UNIVERSITY TECHNOLOGY MARA FACULTY OF CHEMICAL ENGINEERING ENGINEERING
PRODUCTION OF PROPENE
SYAZWINA BINTI AZIZI
2015635202 20156352 02
SITI KHAIRUNNISA BINTI AHMAD
2015655912 20156559 12
HARRIS SHAZLAN BIN MUHAMMAD SHAZWAN
2015635236 20156352 36
NINA SUHAILA BINTI NOR WAHID
2015237884 20152378 84
SABRINA BINTI KHAMIS
2015687078 20156870 78
1
TABLE OF CONTENT Topic 1.0 Introduction 1.1 Market Survey 1.2 Location 1.3 Process Selection 1.4 Process Modification 2.0 Mass Balance 2.1 Basic Mass Balance 2.2 Separator 1 2.3 Mixer 2.4 Reactor 2.5 Separator 2 2.6 Separator 3 2.7 Separator 4 3.0 Energy Balance 3.1 Heat Exchanger 1 3.2 Separator 2 3.3 Furnace 3.4 Reactor 3.5 Heat Exchanger 2 3.6 Separator 2 3.7 Heat Exchanger 3 3.8 Separator 3 3.9 Condenser 3.10 Separator 4 4.0 Heat Integration 5.0 Cost Calculation 5.1 Bare Module Costs 5.2 Fixed Capital Investment 5.3 Cost of Manufacturing 5.4 Revenue 5.5 Cost of Land 5.6 Working Capital 6.0 Profitability Analysis 6.1 Non-discounted Profitability 6.2 Discounted Profitability 7.0 Safety Analysis 8.0 Waste Management 9.0 References
Page Number 3 4 7 8 12 13 13 15 16 17 18 19 20 21 23 24 25 26 27 28 29 30 31 32 33 35 35 49 49 54 54 54 55 55 57 60 64 66
2
TABLE OF CONTENT Topic 1.0 Introduction 1.1 Market Survey 1.2 Location 1.3 Process Selection 1.4 Process Modification 2.0 Mass Balance 2.1 Basic Mass Balance 2.2 Separator 1 2.3 Mixer 2.4 Reactor 2.5 Separator 2 2.6 Separator 3 2.7 Separator 4 3.0 Energy Balance 3.1 Heat Exchanger 1 3.2 Separator 2 3.3 Furnace 3.4 Reactor 3.5 Heat Exchanger 2 3.6 Separator 2 3.7 Heat Exchanger 3 3.8 Separator 3 3.9 Condenser 3.10 Separator 4 4.0 Heat Integration 5.0 Cost Calculation 5.1 Bare Module Costs 5.2 Fixed Capital Investment 5.3 Cost of Manufacturing 5.4 Revenue 5.5 Cost of Land 5.6 Working Capital 6.0 Profitability Analysis 6.1 Non-discounted Profitability 6.2 Discounted Profitability 7.0 Safety Analysis 8.0 Waste Management 9.0 References
Page Number 3 4 7 8 12 13 13 15 16 17 18 19 20 21 23 24 25 26 27 28 29 30 31 32 33 35 35 49 49 54 54 54 55 55 57 60 64 66
2
1.0 INTRODUCTION Propene, also known as propylene or methyl ethylene, is an unsaturated organic compound having the chemical formula C3H6. It has one double bond, and is the second simplest member of the alkene class of hydrocarbons.
At room temperature and atmospheric pressure, propene is a gas, and as with many other alkenes, it is also colorless with a weak but unpleasant smell. Propene has a higher density and boiling point than ethylene due to its greater mass. It has a slightly lower boiling point than propane and is thus more volatile. It lacks strongly polar bonds, yet the molecule has a small dipole moment due to its reduced symmetry (its point group is Cs). Propene has the same empirical formula as cyclopropane but their atoms are connected in different ways, making these molecules structural isomers. Propene is produced from fossil fuels which are petroleum, natural gas, and, to a much lesser extent, coal. Propene is a by-product of oil refining and natural gas processing. During oil refining, ethylene, propene, and other compounds are produced as a result of cracking larger hydrocarbon molecules to produce hydrocarbons more in demand. A major source of propene is naphtha cracking intended to produce ethylene, but it also results from refinery cracking producing other products. Propene can be separated by fractional distillation from hydrocarbon mixtures obtained from cracking and other refining processes; refinery-grade propene is about 50 to 70%. Propene 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 The raw material that is utilized in this process is liquefied petroleum gas (LPG). The term LPG actually encompasses more than one variety of gaseous fuel. There are a number of hydrocarbon gases that fall into the classif 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. 3
Propane dehydrogenation (PDH) is used to produce polymer-grade propylene from propane. The productions of dehydrogenation process normally consist of isobutylene, nbutenes, or propylene from isobutane, n-butane, or propane respectively.
The process
consists of a reactor section, product recovery section and catalyst regeneration section. Hydrocarbon feed is mixed with hydrogen-rich recycle gas and is fed into a heater to be heated to over 540°C (1000°F) and then enters the reactors to be converted at high monoolefin selectivity. Several inter-stage heaters are used to maintain the reaction through supplying heat continuously, since the reaction is endothermic. This process converts propane to propylene on a fixed-bed of chromia-alumina catalyst in order to recycle the unconverted propane and leave propylene as the net product. The chemical reaction occurs during the gaseous phase at high temperatures (ranging from 540 to 820 degrees C) that causes
an
endothermic
equilibrium
reaction.
Conversion
rates
depend
on
the
thermodynamic limits from specific partial pressures and the actual temperature used in the immediate process. Low partial pressures of the hydrocarbons eventually create the dehydration reaction as well as the high temperatures. The propane through the catalytic dehydration is processed in dehydrating reactors to create the by-products and compressors separate these substances later on.
1.1 Market Survey
4
The
dominant
outlet
for propylene is polypropylene (PP), accounting for nearly two-thirds of global propylene consumption. Second largest sales market
for
of propylene
propylene oxide,
is
the
followed
production by
the
manufacturing of acrylonitrile, butyraldehyde, cumene, and acrylic acid. The derivatives of propylene oxide are needed as raw materials for
the
manufacturing
like polyurethane
(PUR),
of
products
lacquers and adhesives, polyester
resins,
cooling
agents,
antifreeze, and solvents. In addition to the production of the engineering plastic acrylonitrile butadiene styrene (ABS), acrylonitrile is also used to produce acrylic fibers that are then processed into textiles. Butyraldehyde is an intermediate in the production of butanol and 2ethylhexanol. Cumene is primarily used to produce phenol and acetone. Acrylic acid is used in the production of acrylic esters for superabsorbers (superabsorbent polymers, SAP) Besides traditional propylene production methods, new on-purpose technologies have penetrated the market. On-purpose sources include propane dehydrogenation (PDH), metathesis, high severity catalytic cracking, olefin cracking, Coal-to-Olefins (CTO) and Methanol-to-Olefins (MTO). On-purpose production currently makes up an estimated 14% of global propylene supply, and is projected to grow to 30% of global supply by 2019.
Figure 2.1: Propylene production by technology
Global on-purpose propylene production grows from 14% to 30% of supply over the next five years as the wave of new propane dehydrogenation (PDH) and coal-to-olefin (CTO)
5
projects start production. The majority of the new on-purpose propylene capacity is being built in China via both CTO and PDH. Based on the recent article, in the Global Propylene Market 2016 stated that propylene is the second most important product of the olefin market (PR Newswire). In 2014, about 90 million tonnes of propylene were processed worldwide and it was estimated the demand to rise to 130 million tonnes worldwide by year 2024 according to the latest data from HIS Chemical North American Propylene Supply Study. Some of the largest demands countries in propylene production are India, China and Japan which together accounted for about 55% of global consumption in 2014. It is anticipated that China will grow its propylene
capacity at the fastest pace of any country in the world over the next five years and the market will exhibit significant growth and account for about 76% of the total increase in propylene demand.
Figure 2.1: Graph of propylene supply and demand from year 2009 to 2024
Currently,
China
is
the
second-largest
producer of propylene and is expected to make up for almost 70% of the Asian capacity addition and 45% of the global capacity addition in the next five years. As a result of being a by-product of refinery sources, the availability of propylene has been widespread. The by far largest share of global propylene output, about two 6
Figure 2.2: Pie chart shows world consumption of propylene
thirds, is processed into polypropylene. Demand for this plastic is projected to increase by 3.7% p.a. until 2021 and will therefore dominate demand development on the propylene market.
7
1.2 Location Scoring method Criteria Transportation Human Capital Labor availability Skill levels Source of Raw Materials Market Target Infrastructure Politics Weather Landscape Total Score Rank
China 5
Location India 3
Weightage (%) Malaysia 5
10
5
5
3
10
4
4
3
20
5 5 4 3 4 435 1
3 3 2 4 4 350 3
3 4 3 4 4 375 2
10 10 15 5 20 100
Site Location
This plant located at Pudong, Shanghai, China where these locations are the most nearer to the supplier of raw material which is propane from Shanghai Hynote Import & Export, which take only 20 min to reach. This location also strategic where the products produce can be easily import and export by shipping at Shanghai Port which are huge port in China. Lastly labour cost also low compared to other country which supplied raw material. 8
1.3 Process Selection
Process 1 (Hydro MTO Process) MTO process The MTO process converts crude methanol to olefins, which results in savings for a methanol purification section. The reaction takes place on a zeolite catalyst in a fluidized-bed reactor and MTO can use crude methanol. The catalyst is deactivated over time by the building of coke, which is why a portion of the spent catalyst is continuously removed from the reactor to a regeneration reactor. Air or oxygen is introduced to that regeneration reactor so that the catalyst can be regenerated by building carbon oxides. The product stream leaving the reactor is fed to a separation section to remove water and to recover non-reacted dimethyl ether (DME). The olefin rich stream is then passed to a fractionation section that separates the mixture into the desired ethylene and propylene streams as well as a fuel gas stream and a stream that consists of medium boiling hydrocarbons. The MTO uses a fluidized-bed reactor, which is favourable with respect to the exothermic reaction so that the released heat can be removed faster. Moreover, the spent catalyst can be regenerated easily by continuously removing a part of the catalyst from the reactor and reroute it back after regeneration. However, the fluidized-bed implies the presence of a stable catalyst, which had to be developed for these special conditions. According to the needs, the ratio between propylene and ethylene produced by the reaction can be adjusted in a range of about 1.3 to 1.8. The heavier hydrocarbon stream leaving fractionation is fed into a cracking zone to provide another source for ethylene and propylene 9
production. The product stream of the cracking section is separated into high boiling hydrocarbons, which are removed from the process and an olefin rich stream which is rerouted to fractionation.
Process 2 (Metathesis of Propylene) Metathesis process The metathesis of ethylene and butane to propylene is carried on a metal oxide catalyst and is operated with excess ethylene. The process consists of a reaction section and product recovery section. Fresh ethylene and butane are fed to a treater before it is sent to the heater to be heated to the operating temperature of 453K. Here the feed reacts to form propylene and other byproducts. Both propylene and unreacted ethylene travels to a depropylenizer column and deethylenizer column to be separated. The Propylene is harvested and unreacted ethylene is sent back to the feed treater to be mixed with fresh feed for higher conversion. Theoretically, if 1-butene was used it will have a 100% selectivity for propylene
10
Process 3 (Dehydrogenation of Propane)
Propane dehydrogenation Propane dehydrogenation (PDH) is used to produce polymer-grade propylene from propane. The productions of dehydrogenation process normally consist of isobutylene, nbutenes, or propylene from isobutane, n-butane, or propane respectively.
The process
consists of a reactor section, product recovery section and catalyst regeneration section. Hydrocarbon feed is mixed with hydrogen-rich recycle gas and is fed into a heater to be heated to over 540°C (1000°F) and then enters the reactors to be converted at high monoolefin selectivity. Several inter-stage heaters are used to maintain the reaction through supplying heat continuously, since the reaction is endothermic. This process converts propane to propylene on a fixed-bed of chromia-alumina catalyst in order to recycle the unconverted propane and leave propylene as the net product. The chemical reaction occurs during the gaseous phase at high temperatures (ranging from 540 to 820 degrees C) that causes
an
endothermic
equilibrium
reaction.
Conversion
rates
depend
on
the
thermodynamic limits from specific partial pressures and the actual temperature used in the immediate process. Low partial pressures of the hydrocarbons eventually create the dehydration reaction as well as the high temperatures. The propane through the catalytic dehydration is processed in dehydrating reactors to create the by-products and compressors separate these substances later on.
11
Catalyst activity is maintained by the continuous catalyst regenerator (CCR) or by shutting down reactors one by one and regenerating the reactor by using regeneration air. In the continuous catalyst regenerator, catalyst is continuously withdrawn from the reactor, regenerated, and then fed back into the reactor bed. The reactor effluent is compressed, dried and sent to a cryogenic separator where hydrogen is recovered. The olefin product is sent to a selective hydrogenation process (SHP) where dienes and acetylenes are saturated to mono-olefins. The propylene stream goes to a deethanizer where light-ends are removed prior to the propane-propylene (P-P) splitter. Unconverted feedstock is recycled back to the depropanizer where it combines with fresh feed before being sent back to the reactor section.
SCREENING
Processes Criteria
MTO
Metathesis
Dehydrogenation
Standard
Capital Cost
0
+
-
0
Cost of
+
0
+
0
+
0
0
0
2
1
0
0
Reactants Environmental Impact Total
Advantages
Disadvantages
Cheap material cost
Reaction is highly endothermic
High propylene selectivity
Requires multiple reactors
Produces other profitable products
Catalyst is easily deactivated
(benzene)
12
1.4 Process Modification
(Modified Dehydrogenation Process)
In the modification of the dehydrogenation of propane we added 3 more reactors and heaters to the original 1 to increase the conversion of propane to its products. At the end of the reaction process, we also added a heat exchanger to act as a preheater for the propane feed that is about to enter the first heater as this would reuse the heat released by the products. The heat exchanger also reduces the temperature of the product so that it would not damage the compressor equipment that is designed to withstand temperatures below 100 degree Celsius only.
13
2.0 MASS BALANCE 2.1 Basic Mass Balance
Temperature = 500ºC
F lb/hr C3H8
Pressure = 1 Bar
R lb/hr C3H8
C2H8 Stoichiometry Mass Flowrate(lb/hr) Molecular Weight (lb/lb mol) n2, (lbmole/hr)
C3H6
H2
1
1
1
m1
m2 = 110 000
m3
44.10
42.08
2.01
2614.07
2614.07
2614.07
→
Basis = 110 000 lb/hr of Propylene produced Conversion, X = 0.84
Number of moles of Propylene formed
ℎ 110 000 = 42.08 = 2614.07 ℎ Assuming 100% conversion was achieved, m 1
= 2 614.07×44.1= 115 280.49ℎ 14
Mass flowrate of recycle, R
= 1− × = 1−0.84 ×115 280. 4 2 0.84 ℎ =21 903.28 ℎ Mass flowrate of H 2, m3
= × = 2 614.07 ℎ ×2.01 =5 254.28 ℎ Mass flowrate of feed to reactor
= + = 115 280.42 ℎ +21 903.42 ℎ = 137 183.84ℎ
15
2.2 Mass Balance For Separation Unit 1
2 D1=126 920.68 lb/hr 0.910 Propane 0.007 Butane 0.002 Pentane S-1 1 F1=145 587.87 lb/hr 0.79 Propane 0.18 Butane 3
0.03 Pentane
B1=18 666.99 lb/hr 0.87 Butane 0.13 Pentane
Stream 1 No
Component
1
Stream 2
Stream 3
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Propane
0.79
115 014.42
0.91
115 014.42
0
0
2
Butane
0.18
26 205.82
0.07
8 929.69
0.87
17 275.83
3
Pentane
0.03
4 367.64
0.02
2 976.56
0.13
1 391.06
TOTAL
F1=145 587.87 lb/hr
D1=126 920.68 lb/hr
16
B1=18 666.99 lb/hr
2.3 Mass Balance For Mixer
2
D1=126 920.68 lb/hr
3
M-1
F =148 828.19 lb/hr
0.91 Propane
A1 Propane
0.07 Butane
A2 Butane A3 Pentane
19
0.02 Pentane
R =21 907.51 lb/hr 1.0 Propane Stream 2 No
Component
1
Stream 3
Stream 19
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Propane
0.91
115 014.42
A1=0.92
136 921.93
1
21 907.51
2
Butane
0.07
8 929.69
A2=0.06
8 929.69
0
0
3
Pentane
0.02
2 976.56
A3=0.02
2 976.56
0
0
TOTAL
D1=126 920.68 lb/hr
F=148 828.19 lb/hr
17
R=21 907.51 lb/hr
2.4 Mass Balance For Reactor Unit 1 (R-1)
R-1
4
5
W= 148 828.19 lb/hr
Y= 148 828.19 lb/hr
0.92 Propane
N1 Propane
0.06 Butane
N2 Butane
0.02 Pentane
N3 Pentane N4 Propene N5 Butene N6 Hydorgen
Stream 4 No
Component
1
Stream 5
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Propane
0.92
136 921.93
N1=0.15
21 907.51
2
Butane
0.06
8 929.69
N2=0.006
892.97
3
Pentane
0.02
2 976.56
N3=0.02
2 976.56
4
Propene
-
-
N4=0.74
110 000
5
Butene
-
-
N5=0.052
7 796.87
6
Hydrogen
-
-
N6=0.035
5 254.28
TOTAL
W =148 828.19 lb/hr
18
Y =148 828.19 lb/hr
2.5 Mass Balance For Separation Unit 2
9 D2 = 143 573.91 lb/hr 1.0 Hydrogen
S-2 7 F2=148 828.19 lb/hr 0.15 Propane 0.006 Butane 8
0.02 Pentane
B2 = 5 254.28 lb/hr 0.74 Propene
0.15 Propane
0.052 Butene
0.006 Butane
0.035 H dro en
0.021 Pentane 0.77 Propene 0.054 Butene
Stream 7 No
Component
Stream 8
Stream 9
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
1
Propane
0.15
21 907.51
0.15
21 907.51
0
0
2
Butane
0.006
892.97
0.006
892.97
0
0
3
Pentane
0.02
2 976.56
0.021
2 976.56
0
0
4
Propene
0.74
110 000
0.77
110 000
0
0
5
Butene
0.052
7 796.87
0.054
7 796.87
0
0
6
Hydrogen
0.035
5 254.28
0
0
1.0
5 254.28
TOTAL
F2 = 148 828.19 lb/hr
B2 = 5 254.28 lb/hr
19
D2 = 143 573.91 lb/hr
2.6 Mass Balance For Separation Unit 3
11 D3 = 131 907.51 lb/hr 0.166 Propane 0.834 Propene S-3 10
F3 = 143 573.91 lb/hr
0.15 Propane 0.006 Butane 12
0.021 Pentane
B3 = 11 666.4 lb/hr 0.77 Propene
0.08 Butane
0.054 Butene 0
0.26 Pentane
H dro en
0.67 Butene
Stream 10 No
Component
Stream 11
Stream 12
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
1
Propane
0.15
21 907.51
0.166
21 907.51
-
-
2
Butane
0.006
892.97
-
-
0.08
892.97
3
Pentane
0.021
2 976.56
-
-
0.26
2 976.56
4
Propene
0.77
110 000
0.834
110 000
-
-
5
Butene
0.054
7 796.87
-
-
0.67
7 796.87
6
Hydrogen
0
0
-
-
-
-
TOTAL
F3 = 143 573.91 lb/hr
D3 = 131 907.51 lb/hr
20
B3 = 11 666.4 lb/hr
2.7 Mass Balance For Separation Unit 4
20 D4= 110 000 lb/hr 1.0 Propene
S-4 15
F4= 131 907.51 lb/hr
0.166 Propane 0.834 Propene 17 B4 = 21 907.51lb/hr 1.0 Propane
Stream 15 No
Component
Stream 20
Stream 17
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
Mole Fraction
Mass Flowrate (lb/hr)
1
Propane
0.166
21 907.51
0
-
1.0
21 907.51
2
Propene
0.834
110 000
1.0
110 000
0
-
TOTAL
F4= 131 907.51 lb/hr
D4= 110 000lb/hr
21
B421 907.51lb/hr
3.0 ENERGY BALANCE
Reference = (liquid phase, 25˚C, 1bar)
Table 3.1: Table of Data for Heat of Capacities
Cp=A+B*T+C*T^-2+D*T^-3 Compound
Molecular Weight
∆Hf
∆Hv Kj/mol
A*10^3
B*10^5
C*10^8
D*10^12
Average Cp for liquid 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
22.306
92.3
27.88
-15.47
34.98
0.13367
Cp=A+B*T+C*T^-2+D*T^-3
22
Butene
56.1
1.17
21.916
82.88
25.64
-17.27
50.50
0.09396
Pentane
72.15
-146.4
25.77
114.8
34.09
-18.99
42.26
0.167
Hydrogen
2.016
0
0.904
28.84
0.00765
0.3288
-0.8698
-
Formula
=− =∆+∆ +∆ : 1. ℎ , =0 2. ℎ , ∆ = 0 3.ℎ ℎ ℎ , ∆ = 0 ℎ: =∆ =̇ ∆̂ −̇ ∆̂ ̇ ∆̂ −̇ ∆̂ =∆º+ 23
∆º = ∣ ∣ ∆̂ − ∣ ∣ ∆̂
3.1 Heat Exchanger (E-1)
Stream inlet (liq) T=293.15K
H1
H4
H2
H5
H3
H6
Stream inlet (liq) T=325.15K
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (liq)
1 182 354.66
H1 =-625 702.09
1182354.66
H4=3 378 791.27
Butane (liq)
204 520.26
H2=-136 691.12
204520.26
H5=754 698.17
Pentane (liq)
27 460.48
H3=-22 929.50
27460.48
H6=123 819.30
24
∆H=-78
522.71
∆H=4
257 308.74
Q=4 335 831.45
3.2 Seperator Unit 1 (S-1)
2 H9
Stream outlet (vap) T=320.15K
H10 H11
S-1 Stream inlet (liq) T=325.15K
1 H4 H5 H6
3(i) H7 H8
REFERENCE (LIQ,298.15K)
25
Stream outlet (liq) T=293.15K
Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (liq)
1182354.66
H4=3 378 791.27
-
-
Butane (liq)
204520.26
H5=754 698.17
134 825.80
H7=1 544 499.51
Pentane (liq)
27460.48
H6=123 819.30
8 745.26
H8=125 161.29
Propane (vap)
-
-
1 183 254.66
H9=26 554 956.1
Butane (vap)
-
-
69 689.93
H10=1 854 086.65
Pentane (vap)
-
-
18 715.22
H11=558 741.02
∆H=4257308.74
∆H=30
637 444.57
Q=26 380 135.83
3.3 Furnace (H-1)
Stream inlet (vap) T=320.15K
3
4
H9
H12
H10
H13
H11
H14
Stream outlet (vap) T=873.15
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (vap)
1 183 254.66
H9=26 554 956.1
1 183 254.66
H12=129 312 751.8
Butane (vap)
69 689.93
H10=1 854 086.65
69 689.93
H13=8 293 798.57
Pentane (vap)
18 715.22
H11=558 741.02
18 715.22
H14=140 346 458.50
∆H=28
967 783.77
∆H=277
Q=248 985 225.10 kJ/hr 26
953 008.90
3.4 Reactor Unit 1 (R-1)
R-1
Stream inlet (vap) T=873.15K
5
4 H12
H15
H13
H16
H14
H17 H18 H19 H20
Heat of Reaction (kJ/mol) C3H4 --- C3H6 + H2
124.21
27
Stream outlet (vap) T=873.15K
C4H10 ---- C4H8 + H2
125.87
Extent reaction (mol/hr)
1 183 204.77
295 895 848.90 kJ/hr
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (vap)
1 183 254.66
H12=129 312 751.80
225 430.87
H15=25 129 004.51
Butane (vap)
69 689.93
H13=8 293 798.57
6 967.18
H16=989 304.72
Pentane (vap)
18 715.22
H14=140 346 458.50
18 715.22
H17=3 215 057.64
Propene (vap)
-
-
1 185 722.21
H18=109 892 734.40
Butane (vap)
-
-
63 040.27
H19=8 061 337.57
Hydrogen (vap)
-
-
1 182 193.26
H20=21 580 937.96
∆H=277
953 008.90
∆H=168
868 376.80
Q= 186 811 216.80 kJ/hr
3.5 Heat Exchanger (E-2)
Stream inlet (vap) T=873.15K
6
7 H21
H15
Stream outlet (liq) T=136.05K
H22
H16
H23
H17
H24
H18
H25
H19
H26
H20 REFERENCE (LIQ, 298,15K) Component
Inlet stream
Outlet stream
28
n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (vap)
225 430.87
H15=25 129 004.51
-
-
Butane (vap)
6 967.18
H16=989 304.72
-
-
Pentane (vap)
18 715.22
H17=3 215 057.64
-
-
Propene (vap)
1 185 722.21
H18=109 892 734.40
-
-
Butane (vap)
63 040.27
H19=8 061 337.57
-
-
Hydrogen (vap)
1 182 193.26
H20=21 580 937.96
-
-
Propane (liq)
-
-
225 430.87
H21= -3 868 393.73
Butane (liq)
-
-
6 967.18
H22=-150 976.62
Pentane (liq)
-
-
18 715.22
H23=-506 621.0
Propene (liq)
-
-
1 185 722.21
H24=-23 050 439.76
Butane (liq)
-
-
63 040.27
H25=-960 161.06
1 182 193.26
H26=-5 520 842.52
Hydrogen (liq) ∆H=168
868 376.80
∆H=-34
057 434.69
Q= -202 925 810.80 kJ/hr
3.6 Separation Unit 2 (S-2)
9 H27
Stream inlet (liq) T=136.05K
Stream outlet (vap) T=136.05K
S-2 7 H21 H22 8
H23 H24
H21
H25
H22
H26
H23 29
H24 H25
Stream outlet (liq) T=136.05K
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (liq)
225 430.87
H21=-3 868 393.73
225 430.87
H21=-3 868 393.73
Butane (liq)
6 967.18
H22=-150 976.62
6 967.18
H22=-150 976.62
Pentane (liq)
18 715.22
H23=-506 621.0
18 715.22
H23=-506 621.0
Propene (liq)
1 185 722.21
H24=-23 050 439.76
1 185 722.21
H24=-23 050 439.76
Butane (liq)
63 040.27
H25=-960 161.06
63 040.27
H25=-960 161.06
Hydrogen (liq)
1 182 193.26
H26=-5 520 842.52
-
-
Hydrogen (vap)
-
-
1 182 193.26
H27=-29 740 435.84
∆H=-34
057 434.69
∆H=-58
277 028.01
Q= -24 219 593.32 kJ/hr
3.7 Heat Exchanger (E-3)
Stream inlet (liq) T=136.06K
9i H21
H28
H22
H29
H23
H30
H24
H31
H25
H32
Stream outlet (liq) T=315.15K
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol) 30
n (mol/hr)
H (kJ/mol)
Propane (liq)
225 430.87
H21=-3 868 393.73
225 430.87
H28=405 550.14
Butane (liq)
6 967.18
H22=-150 976.62
6 967.18
H29=15 815.50
Pentane (liq)
18 715.22
H23=-506 621.0
18 715.22
H30=53 132.51
Propene (liq)
1 185 722.21
H24=-23 050 439.76
1 185 722.21
H31=2 416 501.86
Butane (liq)
63 040.27
H25=-960 161.06
63 040.27
H32=100 675.31
∆H=-28
536 592.17
∆H=2
991 675.32
Q= 31 528 267.49 kJ/hr
3.8 Separation Unit 3 (S-3)
11
H33
Stream outlet (vap) T=315.15K
H34 Stream inlet (liq) T=315.15K
S-3 10 H28 H29 H30
12
H31
H29
H32 H30 H32 31
Stream outlet (liq) T=315.15K
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (liq)
225 430.87
H28=405 550.14
-
-
Butane (liq)
6 967.18
H29=15 815.50
6 967.18
H29=15 815.50
Pentane (liq)
18 715.22
H30=53 132.51
18 715.22
H30=53 132.51
Propene (liq)
1 185 722.21
H31=2 416 501.86
-
-
Butane (liq)
63 040.27
H32=100 675.31
63 040.27
H32=100 675.31
Propane (vap)
-
-
225 430.87
H33= 4 903 346.85
Propene (vap)
-
-
1 185 722.21
H34=22 188 419.72
∆H=2
991 675.32
∆H=27
261 389.89
Q= 24 269 714.57 kJ/hr
3.9 Condenser (E-15)
Stream inlet (vap) T=315.15K
E-15
15
Stream outlet (liq) T=313.15K
11 H33
H35
H34
H36
REFERENCE (LIQ, 298,15K) Inlet stream
Outlet stream
Component Propane (vap)
n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
225 430.87
H33= 4 903 346.85
-
-
32
Propene (vap)
1 185 722.21
H34=22 188 419.72
-
-
Propane (liq)
-
-
225 430.87
H35=357 894.05
Propene (liq)
-
-
1 185 722.21
H36=2 132 521.395
∆H=27
261 389.89
∆H=2
490 415.45
Q= -25 130 974.44 kJ/hr
3.10 Separation Unit 4 (S-4)
20
Stream outlet (vap) T=313.15K
H37
S-4 Stream inlet (vap) T=313.15K
15 H35 H36 17 H35
REFERENCE (LIQ, 298,15K) 33
Stream outlet (liq) T=313.15K
Inlet stream
Outlet stream
Component n (mol/hr)
H (kJ/mol)
n (mol/hr)
H (kJ/mol)
Propane (liq)
225 430.87
H35=357 894.05
225 430.87
H35=357 894.05
Propene (liq)
1 185 722.21
H36=2 132 521.395
-
-
Propene (vap)
-
-
1 185 722.21
H37=34 361 043.92
∆H=2
490 415.45
∆H=34
718 937.97
Q= 32 228 522.52 kJ/hr
4.0 HEAT INTEGRATION
∆T =10 K Stream
Type
F (kg/s)
Cp (KJ/kg.K)
FCp (kW/K)
Tsupply (K)
Ttarget (K)
Q (kW)
H1
Hot
18.75
0.331
6.21
873.15
683.15
1 179.19
H2
Hot
18.75
0.331
6.21
683.15
493.15
1 179.19
H3
Hot
18.75
0.331
6.21
493.15
349.10
958.19
H4
Hot
18.75
0.331
6.21
349.10
136.05
1 258.07
C1
Cold
26.65
0.557
14.84
320.15
399.57
1 179.19
C2
Cold
11.27
4.187
47.19
300.15
325.15
1 179.19
C3
Cold
18.09
0.296
5.35
136.05
315.15
958.19
Cold Utility (Oxygen)
Cold
24.42
0.919
18.50
70
126.05
1 258.07
34
Stream
H1
H2
H3
H4
C1
C2
C3
CU
FCp (kW/K)
6.21
6.21
6.21
6.21
14.84
47.19
5.35
24.42
873.15 683.15 493.15 399.57 349.10 325.15 320.15 303.15 300.15 136.05 126.05 70
35
Heat Integration PFD
36
5.0 COST CALCULATION
5.1 Bare Module Costs REACTOR
Parameter Volumetric Flowrate, Q Retention time (half-full), t Reactor Volume, V Vessel Inside Diameter, Di Vessel Length, L Design Type Material of Construction
Q=
SI 3812.78 ft3/hr 5 min 635.46 ft3 7.40 ft 14.8 ft Vertical Low- Alloy Steel
137 183.84 × .
Q = 3812.78 ft3/hr Retention time =5 min at half full :
×2 × ×
Volume, V = (3812.78 ft3/hr) (
× .
= 635.46 ft3 (
)
)
=18 m3
Assume L/ D = 2
V = (D/2)2L = ( D3)/2
D = (2V/ )1/3 = [2(635.46)/ ] 1/3 = 7.40 ft L= 2D = 14.8 ft
From figure A.16 (purchased reactor)
log =4.1052+0.5320log18−0.0005log18 = $59 185.74 Bare-Module cost From Table A.7 37
= = 4×$59 185.74 = $236 742.98 2016 =$236 742.98(587 397)=$350 045.67
38
DISTILLATION COLUMN
Sizing and costing of the main vessel: Parameters
S1
Domed head wall thickness, a
0.0137m
Tray spacing, b
0.6096m
Column diameter, c
0.4785m
Column wall thickness, d
0.0127m
Design type
Vertical
Material of Construction
Carbon steel
Material of insulation
Mineral wool,0.06m
Column type
Plate column
Plate type
Sieve
Domed head type
Tori spherical
Sizing and costing of the reflux drum: Parameters
S1
Domed head wall thickness,
0.0137m
Vessel length,L
21.36m
vessel diameter, D
1.792m
Column wall thickness, d
0.01427m
Design type
Vertical
Material of Construction
Carbon steel
Material of insulation
Mineral wool,0.06m
39
Sizing and costing of the condenser: Parameters
S1
Length of tube
2.697m
Area of transfer,Ac
58.75m2
Material of Construction
Carbon steel fixed head, shell tube
Type of HE
exchanger
Sizing and costing of the re-boiler: Parameters
S1
Length of tube
3.22m
Area of transfer,AR
188.53m2
Material of Construction
Carbon steel
Type of HE
u-tube
Total cost: Vessel or Equipment
Cost in $
main vessel
4587.49
the reflux drum
13,190.00
the condenser
100,006.00
the re-boiler
173,972.00
Total
291,755.49
40
Calculation: Vessel Bare module cost: The operating pressure: 320 barg. The pressure factor can be obtained from the following formula:
+1 +0. 0 0315 2850−0.6+1 ,,= 0.0063 320+10. 4 785 +0. 0 0315 2 850−0.6320+1 = 0.0063 =19.04 Since the equipment is vertical process tower made up of carbon steel,
=1 Purchased cost for the equipment can be calculated as follow in which A is the equipment size:
= + + = /4=0.4785 /4= 0.1798 =2.9949, =0.4465, =0.3961 =2.9949+0.44650.1798+0.39610.1798 =2.8821 =762.31 =2.25, =1.82 2001= 2001 + 2001=762.312.25+1.8211 2001= $3 102.611 For tray and packed tower,
,
Constant values for bare module factor of vertical process tower is
Annual index for the year 2001 and 2016 are 397 and 587,
2016=$3102.611(587 397) 41
,
2016= $4587.49×4.43= $ . The plant has 3 vessels of the same size hence
2016 = 3$20 322.57 = $60 967.71 Reflux Drum The pressure factor can be obtained from the following formula:
= −+ = = = 0 = 0 = 1 Since the equipment is vertical process tower made up of carbon steel,
=1 Purchased cost for the equipment can be calculated as follow in which A is the equipment size:
= + + = /4=2 /4 = 3.142 =2.9949, =0.4465, =0.3961 =2.9949+0.44653.142+0.39613.142 =3.32 =2065 For tray and packed tower,
,
Constant values for bare module factor of vertical process tower is
2001= 2001 + 2001=20652.25+1.8211 2001=$8921 Annual index for the year 2001 and 2016 are 397 and 587, 42
=2.25, =1.82
,
2016=$8921(587 397)
2016 = $13190×4.43 = $
Condenser The operating pressure is 17.5 barg. The pressure factor can be obtained from the following formula:
= −+ =0.03881, =−0.11272, =0.08183 =0.025 = 1 .06
Since the equipment is made up of carbon steel,
=1 Purchased cost for the equipment can be calculated as follow in which A is the equipment size:
= + + = 58.75 =4.3247, =−0.3030, =0.1634 =4.1884−0.303058.75+0.163458.75 =4.30 =19954 = 1.63 , =1.66 2001= 2001 + 2001=199541.63+1.661.061 2001=$67636 For fixed tube reboiler,
,
Constant values for bare module factor of u-tube reboiler is
Annual index for the year 2001 and 2016 are 397 and 587,
43
,
2016=$67636(587 397) 2016= $100006×4.43= $ Considering the plant has 2 condensers of the same capacity
2016 = 2$443 027 =$886 054 Reboiler The operating pressure is 17.5 barg. The pressure factor can be obtained from the following formula:
= −+ =0.03881, =−0.11272, =0.08183 =0.025 = 1 .06
Since the equipment is made up of carbon steel,
=1 Purchased cost for the equipment can be calculated as follow in which A is the equipment size:
= + + = 188.53 =4.1884, =−0.2503, =0.1974 =4.1884−0.2503188.53+0.1974188.53 =4.64 =43740 = 1.63 , =1.66 2001= 2001 + 2001=437401.63+1.661.061 For u-tube reboiler,
,
Constant values for bare module factor of u-tube reboiler is
44
,
2001=$117660.60 Annual index for the year 2001 and 2016 are 397 and 587,
2016=$117660.60(587 397) 2016= $173972×4.43= $ Considering the plant has 2 reboilers of the same capacity
2016 = 2$770 694 = $1 541 388
45
PUMP Parameters
Pump
Pump Type
Positive Displacement
Material of Construction
Stainless steel
Inlet Pressure, PI
145 psi
Outlet Pressure, PO
253.8 psi
Assumptions: Supplies power shaft at 7 bar = 162 kW Thus, ∆P = 7.5 bar (109 psi) = 174 kW
To find the bare module cost, CBM, CBM = Cpº FBM Since the unit does not operate at the near-ambient pressure, hence the condition for this equipment is non-base-case.
To find Cpº from Equation A.1, log Cpº = K1+ K2 log (A) + K3 [log(A)]2 -----------> Eq A.1 K1 = 3.4771, K2 = 0.1350, K3 = 0.1438, A = Shaft power = 174 kW log Cpº = 3.4771 + 0.1350 log (174 kW) + 0.1438 [ log (174 kW)]2 = 4.5
Cpº = $31 729
To find FBM from Figure A.18, FBM = B1 + B2FMFP
Value of B1 and B2 from Table A.4 B1=1.89 B2=1.35
46
From Table A.3, for compressor with steam turbine, the identification number to be used with Figure A.18 is 34
From Figure A.18, FM = 2.6 FBM = 1.89 + (1.35)(2.6)(1.0) = 5.4 Hence, CBM (2001) = ($31 279)(5.4) = $31 279 CBM (2016) = CBM (2001) (587/397) = $31 279 (587/397) = $46 249
47
COMPRESSOR The function of a compressor is to increase the pressure and corresponding saturation temperature (boiling point) of the refrigerant vapor to high enough level so the refrigerant can condense by rejecting its heat through the condenser. Parameters Compressor Type Drive Type Material of Construction Inlet Volumetric Flow Rate, QI Inlet Pressure, PI Outlet Pressure, PO Specific Heat Ratio, k
Compressor Centrifugal Steam turbine Stainless steel 3812.78 ft3/min 14.5 psi 145.04 psi 1.10
Assumptions: Supplies power shaft at 7 bar = 162 kW Thus, at ∆P = 9 bar (145.04 psi) = 208 kW
To find Cpº from Equation A.1, log Cpº = K1+ K2 log (A) + K3 [log(A)]2 -----------> Eq A.1 K1 = 2.2897, K2 = 1.3604, K3 = -0.1027, A = Shaft power = 208 kW log Cpº = 2.2897 + 1.3604 log (208 kW) - 0.1027 [ log (208 kW)]2 = 4.9 Cpº = $77 865
From Table A.6, for compressor with steam turbine, the identification number to be used with Figure A.19 is 2 FBM = 5.8 CBM (2001) = CpoFBM = 77 865 x 5.8 = $451 618
CBM (2016) = CBM (2001) (587/397) = $451 618 (587/397) = $667 757
48
HEAT EXCHANGER Heat exchanger type Design type Heat exchanger orientation Tube inlet direction Heat duty (kJ/s) Heat duty (Btu/hr)
2 shell and 4 tubes Fixed Head Horizontal Horizontal 1 179.2 4.02 x 10 6
= 1179.2 × 3412.ℎ14 =4.02×10 ℎ ∆ = ∆Tlog −∆T ∆T∆T ∆T ∆T = ℎ ℎ ℎ 7K−683.15−320.15K = 54.11 ∆ = 873.15−399.5log 473.36358 = ∆ = 0.45 1179 = 48. 2 0 .×54.11 Where
Costing
=4.3247, =−0.3030, =0.1634 log = +log+log = 4.3247−0.3030log48.20+0.1634 log48.20 = $ 18 949.22 =1.63, =1.66 =−0.00164, =−0.00627, =−0.0123 Operating pressure = 10 barg
log = +log+log =−0.00164−0.00627 log10−0.0123 log10 =0.95 49
=1 = [ +] = 18 949.221.63+1.660.951 = $ 60 770.15 Considering that we have 3 of the same heat exchanger,
= 360 770.25 = $182 310.75 Equipment
CBM (2001) ($)
Reactor
59 185.57
Vessel (Distillation Column)
9 307.83
Reflux Drum (Distillation Column)
8 921.00
Condenser (Distillation Column)
135 272.00
Reboiler (Distillation Column)
235 321.20
Pump
31 279.00
Compressor
451 618.00
Heat Exchanger
182 310.75
TOTAL
$1 094 986.69
According to the Chemical Engineering Plant Cost Index, 2016 has a cost index of 587 therefore
2016 =$1 094 986.69 (587 397) = $1 619 035.74
50
5.2 Fixed Capital Investment The fixed capital investment can be determined by the Grass Roots cost
= 1.18 Σ =1.18$1 619 035.74 =$1 910 462.17 = +0.50 Σ =$1 910 462.17+0.50$1 619 035.74 =$2 719 980.04 5.3 Cost Of Manufacturing
COST OF OPERATING LABOUR The technique used to estimate operating labour requirements is based on data obtained from five chemical companies and correlated by Alkhayat and Gerrrard (T.Richard, 2013). The operating labour requirement for chemical processing plants is calculated using equation below:
NOL =6.29+31.7P +0.23N. NOL
= the number of operator per shift
P
= the number of processing steps involving the handling of particulate solids
Nnp
= the number of non-particulate processing steps
N = Equipment Based on the process flow diagram,
P
0
Nnp
17
NOLno.of operators required per shift=6.29+31.7 0 +0.2317. =3.19 Based on eligible sources, the number of operators needed to provide this number of shifts is [(1095 shifts/y)/(245 shifts/operators/y)] or approximately 4.5 operators.
51
Operating Labour = (4.5)(3.19) = 14.36 = 14 operators
From the data obtained from the Bureau of Labour Statistics, they give the hourly miscellaneous plant and system operators at $8012.28 for a year.
Labour Cost (2015) = 14 x $8012.28 = $112 172
RAW MATERIAL COST Stream factor, SF = 0.92 Flowrate of propane = 52169.66kg/hr (stream 1) Flowrate of pentane = 1981.13kg/hr (stream 1) Flowrate of butane = 11886.76kg/hr (stream 1) Yearly cost of propane
365 52169.66 $ . 0.92=$239,654,062.60/. =24 Yearly consumption of propane
ℎ (24 )(365 )(52169. 6 6 )0. 9 2 ℎ = =420,445.72/ 1000 Yearly cost of pentane
365 1981.13 $ . 0.92 =$1,596,632.29/. =24 Yearly consumption of pentane
ℎ (24 )(365 )(1981. 1 3 )0. 9 2 ℎ = =15,966.32/ 1000 Yearly cost of butane
365 11886.76 $ . 0.92 =$12,928,347.82/. =24 Yearly consumption of butane
52
ℎ)0.92 ℎ )(365 (24 )(11886. 7 6 = =95,797.78/ 1000 Total yearly cost of crude oil
= $239,654,062.60+$1,596,632.29+$12,928,347.82= $254,179,042.70 Total yearly consumption of crude oil
=420,445.72+15,966.32+95,797.78=532,209.82 UTILITIES COST 1kW.h = 0.567 CNY (Chinese Yuan) Furnace Heat duty
= 69 162.56 kW
Yearly cost
= 69 162.56kW x 0.567/kWh x 0.92 x 24 hr x
= 316 042 910.3 CNY = $ 45 568 215.72 Reactor 923m3 = 186kW
From journal assumption
1m3 = 0.2015kW 18m3 = 3.627kW
For this reactor
Based on heuristics, power input to a homogeneous reaction stirred tank is 0.1-0.3kW/m3. So for reactor’s volume = 1 8 m3, need 3.627kW, efficiency = 84%
S = . kW = 4.318 kW .
Electric Power
=
Yearly cost
= Electric power x Cost x Stream factor
Yearly cost
=
4.3
kW
x
0.567
= 19 649.14 CNY 53
CNY
x
0.92
x
24
hr
x
= $2 833.45
Distillation Column Reboiler duty is 5.3 GJ.
Power consumption cost at reboiler, =
. ×5.3 GJ× ^ k × × × y ×0.92 kW. y y
= 280 309.05 CNY
Condenser duty is 5.1 GJ.
Power consumption cost at condenser, =
. ×5.1 GJ× ^k × × × y ×0.92 kW. y y
= 269 731.35 CNY
Total consumption power of tower
= 280 309.05 CNY + 269 731.35 CNY = 550 040.40 CNY x 4 tower (consider same) = 2 200 161.6 = $317 227.29
Pump Shaft work = 174kW Efficiency =
0.86 (interpolation from efficiency of reciprocating pump based on shaft work from heuristic)
Electric Power
=
S = kW = 200kW .
Yearly cost
= Electric power x Cost x Stream factor
Yearly cost
=
200
kW
x
0.567/kWh
x
0.92
= 913 913.28 CNY = $128 947.20
Compressor Shaft work = 208kW Efficiency =
0.93 (from figure 8.7 (Turton et, al, 2013) ) 54
x
24
hr
x
Electric Power
=
S = kW = 223.7 kW .
Yearly cost
= Electric power x Cost x Stream factor
Yearly cost
=
223.7
kW
x
0.567/kWh
x
0.92
x
24
hr
= 1 022 212.0 CNY
x
= $ 147 386.24
Heat Exchanger Heat duty
= 1 179.2 kW
Yearly cost
= 1 179.2kW x 0.567/kWh x 0.92 x 24 hr x
= 5 388 432.7CNY x 3 Heat Exchanger (considering same) = 16 165 298.10 CNY = $2 330 771.36 Utilities cost, CUT = $ 48 495 381.20
Type of Manufacturing Cost
Cost ($)
Operating
112 172
Raw Material
254 179 042.70
Utility
48 495 381.20
Total
$302 786 595.90
=0.18+2.73 +1.23 + + =0.18$2 719 980.04 +2.73$112 172 +1.23$48 495 381.20+$254 179 042.70 The cost of waste treatment is not available with reason that the by-product is not considered waste and can be sold to other manufacturing companies.
= $373 085 367.40 55
5.4 Revenue Flowrate of propylene = 49895.16kg/hr (stream 20) Yearly revenue from propylene sales
ℎ )(365 )(49895.16 )($ 0.958)( 1 )0.92 =$385,226,240.40/ =(24 ℎ 1 5.5 Cost Of Plot Of Land Our plant will be based in Pudong, Shanghai. The cost for 5 hectares of land is $500 000 per 5 hectares
5.6 Working Capital The working capital is 20% of the fixed capital investment therefore
= 0.2$2 719 980.04 = $543 996.01 6.0 PROFITABILITY ANALYSIS
COMd = $373
085 367.40
Land Cost = $500 000 Fixed Capital Investment = $2 719 980.04 Revenue = $385 226 240.40 Corporate Tax Rate in Shanghai = 25% Project life = 10 years Working Capital = $543 996.01
56
6.1 Non-discounted Profitability After Tax Cash Flow
End of Year
Investment
depreciation
FCI-d
Revenue
COMd
0
-
2,719,980.04
-
-
543,996.01
2,175,984.03
353,124,053.70
2
-500,000.00 3,263,976.05 0.00
870,393.61
1,849,586.43
385,226,240.40
373,083,367.40 14,833,486.27 373,083,367.40 9,324,753.15
3
0.00
522,236.17
2,197,743.87
385,226,240.40
373,083,367.40
9,237,713.79
4
0.00
313,341.70
2,406,638.34
385,226,240.40
373,083,367.40
9,185,490.18
5
0.00
313,341.70
2,406,638.34
385,226,240.40
373,083,367.40
9,185,490.18
6
0.00
156,670.85
2,563,309.19
385,226,240.40
373,083,367.40
9,146,322.46
7
0.00
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
8
0.00
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
9
0.00
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
10
1,043,996.01
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
1
End of
After Tax Cash
Investment
depreciation
FCI-d
Revenue
COMd
0
-500,000.00
-
2,719,980.04
-
-
1
-3,263,976.05
543,996.01
2,175,984.03
353,124,053.70
373,083,367.40 -14,833,486.27 -18,097,462.32
-18,597,462.32
2
0.00
870,393.61
1,849,586.43
385,226,240.40
373,083,367.40
9,324,753.15
9,324,753.15
-9,272,709.17
3
0.00
522,236.17
2,197,743.87
385,226,240.40
373,083,367.40
9,237,713.79
9,237,713.79
-34,995.38
4
0.00
313,341.70
2,406,638.34
385,226,240.40
373,083,367.40
9,185,490.18
9,185,490.18
9,150,494.80
5
0.00
313,341.70
2,406,638.34
385,226,240.40
373,083,367.40
9,185,490.18
9,185,490.18
18,335,984.97
6
0.00
156,670.85
2,563,309.19
385,226,240.40
373,083,367.40
9,146,322.46
9,146,322.46
27,482,307.44
7
0.00
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
9,107,154.75
36,589,462.19
8
0.00
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
9,107,154.75
45,696,616.94
9
0.00
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
9,107,154.75
54,803,771.69
10
1,043,996.01
0.00
0.00
385,226,240.40
373,083,367.40
9,107,154.75
10,151,150.76
64,954,922.45
Year
Flow -
Cash Flow
Cum cash flow
-500,000.00
-500,000.00
Payback Period
+ =$500 000+$543 996.01=$1 043 996.01 We observe from table that the working capital was recovered between year 3 and 4 therefore
976.05−−$34 995.38+3 = 3.5 = 3−4 $3$9 763 150 494.17−−$34 995.38 ℎ =$64 954 922.45 = ∑ ℎ ⁄∑ ℎ $83 552 384.77 ℎ = $500 000+$18 097 462.32 57
== 4.5 > 1 ℎℎ ℎℎ ℎ ℎ ℎ ℎ
= $64 954 922.45 = + ++ 10× 10 × $3 763 763 976. 976.0505 =1.73 Cumulative Cumulati ve Cash Flow Diagram 60,000,000.00 50,000,000.00 40,000,000.00 30,000,000.00 20,000,000.00 10,000,000.00 0.00 1
2
3
4
5
-10,000,000.00 -20,000,000.00 -30,000,000.00
58
6
7
8
9
10
6.2 Discounted Profitability Discount factor = 16% End of Year
Cash Flow (non -discounted)
Cash Flow (discounted)
Cumulative Discounted Cash Flow
0.00
-500,000.00
-500,000.00
-500,000.00
1.00
-18,097,462.32
-15,601,260.62
-16,101,260.62
2.00
9,324,753.15
6,929,810.61
-9,171,450.02
3.00
9,237,713.79
5,918,212.23
-3,253,237.79
4.00
9,185,490.18
5,073,064.45
1,819,826.66
5.00
9,185,490.18
4,373,331.43
6,193,158.09
6.00
9,146,322.46
3,754,037.21
9,947,195.30
7.00
9,107,154.75
3,222,380.28
13,169,575.59
8.00
9,107,154.75
2,777,914.04
15,947,489.62
9.00
9,107,154.75
2,394,753.48
18,342,243.11
10.00
10,151,150.76
2,301,099.43
20,643,342.54
$3 763 976. 976.05− 05 − $1 819 826.66 66 + 4 = 4.5 = 4 −$65$3 193 158.09− 193 09 − $1 819 826. 826.66 == $20 643 643 342. 342.54 744 603.16 = $36 $16 101 260.62 =2.28>1 ℎ = ℎℎ ℎ ℎ
59
End of
Cash Flow
NPV
NPV
NPV
NPV
Year
(Discounted)
(i = 0)
(i = 0.1)
(i = 0.2)
(i = 0.28)
0
-500,000.00
i=0
-500,000.00
-500000
-500,000.00
1
-15,601,260.62
-500,000.00
-14,182,964.20
-13001050.52
-12,188,484.86
2
6,929,810.61
-15,601,260.62
5,727,116.20
4812368.477
4,229,620.73
3
5,918,212.23
6,929,810.61
4,446,440.44
3424891.335
2,822,023.50
4
5,073,064.45
5,918,212.23
3,464,971.28
2446500.99
1,889,863.78
5
4,373,331.43
5,073,064.45
2,715,494.73
1757543.815
1,272,806.96
6
3,754,037.21
4,373,331.43
2,119,056.14
1257219.467
853,569.24
7
3,222,380.28
3,754,037.21
1,653,590.60
899307.1976
572,409.73
8
2,777,914.04
3,222,380.28
1,295,917.40
646054.0213
385,513.02
9
2,394,753.48
2,777,914.04
1,015,609.25
464119.2682
259,639.69
10
2,301,099.43
2,394,753.48
887,173.44
371640.4053
194,910.70
TOTAL
36,744,603.16
2,301,099.43
8,642,405.29
2,578,594.46
-208,127.51
Discount Rate (%)
NPV
0
2,301,099.43
10
8,642,405.29
20
2,578,594.46
28
-208,127.51
By interpolation we acquire the discount rate at NPV value of zero
28−20$0$0−− $2 578 28−20 578 594 594..46 + 20 = −$20 −$2088 127.51− 51 − 2 578 594. 594.46 =20+7.4=27.4% 60
Discounted Cumulative Cash Flow 10,000,000.00
5,000,000.00
0.00 1
2
3
4
5
-5,000,000.00
-10,000,000.00
-15,000,000.00
-20,000,000.00
61
6
7
8
9
10
11
7.0 SAFETY ANALYSIS Material Safety Data Sheet – Safety handling and storage
Propylene Propylene is extremely flammable. It may explode when heated. It may explode at high temperatures and pressure (955 atm; 327°C). Safety handling Ground and bond all lines and equipment associated with product system. All equipment should be non-sparking and explosion proof. Remove all sources of ignition. Ensure adequate ventilation. "NO SMOKING" signs should be posted in storage and use areas. Safety storage Outside or detached storage is preferred. Protect from physical damage. Cylinders should be stored upright with valve protection cap in place and firmly secured to prevent falling. Store in cool, dry, well-ventilated area of non-combustible construction away from heavily trafficked areas and emergency exits. Keep at temperatures below 52°C / 125°F. Propane Safety handling
Safety storage
Do not drop or roll any container. Keep all container valves closed when not in use. Keep protective caps (if applicable) on containers when not in use. Keep containers away from heat sources or temperatures exceeding 130°F. Store and transport containers with relief valves in vapour space. Containers, even those that have been emptied, can contain explosive vapours. Do not cut, drill, grind, weld or perform similar operations on or near containers. Pentane
Safety handling
Safety storage
Keep away from heat. Keep away from sources of ignition. Ground all equipment containing material. Do not ingest. Do not breathe gas/fumes/vapour/spray. Wear suitable protective clothing. In case of insufficient ventilation, wear suitable respiratory equipment. If ingested, seek medical advice immediately and show the container or the label. Avoid contact with skin and eyes. Keep away from incompatibles such as oxidizing agents. Store in a segregated and approved area. Keep container in a cool, well-ventilated area. Keep container tightly closed and sealed until ready for use. 62
Avoid all possible sources of ignition (spark or flame). Do not store above 25°C (77°F). Butane
Safety handling
Safety storage
Keep away from flame, sparks temperatures. Bond and ground containers. Use only in well ventilated areas. Keep away from flame, sparks temperatures. Store only in approved containers.
and
excessive
and
excessive
Butene It is flammable in air. Butene will rapidly or completely vaporize at normal pressure and temperature, or is readily dispersed in air and will burn readily. Safety handling Earth-ground and bond all lines and equipment associated with the butene system. Electrical equipment should be non-sparking or explosion proof. Compressed gas cylinders should not be refilled except by qualified producers of compressed gases. Safety storage Store only where temperature will not exceed 125°F (52°C). Post “No Smoking or Open Flames” signs in storage and use areas. There must be no sources of ignition Cylinders should be stored and used in dry, cool, well ventilated areas away from sources of heat or ignition. Do not store with oxidizers Hydrogen Safety handling
Keep away from heat, hot surfaces, sparks, open flames and other ignition sources. No smoking. Use only nonsparking tools. Use only explosion-proof equipment. Safety storage Keep away from heat, hot surfaces, sparks, open flames and other ignition sources. No smoking. Use only nonsparking tools. Use only explosion-proof equipment. This gas is heavier than air and in an enclosed space tends to accumulate near the floor, displacing air and pushing it upward. This creates an oxygen-deficient atmosphere near the floor. Ventilate space before entry. Verify sufficient oxygen concentration. When handling product under pressure, use piping and equipment adequately designed to withstand the pressures to be encountered. Hazard and Operability Study (HAZOP)
63
HAZOP for CSTR Reactor The reaction in this reaction is exothermic, and a cooling system is provided to remove the excess energy of reaction. If the cooling flow is interrupted, the reactor temperature increases, leading to an increase in the reaction rate and the heat generation rate. This situation may cause a runaway reaction with a subsequent increase in the vessel pressure possibly leading to a rupture of the vessel. The temperature within the reactor is measured and is used to control the cooling water flow rate by a control valve. The HAZOP analysis which can be adapted are in the plant for potential process modifications are: 1. Installation of a cooling water flow meter and low flow alarm to provide an immediate indication of cooling loss. 2. Installation of a high temperature alarm to alert the operator in the event of cooling function loss. 3. Installation of a high temperature shutdown system that would automatically shut down the process in the event of a high reactor temperature. The shutdown temperature would be higher than the alarm temperature to provide the operator with the opportunity to restore cooling before the reactor is shutdown. 4. Installation of a check valve in the cooling line to prevent reverse flow. A check valve could be installed both before and after the reactor to prevent the reactor contents from flowing upstream and to prevent the backflow in the event of a leak in the coils. 5. Periodic inspections and maintenance of the cooling coil to insure its integrity. 6. Evaluation of the cooling water source to consider any possible interruption and contamination of the supply.
Figure shows the exothermic reaction with cooling system in a reactor (copyright by RMIT University 2003)
64
Possible hazard that may happen during handling reactor: 1. Explosions from uncontrolled chemical reactions due to overpressure or due to release of energy from exothermal reactions (a result of mechanical damage or corrosion). 2. Exposure to carcinogenic, mutagenic and teratogenic substances present in the reactor, or released from the reactor during cleaning and maintenance work. 3. Exposure to high temperatures and heat-stress from chemical reactors. Preventative measure that can be done if any possible hazard occur: 1. Check periodically electrical equipment for safety before use and call a qualified electrician for testing and repair of faulty or suspect electrical equipment. 2. Install effective exhaust ventilation and air conditioning, especially in the operator's workstation, in order to prevent air contamination and heat stress; if necessary use chemicals to eliminate odours. 3. Use suitable protective clothing: appropriate headgear, goggles, earmuffs, gloves, boots, etc.
65
8.0 WASTE MANAGEMENT In PDH process, propane is selectivity dehydrogenated into propylene. Based on the process, it is produces lower paraffin hydrocarbon which is the propane. Waste from the plant can be treated by using Conventional Activated Sludge (AS) Plant; the proposed of the plant is composing of two parts; first is the plug flow aeration tank unit for the biological biodegradation of the waste compounds and second is the circular secondary settling tank. or Sequencing Batch Reactor (SBR) Plant; the proposed of this plant is composing of two parts; first is the sequencing batch reactor unit for the biological biodegradation and settling of the waste compounds, and second part is the equalization tank. The sequencing batch reactor (SBR) as in Figure 1 is a fill-and draw activated sludge system for wastewater treatment. In this system, wastewater is added to a single "batch" reactor, treated to remove undesirable components, and then discharged. Equalization, aeration, and clarification can all be achieved using a single batch reactor. To optimize the performance of the system, two or more batch reactors are used in a predetermined sequence of operations. SBR treatment process is as stated: 1. Collection cum Equalization Tank: The Influent is made to pass through the screens to remove bigger objects before entering the equalisation tank. The Screened Influent is collected and equalized in this tank. Coarse bubble diffusers are provided to keep the Influent in suspension and to give mobility to the liquid in the equalization tank. 2. Sequential Batch Reactor Tank: A specially designed tank is used for the processes of BOD removal, nitrification/denitrification, phosphorus removal and sludge separation which are achieved continuously and the process operates on the batch inflow and batch outflow principle. The phases of aeration, settling, decantation occur sequentially and are controlled by an automation system. The aeration is through micro fine bubble diffusion and here are the details: 1. Disinfection; to ensure that the treated water is free from bacteria, disinfectant is dosed. 2. Pressure sand filter; suspended particles in the treated water are removed by passing through the Sand Filter Vessel. The vessel system has backwash provision. 3. Activated carbon filter; the treated water is passed through this media in order to remove bad odour and colour if any. 66
Besides, SBR process advantages over conventional treatment process such as no settling tank and return sludge pumps, the state-of-the-art PLC control system provides a completely automatic operation of plant, the automatic operation results in minimal supervision requirements, contributing to lower operating and maintenance costs, and lastly, low F/M ratios result in significantly less waste sludge and due to its complete stabilization, sludge can be directly disposed into environment without further treatment.
Figure 1: Sequencing batch reactor (SBR) process flow diagram.
67
REFERENCES
1. By 2023, Approximately 30% of Global Propylene Production Will Be On-purpose, Which Is Derived from Metathesis and Propane Dehydration Processes, IHS Said. "Worldwide Demand for Propylene to Rise to 130 Million Tonnes by 2023, Says IHS." Chemicals Technology . N.p., n.d. Web. 26 Sept. 2016. 2. Zhitong Z, Y. L. (2015). Life cycle assessment of primary energy demand and greenhouse gas. Journal of Cleaner Production, 1-8. 3. Propane
dehydrogenation
http://www.ehow.com/how-does_5272515_propane-
dehydrogenation-process.html 4. Collins Discovery Encyclopedia (2005). (1st ed.). Glasgow: HarperCollins. 5. Product safety assessment finder (1995). Retrieved September 25, 2016, from http://www.dow.com/en-us/markets-and-solutions/product-safety/product-safetyassessment/product-safety-assessment-finder 6. Retrieved
September
25,
2016,
from
https://pubchem.ncbi.nlm.nih.gov/compound/Propene#section=Use-andManufacturing 7. Retrieved 13 December 2016, China Average Yearly Wages from http://www.tradingeconomics.com/china/wages 8. Retrieved 13 December 2016, Wages and Employment from http://www.clb.org.hk/content/wages-and-employment 9. Barrientos, Index Mundi (2015, Dec 12). Commodity Price: Propane. Retrieved from http://www.indexmundi.com/commodities/?commodity=propane 10. Brog,
Platts
(n.d.).
Platts
Global
Petrochemical
Index.
Retrieved
from
http://www.platts.com/news-feature/2014/petrochemicals/pgpi/propylene 11. US. Bureau of Labor Statistics (2016, Dec 14). Producer Price Index by Industry: Natural
Gas
Liquid
Extraction:
Butane.
Retrieved
from
https://fred.stlouisfed.org/series/PCU21111221111213 12. US. Energy Information Administration. Natural Gas Weekly Update [Washington, DC], (2016, Dec 6). Retrieved from http://www.eia.gov/naturalgas/weekly/ 13. Felder, R.M., Rousseau, R.W. (1999). Elementary Principles of Chemical Processes . Washington: Wiley Publication. 14. Turton, R., Bailie, R.C., Whiting, W.B., Shaeiwitz, J.A., Bhattacharyya, D. (2013). Analysis, Synthesis and Design of Chemical Processes. United States: Pearson
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68
