4. 5.

TEAM MEMBERS EMAD MOHAMMED NOMAN AL-DHUBHANI MUHAMMAD FAIRIS BIN HADIPORNAMA KELVIN RAWING SEBASTIAN NUR FADZLYANA BINTI HAMDAN MIMI KHAIRIAH BINTI AWANG

MATRIC NO AA103001 A11KK0022 A11KK0065 A11KK0035 A11KK0169

1

TABLE OF CONTENTS

Page TABLE OF CONTENTS

2

CHAPTER 1 INTRODUCTION 1.1

Background of Propylene

5

1.2

Uses of Propylene

6

1.3

Propylene Manufacturing

8

1.4

Market Survey

9

1.5

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

Raw Materials

14

1.5.1 Source of Raw Materials

14

1.5.2 Raw Material Specifications

14

1.5

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

24

Basic Material Balances

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

36

Mass Balance for Mixer

2

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

6.3

6.4

6.2.1

Sizing of reactor

69

6.2.2

Costing of reactor

70

Pump

70

6.3.1

Sizing of pump

70

6.3.2

Costing of pump

70

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 3

6.5

6.6

Compressor

74

6.5.1

Sizing of compressor

74

6.5.2

Costing of compressor

74

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

4

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.

5

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.

6

Table 1.1: Selected Propylene Application Product Application

Polypropylene

Application Description    

Propylene Oxide

 



Isopropanol

  

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. 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.

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 aftershaves; 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.



7

Oxo Process Chemicals

 

Polygas Chemicals (nonene, dodecene, heptenes)

  

1.3

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. 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.

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.

8

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 Europe US

80.0 million tones 14.3 million tons 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.

9

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:

Propylene Top Producers LyondellBasel Industries, Netherlands

Sinopec, China

Total SA, France

Ineos Group, England

ExxonMobil Chemical, USA

SABIC, KSA

PetroChina, Bejing, China

Formosa Plastics Group, Taiwan

Relliance Industries, India

3.8 4 4 4

5

5.8 6.1

9

15.8

Figure 1.2: Propylene top producers 10

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.

11

other western europe; 2 other asia; 3 canada; 0.5 Mexico; 1 Oceania; 1 Brazil; 3 USA; 20 France; 3 India; 4 Taiwan; 5

Benelux; 6 China; 19 Germany; 6 Japan`; 7 Middle east ; 7 Rep. of korea ; 8

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.

12

Table 1.3: Production, Import, Export and Consumption of PP in Malaysia Product Propylene

1.4.5

Unit: KTPA Production Import Export Consumption

2007 839 33 78 765

2008 870 40 97 811

2009 867 25 95 797

2010 808 8 50 744

change 0.3% 0% -2% -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 Propylene Propane Ethylene Butene Methanol

Price RM/Ib 1.962 0.79 2.158 1.118 1.635

RM/Kg 4.326 1.742 4.758 2.465 3.609

13

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. gases

that fall into

There are a number of hydrocarbon

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 14

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: 1. 2. 3. 4. 5.

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 15

A. 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)

16

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 deethanizer where light-ends are removed prior to the propane-propylene splitter.

17

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

1.81 kg/m3, gas (1.013

0.08988 g/L (at 0 °C,

°C, 101.3 kPa)

bar, 15 °C)

101.325 kPa) 3

613.9 kg/m , liquid Melting point

Boiling point

−187.7 °C; −305.8 °F;

− ,185.2 °C (−301.4 °F;

13.99 K (−259.16 °C,

85.5 K

88.0 K)

−434.49 °F)

−42.25 to −42.04 °C;

− 47.6 °C (−54 °F;

20.271 K (−252.879 °C,

−44.05 to −43.67 °F;

226 K)

−423.182 °F)

230.90 to 231.11 K 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 −105.2–−104.2 kJ mol−1

+20.41 kJ/mol

0

Std enthalpy of

−2.2197–−2.2187 MJ

-2058.4 kJ/mol

-285.84 kJ/mol

combustion ΔcHo298

mol−1

Std enthalpy of o

formation ΔfH 298

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

18

C3H8

C3H6

H2

1

1

1

Molecular weight

44.09

42.08

2.016

lb

44.09

42.08

2.016

1.0478

1

0.048

0.79

1.96

11.30

lbmol

lb/lb of propylene RM/lb

Gross profit for reaction path 1 = 1.96(1) + 11.30(0.048) – 1.0478(0.79) = RM 1.67 /lb propylene B. 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

19

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.

20

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 C 2 tower bottom is processed in the C3 tower to produce propylene product and a C 4 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

BUTENE

ETHYLENE



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

Boiling point

Solubility in water

−169.2 °C (104.0 K,

−185.3 °C (−301.5 °F;

− ,185.2 °C (−301.4 °F;

-272.6 °F)

87.8 K)

88.0 K)

−103.7 °C (169.5 K,

−6.47 °C (20.35 °F;

− 47.6 °C (−54 °F; 226 K)

-154.7 °F)

266.68

3.5 mg/100 mL (17 °C)[

0.61 g/m3

Thermochemistry Std enthalpy of

52.28 kJ mol−1

1.17 kJ/mol

+20.41 kJ/mol

-1410.99 kJ mol−1

-2718.6 kJ/mol

-2058.4 kJ/mol

o

formation ΔfH 298 Std enthalpy of combustion ΔcHo298

21

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

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

lbmol

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

22

Table 2.3: Summary of Review and Screening of Alternative Processes Dehydrogenation of propane C3H8  C3H6 + H2 Gross Profit (Appendix 1) Type of process

Metathesis of from Ethylene & Butenes C2H4 + C4H8  2C3H6

RM 1.67 / lb propylene

RM 0.46/lb propylene

Continuous process

Continuous process Butane and ethylene is

Safety

Propane is flammable.

flammable, and ethylene also may cause dizziness

By-product Operating condition

Hydrogen Temperature: 560 – 650 ºC Pressure : slightly below atmospheric pressure

No by-product 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 23

2.2.1

Basic Material Balance Reactor T = 500 OC P = 1 bar gfrgfr

m1C3H8

F lb/hr C3H8

m2C3H6 m3H2

gd

R 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 24

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

25

2.3

Step 3 – Eliminate Differences in Composition

0.99 C3H8 0.009 C4H10 0.001 C5H12 52.°C

Pt-Sn Reactor 600˚C, 1.0 bar

H2 C3H8 C3H6 C4H8 C4H10 C5H12

H2 -137.1˚c

C3H6 C3H8 40˚C

10 bar

17.5 bar

S2 S1 15 bar

LPG: C3 H 8 C4H10 C5H12 110.7° C C4H10 C5H12

137.1° C

C3H6 33˚C

S3

S4 15 bar

42˚C -137.1˚C C3H8 C3H6 C4H8 C4H10 C5H12

108.4˚ C C4 H 8 C4H10 C5H12

42˚C C3 H 8

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 26

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 C 3 is very low, - 48˚C, and hence if C 3 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

H2 C3H8 C3H6

Normal boiling point (1atm, ˚C) -252.78 -42.11 -47.62

Critical constant

Boiling point (˚C) 15 bar

17.5 bar

20 bar

Tc (˚C)

PC (bar)

41.00 33.00

45.00 35.85

53.55 42.65

-240.01 96.74 91.06

12.96 42.51 45.55

27

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. 28

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: 1. The product mixture is preheated before it is introduced to the reactor. The reaction occurs at around 600oC and 1 bar. 2. The products mixture is then cooled to its dew point -137.1oC at 10 bar. 3. 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 -137.1 oC at 10 bar from the condenser is operated upon as follows: 1. Its pressure is increased to 17.5 bar. 2. The temperature is then raised to a liquid at its bubble point, 42oC at 17.5 bar. 3. 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 40 oC. The propylene gas with a boiling point of 33 oC at 15 bar is come out as an upper product from separation column (S4). Finally, the propane liquid from the recycle stream (at 42 oC and 15 bar) undergoes the operation where its temperature is raised to the mixing temperature at 47 oC at 15 bar.

29

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: 1. 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. 2. Depropanizer A propane rich liquefied petroleum gas (LPG) feedstock is sent to a depropanizer to reject butanes and heavier hydrocarbons. 3. 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. 4. 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 5. 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 6. 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 30

condensate will contain more volatile components. At the same time, the original mixture will contain more of the less volatile material. 7. 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 C 3 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, C 3 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.

31

Figure 2.7:

P-P Splitter

Flowsheet task integration for the propylene production process

32

CHAPTER 3

MATERIAL AND ENERGY BALANCES AND PROCESS SIMULATION

3.1

MATERIAL BALANCES

3.1.1

Overall Mass Balance Reactor T = 500 OC P = 1 bar gfrgfr

m1C3H8

F lb/hr C3H8

m2C3H6 m3H2

gd

R 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 33

MW (lb/lbmol)

44.10

42.08

2.01

n, (lbmole/hr)

2376.43

2376.43

2376.43

34

3.1.2

Mass Balance for Separation Unit 1

2

D1 lb/hr C3H8 C4H10 0.001 C5H12 0.990 0.009

1

F1 lb/hr

S-1 C3H8 0.18 C4H10 0.02 C5H12 0.80

3

B1 lb/hr 0.90 C4H10 0.10 C5H12

Stream 1

Stream 2

Stream 3

No .

Component

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

Mass Flowrate, (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

35

3.1.3

Mass Balance for Mixer

D1 = 106263.1 lb/hr C3H8 0.009 C4H10 0.001 C5H12 0.990

3

F = 123323.6 lb/hr

M-1

4

a1 C3H8 a2 C4H10 a3 C5H12

10 R = 17060.53 lb/hr 1.0

Stream 3

C3H8 Stream 4

Stream 10

No .

Component

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

Mass Flowrate, (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

36

3.1.4

Mass Balance for Reactor

121861.09 lb/hr C3H8 4 + 1116.11 lb/hr C4H10 346.38 lb/hr C5H12 0.991 C3H8 0.007 C4H10 0.002 C5H12

123299.61 lb/hr

5

R-1

C3H8 C4H10 C5H12 C3H6 C4H8 H2

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.

Stream 4

Stream 5

No.

Component

Mole Fraction

Mass Flowrate, (lb/hr)

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

Mole Fraction

Mass Flowrate, (lb/hr)

H2 4811.36 lb/hr 123299.61 lb/hr

3.1.5

6 5

C3H8 C4H10 C5H12 Mass Balance for Separation Unit 2 C3H6 C4H8 H2

S-2 7 C3H8 C4H10 C5H12 C3H6 C4H8

118488.25 lb/hr

37

Stream 5

Stream 6

Stream 7

No.

Component

Mole Fraction

Mass Flowrate, (lb/hr)

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

-

-

Mole Fraction

Mass Flowrate, (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

38

3.1.6

Mass Balance for Separation Unit 3 117060.68 lb/hr 8

118488.25 lb/hr

7

C3H8 C4H10 C5H12 C3H6 C4H8

C3H8 C3H6

S-3

1427.57 lb/hr C4H10 C5H12 C4H8

9

Stream 7

Stream 8

Stream 9

No.

Component

Mass Mole Flowrate, Fraction (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

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

Mole Fraction

Mass Flowrate, (lb/hr)

39

3.1.7

Mass Balance for Separation Unit 4

11

100000 lb/hr C3H6

117060.68 lb/hr

8

S-4

C3H8 C3H6 17060.53 lb/hr 10

Stream 8

C3H8

Stream 1

Stream 11

No.

Component

Mole Fraction

Mass Flowrate, (lb/hr)

Mole Fraction

Mass Flowrate, (lb/hr)

1

Propane

0.140

17060.53

1

17060.56

0

0

2

Propene

0.860

100000

0

0

1

100000

Mole Fraction

Mass Flowrate, (lb/hr)

40

3.2

ENERGY BALANCES Table 3.1: Table of Data for Heat of Capacities

Cp=A+B*T+C*T^-2+D*T^-3

Compound

Molecula r Weight

Cp=A+B*T+C*T^-2+D*T^-3 ∆H f

∆H v Kj/mol

A*10^ 3

B*10^ 5

C*10^ 8

D*10^1 2

Average Cp for liquid KJ/ (mol.K)

Propane

44.09

103. 8

Propene

42.08

20.4 1

18.4 2

59.58

17.71

-10.17

24.6

0.1199

Butane

58.12

124. 7

22.3 06

92.3

27.88

-15.47

34.98

0.13367

Butene

56.1

1.17

21.9 16

82.88

25.64

-17.27

50.50

0.09396

Pentane

72.15

146. 4

25.7 7

114.8

34.09

-18.99

42.26

0.167

Hydroge n

2.016

0

0.90 4

28.84

0.007 65

0.3288

0.8698

-

18.7 7

68.02 3

22.59

-13.11

31.71

0.10584

Energy Balance 41

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: Qv  n y k [  C po.k (T )dT  H f ,k ] T2

k

T1

(Vapor) k QL  n xk [  C po.k (T )dT  H f , k (T2 )  H vap (T1 )] T2

k

Where,

T1

(Liquid)

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 2

D1 lb/hr C3H8 C4H10 0.001 C5H12 0.990 0.009

1

F1 lb/hr

S-1 C3H8 0.18 C4H10 0.02 C5H12 0.80

3

For Stream 1

B1 lb/hr 0.90 C4H10 0.10 C5H12

Liquid phase Stream temperature, T = 325.13 K and consider datum at 298.13 K 42

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 Stream temperature, T = 320.13 K and consider datum at 298.13 K

Componen t

Flow rate Ibmole/hr

Flow rate mole/hr

Cp liquid

∆H

Propane

2376.43

1077924. 884

0.10584

2509926.5 34

Butane

19.203

8710.288

0.13367

25614.69

Pentane

4.8

2177.23

0.167

7999.14

∑

2543540.36 KJ/hr

43

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.

44

Component

Flow rate Ibmole/hr

Propane

2376.96

Butane

534.7

Pentane

59.41

∑

Flow rate mole/hr 1077948.6 4

Cp

∆H

0.10584

3650882.69

242535.38

0.13367

1037430.53

26948.37

0.167

144012.08

4832325.3 KJ/hr

45

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

1077948.6 4

69.62

75046784.32

Butane

534.69

242535.38

90.91

22048891.4

Pentane

59.41

26948.37

112.17

3022798.66

∑

3.2.4

100118475 KJ/hr

Energy balance for Heat Exchanger 2 (HE2) 46

Stream outlet 1˚C (Mixture phase)

Stream inlet 600˚C (Gas Stream inlet at 600˚C datum at 25˚C Compound H2 C 3 H8 C 3 H6 C4H10 C 4 H8 C5H12 ∑

Flowrate (mol/hr) 2393.7 386.86 2376.4 1.9203 17.283 4.8009 5180.9642

∫CpdT (kJ/mol)

n∆H

16.8053 71.2935 59.5563 93.1652 82.9306 114.9629 438.7138

40226.84661 27580.60341 141529.5913 178.9051336 1433.28956 551.9253866 211501.1614

Stream outlet at 1˚C datum at 25˚C Compound H2 C 3 H8 C 3 H6 C4H10 C 4 H8 C5H12 ∑

Flowrate (mol/hr) 2393.7 386.86 2376.4 1.9203 17.283 4.8009 5180.9642

∫CpdT (kJ/mol)

n∆H

∆Hv

-0.6922 -2.5392 -2.8776 -3.2082 -2.255 -4.008 -15.5802

-1656.91914 -982.314912 -6838.32864 -6.16070646 -38.973165 -19.2420072 -9541.938571

18.77 18.42 22.306 21.916 25.77 107.182

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

47

3.2.5

Energy Balance for Flash Separator (S2) Stream 6 1˚C

Stream 5 1˚C

Flash separator

Stream 7 1˚C Stream 5 (Feed Stream) at 1˚C datum at 25˚C Compound H2 C 3 H8 C 3 H6 C4H10 C 4 H8 C5H12 ∑

Flowrate (mol/hr) 2393.7 386.86 2376.4 1.9203 17.283 4.8009 5180.9642

∫CpdT (kJ/mol)

n∆H

-0.6922 -2.5392 -2.8776 -3.2082 -2.255 -4.008 -15.5802

-1656.91914 -982.314912 -6838.32864 -6.16070646 -38.973165 -19.2420072 -9541.938571

Stream 6 (Distillate stream) at 1˚C datum at 25˚C Compound H2

Flowrate (mol/hr) 2393.7

∫CpdT (kJ/mol)

n∆H

-0.6922

-1656.91914

∫CpdT (kJ/mol)

n∆H

-2.5392 -2.8776 -3.2082 -2.255 -4.008 -14.888

-982.314912 -6838.32864 -6.16070646 -38.973165 -19.2420072 -7885.019431

Stream 7 (Bottom stream) at 1˚C datum at 25˚C Compound C 3 H8 C 3 H6 C4H10 C 4 H8 C5H12 ∑

Flowrate (mol/hr) 386.86 2376.4 1.9203 17.283 4.8009 2787.2642

There’s no heat transfer from the flash column:

Q= ΔH=0 48

Q = n∆Hout - n∆Hin =(-7885.019431+ (-1656.91914)) – (-9541.938571) = 0 kJ/hr

3.2.6

Energy balance for heat exchanger 3

Stream inlet 1˚C

Stream outlet 42˚C

49

Stream inlet at 1˚C datum at 25˚C Compound C 3 H8 C 3 H6 C4H10 C 4 H8 C5H12 ∑

Flowrate (mol/hr) 386.86 2376.4 1.9203 17.283 4.8009 2787.2642

∫CpdT (kJ/mol)

n∆H

-2.5392 -2.8776 -3.2082 -2.255 -4.008 -14.888

-982.314912 -6838.32864 -6.16070646 -38.973165 -19.2420072 -7885.019431

∫CpdT (kJ/mol)

n∆H

1.2825 1.1118 1.7249 1.5516 2.1421 7.8129

496.14795 2642.08152 3.31232547 26.8163028 10.28400789 3178.642106

Stream outlet at 42˚C datum at 25˚C Compound C 3 H8 C 3 H6 C4H10 C 4 H8 C5H12 ∑

Flowrate (mol/hr) 386.86 2376.4 1.9203 17.283 4.8009 2787.2642

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

3.2.7

Energy Balance for Compressor (C1)

STREAM INLET 1 bar 600˚C

STREAM OUTLET 10 bar 600˚C

Inlet (stream 2) Phase Component

Vapor Mixture 50

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.

T 2 =T 1

P2 P1

(γ −1γ )

( )

T 2 =600

10 1

(1 .31 .3−1 )

For ideal system, γ = 1.3. Hence,

( )

¿ 1020.75˚ C For energy balance,

Q=∆ H =∑ ni ^ H i − ∑ ni ^ Hi ¿

out

T

¿ ∑ ni out

∫

298.15

T

C p dT −∑ ni ¿

∫

C p dT

298. 15

Since there is no change component flow rate,

[∫

1293.9

Q=∆ H =ni

298.15

1293.9

68.023× 10−3 +22.59 ×10−5 T −13.11× 10−8 T 2 +31.71× 10−12 T 3 dT +

∫

59.58× 10−3 +17.71× 10−5

298.15

51

175.44+144.36+227 +197.87+278.61+30.52 ) ¿ 2349.64 ¿ = 2476050.63 kJ/hr

3.2.8

Separation Unit 3 (S3)

117060.68 lb/hr 8

118488.25 lb/hr

7

C3H8 C4H10 C5H12 C3H6 C4H8

C3H8 C3H6

S-3

1427.58 lb/hr 9

For Stream 7 Liquid stream

C4H10 C5H12 C4H8

Stream temperature, T = 315.15 K and consider datum at 298.15 K Component

Flow rate (lbmol/hr)

Flow rate (mol/hr)

Cp

∫CpdT (kJ/mol)

n∆H kJ/hr 52

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 Stream temperature, T = 317.15 K and consider datum at 298.15 K 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 Stream temperature, T = 381.55 K and consider datum at 298.15 K Component

Flow rate (lbmol/hr)

Flow rate (mol/hr)

Cp

∫CpdT (kJ/mol)

n∆H kJ/hr

53

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) 11

D1 lb/hr

C3H6

F1 lb/hr

8 S4 C3H8 C3H6 12

B1 lb/hr

C3H8

For Stream 8 Liquid stream Stream temperature, T = 318.15 K and consider datum at 273.15 K

Component

Flow rate (lbmol/hr)

Flow rate (mol/hr)

Cp

∫CpdT (kJ/mol)

n∆H kJ/hr 54

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 Stream temperature, T = 309.15 K and consider datum at 273.15 K

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 Stream temperature, T = 320.15 K and consider datum at 273.15 K 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 ∑ 55

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)

outlet inlet

Component

Molar Flow Rate (lbmol/hr) C3H8 386.86 C3H6 2376 C4H10 1.92 C4H8 17.28 4.801 C5H12 ∑Fv = 28906.35

Molar Flow Rate, F (kmol/hr) 1.75476 х 102 10.77734 х 102 0.00870 х 102 0.07838 х 102 0.02178 х 102

Molar volume,v(kmol/m3)

Fv

21.9375 23.2486 0.09667 0.09048 0.11452

3849.50 25055.81 0.08410 0.70918 0.24942

Inlet Stream Phase Pressure (bar) Temperature (oC)

Liquid 1 -47

Outlet Stream Phase Pressure (bar) Temperature (oC)

Liquid 18 -47

Q = ∑Fv ( P) Q = 28906.35 х (18 -1) = 4.91408 х 105 kJ/hr 56

3.2.11 Energy Balance for Reactor

873K (600˚C), 1 bar

873K (600˚C), 1 bar

4 121861.09 lb/hr C3H8 + 1116.11 lb/hr C4H10 346.38 lb/hr C5H12

R-1

5

0.991 C3H8 0.007 C4H10 0.002 C5H12

C3H6 + H2

∆^ Hr1

2. C4H10

C4H8 + H2

∆^ Hr2 >

1. C3H8

C3H8 C4H10 C5H12 C3H6 C4H8 H2

>

C3H8 (g), 873K (600˚C), 1 bar

C3H6 (g), 873K (600˚C), 1 bar

> C3H8 (g), 298K (25˚C), 1 bar

C3H6 (g), 298K (25˚C), 1 bar

Reaction path for Propane

^ R 1+ n ^ Q1=n ΔH =∑ n H H p 1+ ∆ ^ H r 1❑

(for Propane)

^ R 2 +n ^ Q2=n ΔH =n ∑ n H H p 2+ ∆ ^ H r 2❑

(for Butane)

Q=Q 1 +Q 2

57

1. Energy balance for Propane ^ H R1

^ H p1

∆^ Hr1

1253192.74

2330911.57

-

-91.8

184.72

-

-115043093.5

430565985.2

140.21

Component Flow rate (mol/hr) Specific ^ Enthalpy, H (kJ/mol) ∆H (kJ/hr) ∑

315523031.9

2. Energy balance for Butane ^ H R2

^ H p2

∆^ Hr2

8708.84

16547.07

-

-119.0

240.61

-

-1036351.96

3981390.5

125.87

Component Flow rate (mol/hr) Specific ^ Enthalpy, H (kJ/mol) ∆H (kJ/hr) ∑

1945164.41

Q=Q 1 +Q 2 Q=315523031.9

kJ kJ +1945164.41 hr hr

Q = 317.468 x 106 kJ/hr 58

59

3.3

Simulation Result from ASPEN HYSYS 3.3.1

3.3.2

3.4

Material Balance Stream no.

Mass (lbmole/hr) (Hysys)

stream 1 stream 2 stream 3 stream 4 stream 5 stream 6 stream 8 stream 9 stream 10 stream 11 stream 12 stream 13 stream 14 stream 15 stream 16 stream 17 stream 23 stream 24 stream 25

3179 3179 769.4 2410 2523 2523 4692 4692 4692 4692 2533 2159 2533 2533 156.3 2377 2263 113.4 113.2

Energy Balance Equipment

Energy, kJ/h (HYSYS)

Q-HE00 Q-HE01 Q-HE02 Q-Furnace Q-Compressor Q-Pump

5.76E+06 1.79E+08 1.65E+07 9.23E+07 6.78E+07 6.90E+04

Percentage Difference between Manual Calculation and HYSYS Calculation 3.4.1

Mass Balance 60

Stream no.

3.4.2

Energy

stream 1 stream 2 stream 3 stream 4 stream 5 stream 6 Equipment stream 8 stream 9 stream 10 Q-HE00 stream 11 Q-HE01 stream 12 Q-HE02 stream 13 stream 14 Q-Furnace stream 15 Q-Compressor stream 16 Q-Pump stream 17 stream 23 stream 24 stream 25

Mass (lbmole/hr) (manual) 2971 2971 721.5 2400.4 2787.3 2787.3 Energy, kJ/h 5181 (manual) 5181 5181 4.83E+06 5181 1.80E+06 2787 6.33E+04 2393 2787 1.00E+08 2787 2.48E+06 165.4 4.91E+05 2663.5 2376.4 124.5 124.5

Mass (lbmole/hr ) (Hysys)

% Diff

3179 3179 769.4 2410 2523 2523 Energy, kJ/h 4692 (HYSYS) 4692 4692 5.76E+06 4692 1.79E+08 2533 1.65E+07 2159 2533 9.23E+07 2533 6.78E+07 156.3 6.90E+04 2377 2263 113.4 113.2

6.54 6.54 6.23 0.40 Balance 10.48 10.48 % Diff 10.42 10.42 10.42 10.4216.09 10.0399.00 10.8499.62 10.03 8.46 10.0396.35 5.82612.70 12.05 5.01 9.79 9.98

CHAPTER 4

HEAT INTEGRATION

4.1

PROCESS ENERGY INTEGRATION

∆Tmin = 10˚C

Table 4.1: Steam Table Data

61

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

Cold

-136.8

30

0.028

T (˚C)

0.04 9

∆T (˚C)

865.70

H1

4.2

C3 Algorithm Table

∑FCpC -∑FCpH (MW/K)

∆Hi (MW)

0 284.4

-0.049

13.94 526.22

-0.001

-0.52622

55.08

14.46 6.23

48.85

0.052

0.32396

C2

14.14

0.04 8

13.85

0.004

0.0554

35.00

14.08 10

0.032

0.3200

C1

13.76

0.05 3 -131.80

Pinch

-13.9356

581.30

25.00

1st Cascade

156.8

-0.021

-3.2928

C3 0.02 8

17.06 10.3

-0.049

-142.10

-0.5047 17.56 Qc

Figure 4.1: Algorithm Table

62

4.3

Heat Exchanger Network

∆H (MW)

T pinch FCp(MW/K)(870.7˚C)

1.5942

0.053

C1 50.08

25.5576

0.048

20 C2

576.3 4.6704

43.85

0.028

C3 30

-136.8

H1

49.3822

0.049 870.7

E11

25.5576

E22

E3

1.5942

C 4.6704

-137.1

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 E2

870.7 349.12

349.12 316.59

43.85 20

576.3 50.08

E3

316.59

221.28

-136.8

30

C

221.28

-137.1

-

-

4.4

Process Flow Diagram Heat Exchanger Network

Figure 4.3: 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: a) Inequality constraints Propane supply

R > 100000 Ib/hr

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

c) 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

6.2.2

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

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 =

6.3.2

ΔP(2.31) SG

=

ΔP ρ

= 356.82 ft

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 Domed head wall

S1 13.7 mm

thickness, a Tray spacing, b

2ft

Column diameter, c

6.05 ft

Column wall thickness, d

Design type

Vertical

Material of Construction

Carbon steel

Material of insulation

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

0.5 in

Mineral wool,60mm

Column type

Plate column

Plate type

Sieve

Domed head type

Torispherical

6.4.2

Sizing and costing of the reflux drum: Parameters Domed head wall

S1 13.7 mm

thickness, Vessel length,L

8.924 ft

vessel diameter, D

17.85 ft

Column wall thickness, d

0.562 in

Design type

Vertical

Material of Construction

Carbon steel

Material of insulation

Mineral wool,60mm

After bare-model: Cost= $ 223290

6.4.3

Sizing and costing of the condenser:

Parameters Length of tube

S1 20 ft

Area of transfer,Ac

273.1 ft2

Material of Construction

Carbon steel

fixed head, Type of HE

shell tube exchanger

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

6.4.4

Sizing and costing of the re-boiler:

Parameters Length of tube

S1 20 ft

Area of transfer,AR

3.146 ft2

Material of Construction

Carbon steel

Type of HE

kettle reboiler

Vessel or Equipment main vessel

Cost in $ 484809

the reflux drum

223290

the condenser

68857

the re-boiler

207080

Total

984036

CB= $ 65325 With bare-module $ 3.17(65325) = $ 207080

Total cost for S1:

6.5

Compresssor

6.5.1

Main Sizing Parameters

Parameters Compressor Type Drive Type Material of Construction Inlet Volumetric Flow Rate, QI Inlet Pressure, PI Outlet Pressure, PO Specific Heat Ratio, k

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)

Compressor Centrifugal Steam turbine Stainless steel 83283.83 ft3/min 14.5 psi 72.52 psi 1.10

Heat exchanger type Design type Heat exchanger orientation Tube inlet direction Heat duty (kJ/s) Heat duty (Btu/hr)

Q

2 shell and 4 tubes Fixed Head Horizontal Horizontal 1594.2 5.44x10^6 Hot

Cold

Tin (˚C)

870.7

43.85

Tout (˚C)

349.12

576.3

1594.2kJ 0.94782 Btu 3600 s x x s kJ hr

Q  5.44 x10 6 Btu / hr

LM 

(349.12  43.85)  (870.7  546.3) 349.12  43.85 ln( ) 870.7  546.3

LM  314.74 o F

870.7  349.12 576.3  43.85

R

R  0.98

S

576.3  43.85 870.7  43.85

S  0.644

From Figure 18.15 (a), FT = 0.85 and 2-4 exchanger is used. Ui = 235.5 Btu/oF.ft2.hr Ai 

5.44 x10 6 Btu / hr 235.5 Btu / oF . ft 2 .hrx 0.85 x314.74 o F

Ai  86.34 ft 2

Velocity of tube-side; 99.59m 3 144in 2 / ft 2 35.3145 ft 3 ui  x x hr 0.302in 2 m3 u i  1.68 x10 6 ft / hr

Cross section are/pass; Aci 

111244.12lb ft 3 hr x x hr 0.421lb 1.68 x10 6 ft

Aci  0.157 ft 2 / pass

By using 0.75 in. O.D. 16 BWG tubing with I.D. of 0.62 in.; 0.302in 2 ft 2 x tube 144in 2

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

0.157 ft 2 / pass 2.097 x10 3 ft 2 / tube

N t  75tubes / pass Area per tube; 86.43 ft 2 4 passx 75tubes / pass

= = 0.288 ft2/tube

L

0.288 ft 2 / tube 0.62in x 12in / ft

L = 5.58 ft 6.6.2

Costing of Heat Exchanger (HE2)

Ai  86.34 ft 2

FBM = 3.17 F M =1.08+(

86.34 0.5 ) 100

= 2.01 FL = 1.25 (Tube length = 5.58 ft2) F P=0.9803+ 0.018

(

145.04 145.04 2 +0.0017 ( ) 100 100

)

= 1.01 Fixed head: C B=exp {11.0545−0.9228 [ ln ( 86.34 ) ] +0.09861 [ ln ( 86.34 ) ]

2

= $7,334.88

}

C P=( 2.01 ) ( 1.25 ) (1.01 )(7,334.88) = $18,613.18 Bare-module cost = 3.17 ( 18,613.18) = $59,003

CHAPTER 7

TOTAL CAPITAL INVESTMENT AND PAYBACK PERIOD

7.1

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

Labe l

Size

Material of Construction

Design Temp. (˚C)

Design Pressure (bar)

Baremodule Cost, CBM

Reactor

R1

Low- Alloy Steel SA-387B

576

1

$ 243, 772

Pump

P1

V=586.17 ft3 Di = 7.20 ft L= 14.4 ft H =356.82 ft

Cast Steel

-137.1

$ 37,183

Distillation Column

S1

D = 6.05 ft t = 0.5 in

Carbon steel

50

Pinlet = 10 Poutlet = 17.5 15

Compressor

C1

Carbon Steel

526

Heat Exchanger

HE2

Q = 83283.83 ft3/min A = 86.34 ft2

Carbon Steel

870.7

Pinlet = 1 Poutlet = 10 10 CTBM ∑

$ 984,036

$ 15,757,144 $ 59,003 $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

7.2

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:

Initial Investment Cash Inflow per Period

Payback Period = ¿

RM 152,764,816.40 RM 167,327.77 /hr

¿ 912.967 hr ×

1day 1 month × 24 hr 30 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 : F 1 = D 1 + 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

F 1 = D 1 + 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=2763.3

lbmol 1 kgmol 1000 mol x x =¿ 1253197.28mol/hr hr 2.205 lbmol 1 kgmol

For reaction of propane 298 K

H R1=

∫

C p(reactant) dt

873 K

298 K

H R1=

∫

( 68.023 ×10−3+ 22.59× 10−5 T −13.11 ×10−8 T 2+ 31.71×10−12 T 3 ) dt

873 K

H R 1 =−91.81kJ /mol 298 K

H R1=

∫

873 K

C p( product ) dt

873 K

H p 1=

∫

( 68.023× 10−3 +22.59 ×10−5 T−13.11× 10−8 T 2 +31.71× 10−12 T 3 ) dt+ ( 59.58 ×10−3+17.71 ×10−5 T −10.17 ×10−8 T 2+24.6 × 10−12 T 3 ) dt + ( 28

298 K

H p 1=91.79 kJ /mol+75.75 kJ /mol+17.18 kJ /mol H p 1=¿ 184.72kJ/mol ∆^ H r 1=∑ v i ∆ ^ H f =( 1 ) ( ∆ ^ H f )C 3 H 6 + ( 1 ) ( ∆ ^ H f ) H 2− ( 1 ) ( ∆ ^ H f ) C3 H 8 ∆^ H r 1=

(−119.8 ) kJ 20.41 kJ +0− =140.21 kJ /mol mol 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) × (

5 min ×1 hr 60 min

×2 ) = 586.17 ft3

Assume L/ D = 2 V= π

(D/2)2L = ( π D3)/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 ( 10993.86 ) ( 1.0 )−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

=

ΔP ρ

= 108.78 psi x

H = 356.82 ft

1 lb /¿ 2 1 psi

x

ft 3 43.9 lb

x

144 ∈2 1 ft 2

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 =

QHρ 33000 = 413.09

= 26.16

gal min

x 356.82 ft x

lb . ft min

ln Q = 6.024 ηp = -0.316 +0.24015 (6.024) – 0.01199(6.024)2 = 0.6956 PB =

PT ηp

=

26.16 0.6956

= 37.61

lb . ft min

ln PB = 3.627 ηm = 0.80 + 0.0319(3.627) – 0.00182(3.627)2 = 0.892

43.9 lb ft 3

x

0.1334 ft 3 1 gal

x

1 33000

PC =

PT η pηm

=

26.16 (0.6956)(0.892)

= 42.16

lb . ft min

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 425100 ) 3600 DT= =6.05 ft=1.844 m 0.8(2.758)(3.14)(1−0.10378)(2.09) 4(

Purchase costs of the vessel P0=1500 Kpa=217.55 psig Pd= 5.582 psig Di=6.05 ft L= 50 ft (5.582)(6.05)(12) tp= 2 ( 15000 ) ( 0.85 )−(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 0.63316 500.8016 = $21642 CPL= 300.9∗6.05

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

5

Qc= -0.126* 10

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 ×

(

(( )

1.10 83283.83× 14.5 72.52 × × 1.10−1 0.4952 14.5

)

1.10−1 1.10

)

−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 =

18430.94 =24574.59 Hp 0.75

Base purchase cost, C B=exp {7.5800+0.80 [ ln ( PC ) ] } C B=exp {7.5800+0.80 [ ln ( 24574.59 ) ] } C B=$ 6,372,960

Total purchase cost, C P=F D F M C B C P=1.15 ×1.00 × $ 6,372,960

C P=$ 7,328,904

HEAT EXCHANGER (HE2) Sizing of Heat exchanger (HE 2) Heat exchanger type Design type Heat exchanger orientation Tube inlet direction Heat duty (kJ/s) Heat duty (Btu/hr)

Tin (˚C)

2 shell and 4 tubes Fixed Head Horizontal Horizontal 1594.2 5.44x10^6 Hot

Cold

870.7

43.85

Tout (˚C)

Q

349.12

576.3

1594.2kJ 0.94782 Btu 3600 s x x s kJ hr

Q  5.44 x10 6 Btu / hr

LM 

(349.12  43.85)  (870.7  546.3) 349.12  43.85 ln( ) 870.7  546.3

LM  314.74 o F

R

870.7  349.12 576.3  43.85

R  0.98

S

576.3  43.85 870.7  43.85

S  0.644

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

Ui = 235.5 Btu/oF.ft2.hr 5.44 x10 6 Btu / hr Ai  235.5 Btu / oF . ft 2 .hrx 0.85 x314.74 o F

Ai  86.34 ft 2

Velocity of tube-side;

ui 

99.59m 3 144in 2 / ft 2 35.3145 ft 3 x x hr 0.302in 2 m3

ui  1.68 x10 6 ft / hr

Cross section are/pass; Aci 

111244.12lb ft 3 hr x x hr 0.421lb 1.68 x10 6 ft

Aci  0.157 ft 2 / pass

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

0.302in 2 ft 2 x tube 144in 2

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

0.157 ft 2 / pass 2.097 x10 3 ft 2 / tube

N t  75tubes / pass Area per tube;

86.43 ft 2 4 passx 75tubes / pass

= = 0.288 ft2/tube

L

0.288 ft 2 / tube 0.62in x 12in / ft

L = 5.58 ft Costing of Heat Exchanger (HE2)

Ai  86.34 ft 2 FBM = 3.17

0.5

F M =1.08+(

86.34 ) 100

= 2.01 FL = 1.25 (Tube length = 5.58 ft2) 2

145.04 145.04 F P=0.9803+ 0.018 +0.0017 ( ) 100 100

(

)

= 1.01 Fixed head: C B=exp {11.0545−0.9228 [ ln ( 86.34 ) ] +0.09861 [ ln ( 86.34 ) ]

2

= $7,334.88 C P=( 2.01 ) ( 1.25 ) (1.01 )(7,334.88) = $18,613.18 Bare-module cost = 3.17 ( 18,613.18) = $59,003

}

PROCESS FLOW DIAGRAM (HEAT EXCHANGER NETWORK)