PROPYLENE PRODUCTION PLANT
1
TABLE OF CONTENTS
Page TABLE OF CONTENTS
2 2
CHAPTER 1 INTRODUCTION
1.1
Background of Propylene
5
1.2
Uses of Propylene
6
1.3 1.3
Prop Propyl ylen enee Manu Manufa fact ctur urin ingg
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. 1.4.44 Outl Outloo ookk for Prod Produc ucti tion on of Propy ropyle lene ne in Mala alaysia ysia
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 1-Eliminate Differenc Differences es in Molecular Molecular Type
15
2.2
Step 2- Distribute Distribute the Chemicals Chemicals
24
3
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
36
Mass Balance for Mixer
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
Percentage Differences between Manual Calculation and HYSYS
61
3.4
4
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
6.2.1
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 5
6.4
6.5
6.6
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
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
6
7
CHAPTER 1
INTRODUCTION
1.1
Background of Propylene 8
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 C 3H6. 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.
9
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.
10
Table 1.1: Selected Propylene Application
Product Application
Polypropylene
Application Description
●
●
●
●
Propylene Oxide
●
●
●
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. 11
Isopropanol
●
●
●
Cumene
Ethylene-Propylene Elastomers
● ●
●
●
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.
Alkylation of benzene with propylene yields cumene. Consuming in phenol production for the manufacture phenolic resins, caprolactam and bisphenol A.
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.
12
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 13
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 o ther 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
14
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.
15
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:
16
Figure 1.2: 1.4.3
Propylene top producers
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 17
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.
18
Figure 1.3: World Consumption of Propylene in 2010
1.4.4
Outlook Outlook for Producti Production on of Propylen Propylenee in Malaysia Malaysia
19
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%
Consumptio
765
811
797
744
-5%
n
1.4.5
Market Market Prices Prices of Polyprop Polypropylene ylene
20
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
Price Product 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
21
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.
22
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
23
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. Eliminate differences in molecular types 2. Distribute the chemicals by matching sources and sinks 24
3. Eliminate differences in composition 4. Eliminate differences in temperature, pressure, and phase 5. Task integration; combination of operations into unit processes and decide between continuous and batch processing
2.1
Step 1 – Eliminate Differences in Molecular Type
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. 25
Figure 2.1:
Commercial process flow diagram (Dehydrogenation)
26
Figure 2.2: Detailed process flow diagram
(Dehydrogenation)
Process Description of Propylene Dehydrogenation
27
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.
28
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 o formation Δ H f 298
−105.2–−104.2 kJ mol −1
+20.41 kJ/mol
0
29
Std enthalpy of combustion Δc H o298
−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
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
0.79
1.96
11.30
lbmol
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
30
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
31
Figure 2.3: Commercial process flow diagram (Metathesis)
Figure 2.4: Detailed process flow diagram (Metathesis)
Process Description of Metathesis of Ethylene and Butene
32
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 C 4s (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 C 2 tower bottom is processed in the C3 tower to produce propylene product and a C4 recycle stream. Purge streams containing non-reactive light material, C 4s 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 C 4 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 33
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 o formation Δ H f 298
52.28 kJ mol−1
1.17 kJ/mol
+20.41 kJ/mol
Std enthalpy of combustion Δc H o298
-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
lbmol Molecular weight
→
C2H4
C4H8
C3H6
1
1
2
28.05
56.10
42.08
2C3H6
34
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
35
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
RM 1.67 / lb propylene
RM 0.46/lb propylene
Continuous process
Continuous process
(Appendix 1) Type of process
Butane and ethylene is Safety
Propane is flammable
.
flammable, and ethylene also may cause dizziness 36
By-product
Hydrogen Temperature: 560 – 650 ºC
Operating condition
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.
37
2.2
Step 2 – Distribute the Chemicals
2.2.1
Basic Material Balance
Overall Reaction Equation : C3H8
C3H6 + H2 38
Basis : 100000 lb/hr of propylene (C 3H3) 86% of conversion C3H8
C3H6
H2
1
1
1
m1
m2=100,000.00
m3
MW (lb/lbmol)
44.10
42.08
2.01
n, (lbmole/hr)
2376.43
2376.43
2376.43
stoichiometry Mass flowrate (lb/hr)
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, m 1 = 2376.43 x 44.1 = 104800.56 lb/hr 39
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 H 2 , m3 = (no. of mole) X (molecular weight) = (2376.43) x (2.01) = 4776.62 lb/hr Mass flowrate feed to the reactor,
F = m 1 + R = 104800.56 + 17060.53 = 121861.09lb/h
2.3
Step 3 – Eliminate Differences in Composition
40
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.
41
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 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
Normal boiling point (1atm, ˚C)
Critical constant Boiling point (˚C)
15 bar
17.5 bar
20 bar
T c (˚C)
PC (bar)
42
2.4
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
Step 4 – Eliminate Differences in Temperature, Pressure, and Phase
43
44
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: 1. The product mixture is preheated before it is introduced to the reactor. The reaction occurs at around 600 oC 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).
45
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.
46
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
47
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
48
To separate the mixture of C 3 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 boiling mixture mixture will be richer richer in the components that have lower boiling boiling points. Therefor Therefore, e, when this vapour is cooled and condensed, condensed, the 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.
49
50
process
CHAPTER 3
MATERIAL AND ENERGY BALANCES AND PROCESS SIMULATION
3.1 3.1
MATE MA TERIA RIAL L BALAN BALANCE CES S
3.1.1
Overall Overall Mass Balance Balance
51
Overall Reaction Equation : C3H8
C3H6 + H2
Basis : 100000 lb/hr of propylene (C 3H3) 86% of conversion
stoichiometry Mass flowrate (lb/hr) MW (lb/lbmol)
C3H8
C3H6
H2
1
1
1
m1
m2=100,000.00
m3
44.10
42.08
2.01
52
n, (lbmole/hr)
2376.43
2376.43
2376.43
53
3.1.2
Mass Balance for Separation Unit 1
Stream 1 No .
Component
Mole Fraction
1
Propane
0.80
Mass Flowrate, (lb/hr) 104800.56
Stream 2 Mole Fraction 0.990
Mass Flowrate, (lb/hr) 104800.56
Stream 3 Mole Fraction
Mass Flowrate, (lb/hr)
0
0 54
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
55
3.1.3
Mass Balance for Mixer
Stream 3
Stream 4
No .
Component
Mole Fraction
1
Propane
0.990
104800.56
0.991
2
Butane
0.009
1116.11
3
Pentane
0.001
346.38
Mass Flowrate, (lb/hr)
Mass Flowrate, (lb/hr)
Stream 10 Mole Fraction
Mass Flowrate, (lb/hr)
121861.09
1.0
17060.53
0.007
1116.11
0
0
0.002
346.38
0
0
Mole Fraction
56
3.1.4
Mass Balance for Reactor
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 No.
Component
Mole Fraction
Stream 5 Mass Flowrate, (lb/hr)
Mole Fraction
Mass Flowrate, (lb/hr)
57
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
58
Stream 5
Stream 6
Stream 7
No.
Component
Mole Fraction
Mass Flowrate, (lb/hr)
Mole Fraction
Mass Flowrate, (lb/hr)
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
-
-
59
3.1.6
Mass Balance for Separation Unit 3
Stream 7 No.
Component
Mole Fractio n
Stream 8 Mass Flowrate, (lb/hr)
Mole Fraction
Stream 9 Mass Flowrate, (lb/hr)
Mole Fraction
Mass Flowrate, (lb/hr) 60
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
61
3.1.7
Mass Balance Balance for Separation Separation Unit 4
Stream 8
Stream 1
Stream 11
No.
Component
Mole Fraction
Mass Flowrate, (lb/hr)
Mole Fraction
Mass Flowrate, (lb/hr)
Mole Fraction
Mass Flowrate, (lb/hr)
1
Propane
0.140
17060.53
1
17060.56
0
0 62
2
Propene
0.860
100000
0
0
1
100000
63
3.2
ENERGY BALANCES
Table 3.1: Table of Data for Heat of Capacities Cp=A+B*T+C*T^-2+D*T^-3
Cp=A+B*T+C*T^-2+D*T^-3 Compound
Molecular Weight
Hf
Average Cp for liquid
Hv Kj/mol
A*10^3
B*10^5
C*10^8
D*10^1 2
KJ/(mol.K)
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
22.30 6
92.3
27.88
-15.47
34.98
0.13367
Butene
56.1
1.17
21.91 6
82.88
25.64
-17.27
50.50
0.09396
64
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
0.0076 5
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:
65
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
66
For Stream 1 Liquid phase Stream temperature, T = 325.13 K and consider datum at 298.13 K
Flow rate
∆H
Flow rate
Component
Cp
Ibmole/hr
mole/hr
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
67
Flow rate
Flow rate
Cp
Ibmole/hr
mole/hr
liquid
Propane
2376.43
1077924.88 4
0.10584
Butane
19.203
8710.288
0.13367
25614.69
Pentane
4.8
2177.23
0.167
7999.14
Component
∆H 2509926.534
2543540.36 KJ/hr
68
For Stream 3 /Stream temperature, T = 377.13 K and consider datum at 298.13 K Flow rate
Flow rate
Cp
Ibmole/hr
mole/hr
liquid
Butane
8.86
4018.8
0.13367
42438.24
Pentane
0.756
342.91
0.167
4524.011
Component
∑
∆H
46962.25 KJ/hr
H=46962.25 +2543540.36 -4077274.484 =-1486771.86 KJ/hr
69
3.2.2
Heat exchanger 1
70
Flow rate
Flow rate
Ibmole/hr
mole/hr
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
Component
Cp
H
4832325.3 KJ/hr
71
3.2.3
Furnace
Stream temperature, T = 600 C and consider datum at 47 ˚C
∆H
Flow rate
Flow rate
Ibmole/hr
mole/hr
Propane
2376.43
1077948.64
69.62
75046784.32
Butane
534.69
242535.38
90.91
22048891.4
Component
∫CpdT
KJ/hr
72
Pentane
∑
3.2.4
59.41
26948.37
112.17
3022798.66
100118475 KJ/hr
Energy balance for Heat Exchanger 2 (HE2)
73
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 CpdT (kJ/mol)
n H
H2
Flowrate (mol/hr) 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
Compound
C5H12
Hv
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
74
75
3.2.5
Energy Balance for Flash Separator (S2)
Stream 5 (Feed Stream) at 1˚C datum at 25˚C CpdT (kJ/mol)
n H
H2
Flowrate (mol/hr) 2393.7
-0.6922
-1656.91914
C3H8
386.86
-2.5392
-982.314912
Compound
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
76
Compound H2
Flowrate (mol/hr) 2393.7
CpdT (kJ/mol)
n H
-0.6922
-1656.91914
Stream 7 (Bottom stream) at 1˚C datum at 25˚C CpdT (kJ/mol)
n H
C3H8
Flowrate (mol/hr) 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
Compound
C5H12
4.8009
-4.008
-19.2420072
2787.2642
-14.888
-7885.019431
There’s no heat transfer from the flash column: Q = Δ H = 0 Q = n ∆Hout - n∆Hin =(-7885.019431+ (-1656.91914)) – (-9541.938571)
= 0 kJ/hr
77
78
3.2.6
Energy balance for heat exchanger 3
Stream inlet at 1˚C datum at 25˚C CpdT (kJ/mol)
n H
C3H8
Flowrate (mol/hr) 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
Compound
79
4.8009
-4.008
-19.2420072
2787.2642
-14.888
-7885.019431
CpdT (kJ/mol)
n H
C3H8
Flowrate (mol/hr) 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
Stream outlet at 42˚C datum at 25˚C Compound
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) 80
Inlet (stream 2) Phase Component Pressure (bar) Temperature (oC) Total Molar Flow Rate (kmol/hr)
Vapor Mixture 1 600 2349.64
Outlet (stream 3) Phase Component
Vapor Mixture
Pressure (bar)
10
Temperature (oC)
600
Total Molar Flow Rate (kmol/hr)
2349.64
81
The outlet temperature of a stream by assuming the process is an ideal system. P
T 2 = T 1( P 2 )
(γ−1 γ )
1
For ideal system, γ = 1.3. Hence, T 2 = 6 00(10 1)
(1.3−1 1.3 )
= 1020.75˚C For energy balance, Q
ˆ = ∆ H = ∑ niHˆ i − ∑niH i in
out
= ∑ ni out
T
∫
298.15
C pdT
− ∑ ni in
T
∫
298.15
C pdT
Since there is no change component flow rate, 1293.9
1293.9
1293.9
298.15
298.15
298.15
Q = ∆ H = ni[
∫ 68.023×10−3 + 22.59×10 −5T − 13.11×10−8T 2 + 31.71×10− 12 T 3dT + ∫ 59.58×10−3 + 17.71×10−5T − 10.17×10−8T 2 + 24.6×10−1 2T 3 dT + ∫ 92.3×10−3 + 27.88×10 −5T −
= 2349.64 (175.44 + 144.36 + 227 + 197.87 + 278.61 + 30.52 ) = 2476050.63 kJ/hr 82
3.2.8
Separation Unit 3 (S3)
83
For Stream 7 Liquid stream
Stream temperature, T = 315.15 K and consider datum at 298.15 K Flow rate
Flow rate
(lbmol/hr)
(mol/hr)
Propane
386.86
1.75447 х 105
Butane
1.92
Pentane Propene
Component
∫C pdT
n∆H
(kJ/mol)
kJ/hr
0.10584
1.7993
315681.79
870.748
0.13367
2.2724
1978.69
4.80
2176.871
0.167
2.8390
6180.14
2376.43
10.777 x 105
0.1199
2.0383
2196675.91
C p
84
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 Flow rate
Flow rate
(lbmol/hr)
(mol/hr)
Propane
386.86
1.75447 х 105
Propene
2376.4
1077732.426
Component
∑
∫C pdT
n∆H
(kJ/mol)
kJ/hr
0.10584
2.0110
352823.92
0.1199
2.2781
2.45518 х 106
C p
2.8080 x 106
For Stream 9 85
Liquid stream
Stream temperature, T = 381.55 K and consider datum at 298.15 K Flow rate
Flow rate
(lbmol/hr)
(mol/hr)
Butane
1.9203
870.884
Butene
17.283
Pentane
4.80
Component
∑
∫C pdT
n∆H
(kJ/mol)
kJ/hr
0.13367
11.1481
9708.70
7838.095
0.09396
7.8363
61421.66
2176.871
0.1670
13.9278
30319.02
C p
101449.38
H = 101449.38+ 2.8080 x 106 - 2533034.15 = 376415.23kJ/hr
3.2.9
Separation Unit 4 (S4)
86
For Stream 8 Liquid stream
Stream temperature, T = 318.15 K and consider datum at 273.15 K
Component
Flow rate
Flow rate
(lbmol/hr)
(mol/hr)
C p
∫C pdT
n∆H
(kJ/mol)
kJ/hr
87
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
Propene
Flow rate
Flow rate
∫C pdT
n∆H
lbmol/hr
(mol/hr)
(kJ/mol)
(kJ/hr)
2376.43
10.77929 х 105
2.258
24.33963 х 105
88
24.33963 х 105
∑
For Stream 12 Liquid stream
Stream temperature, T = 320.15 K and consider datum at 273.15 K Component
Propane
Flow rate
Flow rate
(lbmole/hr)
(mol/hr)
386.95
1.75517 х 105
C p
0.10584
∫C pdT
n∆H
(kJ/mol)
(kJ/hr)
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 10 5 kJ/hr
89
3.2.10
Pump (P1)
Componen t
Molar Flow Rate (lbmol/hr)
Molar Flow Rate, F (kmol/hr)
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
Molar volume,v(kmol/m3) 21.9375 23.2486 0.09667 0.09048 0.11452
Fv 3849.50 25055.8 1 0.08410 0.70918 0.24942
∑Fv = 28906.35
90
Inlet Stream
Phase Pressure (bar) Temperature (oC)
Liquid 1 -47
Outlet Stream
Phase Pressure (bar) Temperature (oC) Q=
∑Fv ( P)
Q=
28906.35 х (18 -1) = 4.91408 х 105 kJ/hr
Liquid 18 -47
3.2.11 Energy Balance for Reactor
91
1. C3H8
C3H6 + H2
∆ ˆH r 1
2. C4H10
C4H8 + H2
∆ ˆH r 2
ˆ + nH ˆ + ∆ H ˆ Q1 = nΔ H = ∑nH R1 p1 r 1
(for Propane)
92
ˆ + n ˆH + ∆ ˆ Q2 = nΔ H = n ∑ nH Hr 2 R2 p2
(for Butane)
Q = Q1 + Q2
1. Energy balance for Propane Component Flow rate (mol/hr) Specific ˆ Enthalpy, H (kJ/mol)
ˆ H R1
ˆ H p1
ˆ ∆ H r 1
1253192.74
2330911.57
-
-91.8
184.72
-
93
∆H (kJ/hr)
-115043093.5
∑
430565985.2
140.21
315523031.9
2. Energy balance for Butane Component Flow rate (mol/hr) Specific ˆ Enthalpy, H (kJ/mol) ∆H (kJ/hr) ∑
ˆ H R2
ˆ H p2
ˆ ∆ H r 2
8708.84
16547.07
-
-119.0
240.61
-
-1036351.96
3981390.5
125.87
1945164.41
Q = Q1 + Q2
94
Q =
kJ kJ 315523031.9 hr + 1945164.41 hr
Q = 317.468 x 106 kJ/hr
95
3.3
Simulation Result from ASPEN HYSYS 3.3.1
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
96
3.3.2
Energy Balance
Equipment
Energy, kJ/h (HYSYS)
Q-HE00 Q-HE01 Q-HE02
5.76E+06 1.79E+08 1.65E+07 97
Q-Furnace Q-Compressor Q-Pump
3.4
9.23E+07 6.78E+07 6.90E+04
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 stream 2 stream 3 stream 4 stream 5 stream 6 stream 8 stream 9 stream 10 stream 11 stream 12
2971 2971 721.5 2400.4 2787.3 2787.3 5181 5181 5181 5181 2787
3179 3179 769.4 2410 2523 2523 4692 4692 4692 4692 2533
6.54 6.54 6.23 0.40 10.48 10.48 10.42 10.42 10.42 10.42 10.03 98
stream 13 stream 14 stream 15 stream 16 stream 17 stream 23 stream 24 stream 25
3.4.2
2393 2787 2787 165.4 2663.5 2376.4 124.5 124.5
2159 2533 2533 156.3 2377 2263 113.4 113.2
10.84 10.03 10.03 5.82 12.05 5.01 9.79 9.98
Energy Balance
Equipment
Energy, kJ/h (manual)
Energy, kJ/h (HYSYS)
% Diff
Q-HE00 Q-HE01 Q-HE02 Q-Furnace Q-Compressor Q-Pump
4.83E+06 1.80E+06 6.33E+04 1.00E+08 2.48E+06 4.91E+05
5.76E+06 1.79E+08 1.65E+07 9.23E+07 6.78E+07 6.90E+04
16.09 99.00 99.62 8.46 96.35 612.70
99
CHAPTER 4
HEAT INTEGRATION
4.1
PROCESS ENERGY INTEGRATION
∆Tmin = 10˚C
Table 4.1: Steam Table Data
100
Stream
Type
T supply (˚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
T (˚C)
0.049
865.70
∆T (˚C)
∑FCpC -∑FCpH (MW/K)
∆Hi (MW)
0
H1 284.4
-0.049
-13.9356
581.30
13.94 526.22
-0.001
-0.52622
55.08
14.46 6.23
48.85
1st Cascade
C2
0.052
0.32396 14.14
101
0.048
13.85
0.004
0.0554
35.00
14.08 10
25.00
0.032
13.76
C1 0.053
156.8
-131.80
-0.021
-3.2928 17.06
C3 0.028
10.3
-0.049
-142.10
-0.5047 17.56
Figure 4.1: 4.3
0.3200
Algorithm Table
Heat Exchanger Network
∆H ( MW) 1.5942
FCp(MW/K ) 0.053 50.08
25.5576
20
0.048 576.3
43.85
102
4.6704
0.028 30
49.3822
-136.8
0.049 870.7
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)
T H,out(˚C)
T C,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
-
-
103
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 104
(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
+
105
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 << 104800 Ib/hr
Propene production P1>> 100000 Ib/hr
106
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
107
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
108
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, D i
7.20 ft
Vessel Length, L
14.4 ft
Design Type
Vertical
Material of Contruction
Low- Alloy Steel SA-387B
109
6.2.2
Costing of Reactor
Cost of vessel, Cv = $ 40, 279 Cost of ladders and nozzles, C PL = $ 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 m 3/hr = 413.09 gpm Pump head, H =
Δ P (2.31) SG
=
Δ P = 356.82 ft ρ
110
6.3.2
Costing of Pump
Cost of pump, CP = $ 6577.78 Cost of motor, C P = $ 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
0.5 in Vertical 111
Material of Construction Material of insulation
Carbon steel 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
112
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
113
6.4.3
Sizing and costing of the condenser:
Parameters
S1
Length of tube
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
114
6.4.4
Sizing and costing of the re-boiler:
Parameters
Length of tube
S1
20 ft 115
Area of transfer,AR
3.146 ft2
Material of Construction
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
116
6.5
Compresssor
6.5.1
Main Sizing Parameters
Parameters
Compressor
Compressor Type
Centrifugal
Drive Type
Steam turbine
Material of Construction
Stainless steel
Inlet Volumetric Flow Rate, Q I
83283.83 ft3/min
117
6.5.2
Inlet Pressure, P I
14.5 psi
Outlet Pressure, P O
72.52 psi
Specific Heat Ratio, k
1.10
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 Design type Heat exchanger orientation Tube inlet direction Heat duty (kJ/s) Heat duty (Btu/hr)
Tin (˚C) Tout (˚C)
2 shell and 4 tubes Fixed Head Horizontal Horizontal 1594.2 5.44x10^6 Hot 870.7 349.12
Cold 43.85 576.3 118
119
From Figure 18.15 (a), F T = 0.85 and 2-4 exchanger is used. Ui = 235.5 Btu/oF.ft2.hr
Velocity of tube-side;
Cross section are/pass;
120
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
121
6.6.2
Costing of Heat Exchanger (HE2)
FBM = 3.17 F M = 1.08 + ( 86.34 100 )
0.5
= 2.01 FL = 1.25 (Tube length = 5.58 ft 2) 145.04 F P = 0.9803 + 0.018 (145.04 100 ) + 0.0017 ( 100 )
2
= 1.01 Fixed head: C B = e xp {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) 122
= $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, C TCI. Firstly, we need to prepare an equipment list, giving the equipment tittle, label, size, material of construction, design temperature, and design pressure.
123
Equipment Tittle
Label
Size
Material of Construction
Design Temp.
Design Pressure (bar)
Bare-module Cost, CBM
(˚C) Reactor
R1
V=586.17 ft3 Di = 7.20 ft
Low- Alloy Steel SA-387B
576
1
$ 243, 772
Cast Steel
-137.1
Pinlet = 10
$ 37,183
L= 14.4 ft Pump
P1
H =356.82 ft
Poutlet = 17.5 Distillation Column
S1
Compressor
C1
Heat Exchanger
D = 6.05 ft
Carbon steel
50
15
$ 984,036
Q = 83283.83 ft3/min
Carbon Steel
526
Pinlet = 1
$ 15,757,144
A = 86.34 ft2
Carbon Steel
t = 0.5 in
HE2
Poutlet = 10 870.7
10 CTBM ∑
$ 59,003 $17,081,138
124
After we get the value of total bare module cost, C TBM, we need to find the site development cost, Csite, building cost, C buildings, and offsite facilities cost, C offsite by assuming some factor. The calculation of total capital investment cost is shown below: facilities Assume it is grass-roots plant, the value fo C SITE is 10-20% of CTBM. Assume we take 15% of C TBM.
CSITE = 0.15 (17,081,138) CSITE = $ 2,562,170.75
Assume it is process buildings, the value of C BUILDINGS is 10% of CTBM
CBUILDINGS = 0.10 (17,081,138) CBUILDINGS = $ 1,708,113.80
The value of C OFFSITE 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) 125
CTPI = $ 39,969863.01
The value of C WC can be estimated 17.6% of C TPI
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
126
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 152,764,816.40 RM 167,327.77/ hr
1 month = 912.967 hr × 124day hr × 30 days
= 1 month 9 days
127
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
128
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.
129
In a nutshell, after we had done a simulation, optimization and process integration, our total capital investment is $ 47,004,558.90.
130
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, m 1 = 2376.43 x 44.1 131
= 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 H 2 , m3 = (no. of mole) X (molecular weight) = (2376.43) x (2.01) = 4776.62 lb/hr
Mass flowrate feed to the reactor, F = m 1 + R = 104800.56 + 17060.53 = 121861.09lb/hr
2. Separation Unit 1
Overall mass balance : F1 = D1 + B1
132
From the overall mass balance, we know that the mass flowrate of propane at D 1 is 104800.56 lb/hr and the composition is assume 0.990 of C 3H8, so 0.990 D1 = mass flow rate of C 3H8 D1 = mass flow rate of C 3H8 / 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
133
3. Mixer
Overall mass balance : F = D1 + R
where R is the recycle of propane
from previous calculation, the value of D 1 = 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.
134
(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
135
= 2393.7 x 2.01 = 4811.36 lb/hr
136
APPENDICES B
CALCULATION OF ENERGY BALANCE Sample Calculation for Energy Balance
Molar Flow Rate for Propane 1kgmol
1000mol n = 2763.3 lbmol hr x 2.205lbmol x 1kgmol
= 1253197.28mol/hr
For reaction of propane H R1 = H R1 =
298 K
∫
C p(reactant )dt
873 K
298 K
∫ (68.023×10−3 + 22.59×10 −5T − 13.11×10−8T 2 + 31.71×10− 12 T 3 )dt
873 K
H R1 =− 91.81kJ /mol H R1 = H p1 =
298 K
∫
C p( product )dt
873 K 873 K
∫ (68.023×10−3 + 22.59×10 −5T − 13.11×10−8T 2 + 31.71×10−1 2T 3 )dt + (59.58×10−3 + 17.71×10 −5T − 10.17×10−8T 2 + 24.6×10−1 2T 3 ) dt + (28.84×10−3 + 0.00765×10 −5T + 0.3288×1
298 K
137
H p1
= 91.79kJ /mol + 75.75kJ /mol + 17.18kJ /mol
H p1 =
184.72kJ/mol
ˆ = ∑v ∆ H ˆ = (1)(∆ H ˆ ) + (1)(∆ H ˆ ) − (1)(∆ H ˆ ) ∆ H r 1 i f f C 3 H 6 f H 2 f C 3 H 8 ˆ = 20.41 ∆ H kJ + 0 − r 1 mol
(−119.8)kJ mol
= 1 40.21 kJ /mol
138
APPENDICES C
ASPEN HYSYS
139
140
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˚
141
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
142
Q=FCpC∆T 4.6704 = 0.028 (T—136.8) T = 30˚C
143
APPENDICES E
CALCULATION OF SIZING AND COSTING
REACTOR
Q = 3517.02 ft 3/hr Retention time =5 min at half full : ×1 hr ×2 ) = 586.17 ft 3 Volume, V = (3517.02 ft3/hr) × ( 5min 60 min 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) 144
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 m 3/hr = 413.09 gpm
145
H=
Δ P (2.31) SG
=
Δ P ρ
= 108.78 psi x
1lb/in2 1 psi
x
ft 3 43.9 lb
x
144 in2 1 ft 2
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 H ρ 33000
= 413.09 gal min x 356.82 ft x
43.9 lb ft 3
x
0.1334 ft 3 1 gal
x
1 33000
lb. ft min
= 26.16
ln Q = 6.024 ηp = -0.316 +0.24015 (6.024) – 0.01199(6.024)2 = 0.6956 PB = PT η p =
26.16 0.6956
= 37.61
lb. ft min
ln PB = 3.627 146
ηm = 0.80 + 0.0319(3.627) – 0.00182(3.627)2 = 0.892 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 147
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 4( 425100 )
3600 DT= 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 (5.582)(6.05)(12) tp= 2(15000)(0.85)−(1.2 =0.0158 < tabulated data thus t p=0.375 in *5.582)
148
ts=0.375+0.125=0.5 in
W= π (Di+ts) (L+0.8Di) ts ρ= 21458 Ib Cv= $ 67436 CPL= 300.9 * 6.050.63316 500.8016 = $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 ft 3/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 149
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: 150
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=1 151
CB= $ 65325 With bare-module $ 3.17(65325) = $ 207080
COMPRESSOR (a)
Preliminary estimate of brake horsepower, P B
Inlet volumetric flow rate, Q I = 83,283.83 ft3/min Inlet pressure, P I
= 14.5 psi
Outlet pressure, PO
= 72.52 psi
Specific heat ratio, k
= 1.10
Mechanical efficiency, η B
= 0.4952
1.10 83283.83×14.5×(( 72.52) P B = 0.00436×( 1.10−1)× 0.4952 14.5
(b)
1.10−1 1.10
− 1) = 18430.94 BHp
Purchase cost of compressor
152
Assumption: Drive efficiency, η C = 0.75 Material factor, F M
= 1.00 (carbon steel)
Drive type factor, F D = 1.15 (steam turbine) P C = 18430.94 0.75
= 24574.59 Hp
Base purchase cost, C B = e xp{7.5800 + 0.80[ln ( P C )
]}
C B = e xp{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
153
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 870.7
Cold 43.85 154
Tout (˚C)
349.12
576.3
155
From Figure 18.15 (a), F T = 0.85 and 2-4 exchanger is used. Ui = 235.5 Btu/oF.ft2.hr
Velocity of tube-side;
Cross section are/pass;
156
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) 157
FBM = 3.17 F M = 1.08 + ( 86.34 100 )
0.5
= 2.01 FL = 1.25 (Tube length = 5.58 ft 2) 145.04 F P = 0.9803 + 0.018 (145.04 100 ) + 0.0017 ( 100 )
2
= 1.01 Fixed head: C B = e xp {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) = $59,003 158
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