Production of Propylene from Methanol
Florida Institute of Technology
College of Engineering
Department of Chemical Engineering
Senior Design 2015/16
CHE 4182-Chemcial Engineering Plant Design II
Faculty Advisor: Dr. Jonathan Whitlow
Khalid Almansoori, Abdullah Kurdi and Nasser Almakhmari
AKA 1
April 27, 2016
Dr. Jonathan Whitlow
Florida Institute of Technology
Department of Chemical Engineering
150 W. University Blvd.
Melbourne, FL 32901
Dear Dr. Whitlow,
Enclosed you will find the requested report for the design of a production of
propylene from methanol plant. As requested, the report includes all parameters
and sizes of the new plant, an economic analysis that includes detailed cost estimates
and a sensitivity analysis of several parameters that might affect the rate of return on
investment, and key environmental and safety considerations.
If you have any questions, comments, or concerns about the report, please do not
hesitate to contact us at [email protected], [email protected], or
We thank you for giving us the opportunity to work with you in the design of this
plant, and we look forward to working with you in the future.
Sincerely,
Khalid Almansoori
_________________________
Abdullah Kurdi
_________________________
Nasser Almakhmari
_________________________
AKA 2
Table of Content
Letter of Transmittal 1
Table of Contents 2
Executive Summary 3
Introduction 4
Process Description 9
Process Flow Diagram, PFD 13
Stream Table 14
Utilities Table 17
Equipment Tables 17
Process Design and Simulation 24
Capital Cost 29
Manufacturing Cost 35
Profitability Analysis 41
Sensitivity Analysis 43
Process Control 44
Process Instrumentation Diagram, PID 46
Environmental and Safety Consideration 49
References 52
Appendix A: Equipment Design Methods, Calculations and Assumptions 55
Appendix B: Capital Cost Sample Calculations 77
Appendix C: Manufacturing Cost Sample Calculations 82
Appendix D: Profitability Analysis Sample Calculations 88
Appendix E: Literature Review 91
Appendix F: Project Timeline 105
AKA 3
Executive Summary
The following report explains, with in depth detail, the design and simulation
of a propylene production plant from methanol. The plant is to be sited in the
industrial city of Jubail, in the Kingdom of Saudi Arabia. The capacity of the plant is to
produce 480,000 metric tons of polymer grade propylene (99.6% purity) annually.
The plant also produces side-products of fuel gas (99.9% purity) at a rate of 55,400
metric tons per year, liquid petroleum gas (91.2% purity) at a rate of 285,000 metric
tons per year, and gasoline (99,5% purity) at a rate of 1,176,000 m3 per year. The
plant is assumed to operate 350 days a year, with 15 days for annual maintenance.
The novelty of the plant lies in the second reactor, MTP Reactor, catalyst. The catalyst
used in this process is Mordenite Zeolite (HMOR). This catalyst provides a higher
selectivity and conversion rate that produces twice the amount of propylene
compared to current used catalysts in industry. Market research was performed to
look into the feasibility of the products produced. Detailed design and simulation of
the plant is presented, as well as major environmental and safety considerations.
From the costing analysis, the estimated capital cost was found to be
$175,400,000 and the estimated cost of manufacturing was found to be $956,200,000.
From the profitability analysis, the estimated breakeven point is on the 5th year with
a return on investment rate of 37%. From the conduct sensitivity analysis, the main
affecting factors to the plant were the cost of raw materials, mainly the feed
methanol, and the selling price of propylene, which accounts for 60% of the
revenue.
AKA 4
Introduction
Propylene, also known as propene, is one of the most important raw materials
of the petrochemical industry; it is used in the production of a wide range of
chemical products. There are many ways of producing propylene; the main
industrial routes include Metathesis, Dehydration of Propane (PDH), Methanol-To-
Olefin (MTO) and Methanol-To-Propylene (MTP) (Jasper, 2015). (Refer to appendix E
for more information about the different routes of producing propylene).
During the past few years, the gap between the continuous consumption of the
restricted petroleum reserves and the increasing demand for propylene and its
derivatives has been increasing (Wen, 2016). The traditional petroleum-based
production of propylene (such as refinery fluid catalytic cracking (FCC) and steam
thermal cracking of naphtha) is hardly meeting the market demand (Wen, 2016). As
a result, it has become important to develop economical and energy efficient
processes that can fill the gap and replace the petroleum based production of
propylene (Wen, 2016).
The following design project is a production plant for producing on-purpose
polymer grade propylene using methanol as the feed, also known as Methanol-To-
Propylene (MTP) process. The plant is designed to have a production rate of 480,000
metric tons of propylene annually. The design and simulation of the plant was
conducted using Aspen Plus V8.8.
AKA 5
The plant is to be located in Jubail Industrial City, in the Kingdom of Saudi
Arabia. This city is the capital of petrochemical manufacturing in the kingdom,
where most of the petrochemical plants are located. Furthermore, the kingdom is a
large producer of methanol which comes from it having the 6th largest natural gas
reserves, and it being the 9th largest producer (BP Statistical Review of World
Energy, 2014). The kingdom has recently started looking into diversifying its
sources of income and wants to satisfy local petrochemical demand (Palagacheva,
2015). Another great advantage, from a business prospective, is that the kingdom
has low corporate tax rates.
The process novelty of this plant lies in the second reactor, MTP Reactor,
catalyst. The catalyst used in this process is mordenite zeolite (HMOR). Mordenite
has a silicon to aluminum ratio equal to 5. Comparing mordenite to the currently
used catalyst in industry, HZSM-5, it was found from the experimental results that
HMOR has twice the selectivity for producing propylene than HZSM-5 as well as a
significantly higher conversion rate (Moreno-Pirajan, 2013). This higher selectivity
reduces the number of reactors needed in the process from two to one, which will be
discussed further in this report. HMOR also helps in producing other useful and
valuable products such as fuel gas, liquid petroleum gas (LPG), and gasoline, all of
which with high purities (See Appendix E for more information about mordenite
zeolite and the experimental results).
AKA 6
Polypropylene, propylene oxide, and acrylonitrile are the most common
chemical derivatives from propylene (IHS, 2015). Polypropylene takes about 64% of
the total propylene consumption; propylene oxide accounts for about 7% of the total
propylene consumption; and acrylonitrile takes about 6% of the total propylene
consumption (IHS, 2015). The remaining 23% goes into the production of other
chemical derivatives such as acrylic acid, oxo alcohols, and cumene (IHS, 2015).
Polypropylene is used widely in the clothing industry and many consumer products
such as plastics, ropes, and carpets (CIEC, 2014). Propylene oxide goes into the
production of propylene glycol which is used as antifreeze for cars, deicing of
aircrafts, and goes into making cosmetics (HIS, 2015). Acrylonitrile goes into the
production of acrylic fibers, which are used in clothing and goes into the production
of paints and adhesives (IHS, 2015) (See appendix E for more information about the
products of propylene).
AKA 7
The demand for propylene increases annually and continues to be driven
primarily by developments in the polypropylene industry followed by the
propylene oxide and the acrylonitrile industries (IHS, 2015). Figure 1 shows the
increasing world demand and estimates that the demand for the year 2020 will reach
100 million tons (Galadima, 2015).
Figure 1: Historical and Expected Propylene Worldwide demand (Galadima, 2015)
AKA 8
The consumers of propylene are many, yet the largest consumers in the world
are China, followed by the United States, then Western Europe, together they
account for about 55% of the global propylene consumption (IHS, 2015). South Korea
and Saudi Arabia are also significant consumers of the global propylene market as
well. The following Figure 2, shows the global consumption of propylene as of 2014.
Figure 2: Global Consumption of Propylene by Country as of 2014 (IHS, 2015)
The project timeline is found in appendix F. The timeline highlights the major
tasks performed during the Spring semester of 2016 (1/11/2016 – 4/27/2016) to
complete this project. The period is represented by the semester working weeks,
from week 1, the first week of classes, to week 15, the last week of classes. Most of
the semester was spent on reviewing literature and gathering information as well as
simulating the plant.
AKA 9
Process Description
The Process Flow Diagram, PFD, of the plant is shown in Figure 3, with Stream,
Utility, and Equipment tables succeeding it in Tables 1, 2, and 3 respectively. The
stream tables show the specifications of each stream in the plant; that includes the
temperature, pressure, vapor fraction, mass flowrate, mole flowrate, and the
composition of the stream. The utility table shows the equipment unit number, the
type of utility used, and the amount of utility needed. Finally, the equipment tables
show each type of equipment and the design specifications for it.
The feed methanol (via pipeline from a neighboring plant) is pumped into the
process at 3.35 bar and 45 oC (via P-101) at a flowrate of 350,000 kg/hr; the feed then
goes through a heat exchanger (E-101) to be vaporized using low-pressure steam
(Hong, 2008). Low-pressure steam is converted to boiler feed water that can be re-
used for other purposes within the plant or to be sold. The vaporized feed (stream 3)
then goes through another heat exchanger (E-102) to be superheated to 266 oC
(Hong, 2008), prior to entering the Dimethyl-Ether Reactor, DME Reactor (R-101). In
the DME reactor (R-101), methanol is converted into dimethyl-ether and water via
the following equilibrium reaction (Farsi, 2010): 2𝐶𝐻3𝑂𝐻 ⟺ 𝐶𝐻3𝑂𝐶𝐻3 + 𝐻2𝑂
The DME reactor is a tabular shell and tube reactor (Farsi, 2010) with packed
aluminum oxide catalyst in the tube side, where the reaction occurs (Lurgi, 2003).
The reactor operates isothermally at 300 oC (Hong, 2008). The product of the DME
reactor (stream 5) goes through heat exchanger (E-103) to heat up the stream to 420
oC (Hong, 2008) prior to entering the Methanol-to-Propylene reactor, MTP reactor
(R-102).
AKA 10
In the MTP reactor (R-102), dimethyl-ether and the remaining unreacted
methanol are converted mainly into Ethylene, Propylene, Butene, Pentene, Hexene,
Heptene, Octene, and Water following the two general form of reactions
respectively (Meyers, 2005) (Hadi, 2014):
𝑛𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶𝑛𝐻2𝑛 + 𝑛𝐻2𝑂 𝑛 = 2, … ,8
𝑛𝐶𝐻3𝑂𝐻 → 𝐶𝑛𝐻2𝑛 + 𝑛𝐻2𝑂 𝑛 = 2, … ,8
The MTP reactor is a fixed bed reactor (Jasper, 2015) with mordenite zeolite,
HMOR, catalyst (Moreno-Pirajan, 2013). The reactor operates isothermally at 452
oC(Hong, 2008). Because the reactions taking place in the MTP reactor (R-102) are
exothermic, the stream going to the reactor (stream 6) is split into six streams to feed
the reactor at different levels (Lurgi, 2003); this method optimizes reaction control of
the MTP reactor (R-102) by controlling the flow of feed into the reactor, which then
limits the heat of reaction (Lurgi, 2003). The products of the MTP reactor (R-102) are
in the gaseous phase; the product stream (stream 7) is then compressed to 6.1 bar
via compressor (C-101) and has a temperature of 610 oC. This hot stream (stream 8)
is cooled down by passing through the shell side of the heat exchanger (E-103) to
heat up stream 5 to its desired temperature, and through the shell side of the heat
exchanger (E-102) to heat up stream 3 to its desired temperature. The hot stream
(stream 10) is finally cooled down to 38 oC via cooling water in heat exchanger (E-
104) before entering the flash separator (V-101). At this point, stream 11 is partially
condensed with water being the majority of the liquid composition.
AKA 11
In the flash separator (V-101), water leaves the separator from the bottom
(stream 13), and the gaseous hydrocarbon mixture from the top (stream 12). Stream
13 is to be sent to a neighboring wastewater treatment facility to be treated. Stream
12, containing the gaseous hydrocarbon mixture, is heated up to 104 oC via high-
pressure steam in heat exchanger (E-105) in preparation for compression. High-
pressure steam is converted to boiler feed water that can be re-used for other
purposes within the plant or to be sold. Stream 14 then enters a two stage
compressor series (C-102 & C-103) with equal pressure ratio and intermediate
coolers (E-106 & E-107) to compress the hydrocarbon mixture and filly condense it
in heat exchanger (E-107). Stream 18 exits heat exchanger (E-107) at 25 bar and 75
oC (Lurgi, 2003) prior to entering the first distillation column (T-101).
The first distillation column (T-101) has 44 stages with the feed entering on the
28th stage. In this column, the heavy hydrocarbons, mainly C5+, are separated from
the mixture and leave the column as gasoline in the bottoms product at a purity of
99.5%. The remaining hydrocarbon mixture leaves the column as the distillate
product and is sent to the second distillation column (T-102).
The second distillation column (T-102) has 33 stages with the feed entering on
the 11th stage. In this column, the C2- hydrocarbons, mainly ethylene in this
particular process, are separated from the hydrocarbon mixture as the distillate
product and leaves the column as fuel gas at a purity of 99.9%. The remaining
mixture leaves the column as the bottoms product and is sent to the third, and final,
distillation column (T-103).
AKA 12
The third distillation column has 48 stages with the feed entering on the 22th
stage. In this column, propylene is separated from the hydrocarbon mixture as the
distillate product and leaves the column at a purity of 99.6%. The remaining
hydrocarbon mixture, mainly butene in this particular process, leaves the process as
liquid petroleum gas, LPG, as the bottoms product at a purity of 91.2%. In general,
all product streams (19, 21, 23, and 24) are sent to their designated storage tanks to
be sold.
AKA 13
Figure 3: Process Flow Diagram, PFD
AKA 14
Table 1-1: Stream Summary
Stream Number 1 2 3 4
Temperature (oC) 45 45 97 266
Pressure (bar) 1.00 3.35 3.25 3.05
Vapor Mole Fraction 0 0 1 1
Flowrate (kg/hr) 350,000 350,000 350,000 350,000
Flowrate (kmol/hr) 10,923.11 10,923.11 10,923.11 10,923.11
Component Flowrates (kmol/hr)
Methanol 10,923.11 10,923.11 10,923.11 10,923.11
Water 0 0 0 0
Dimethyl-ether 0 0 0 0
Ethylene 0 0 0 0
Propylene 0 0 0 0
Butene 0 0 0 0
Pentene 0 0 0 0
Hexene 0 0 0 0
Heptene 0 0 0 0
Octene 0 0 0 0
Stream Number 5 6 7 8
Temperature (oC) 300 420 452 610
Pressure (bar) 2.70 2.50 1.60 6.10
Vapor Mole Fraction 1 1 1 1
Flowrate (kg/hr) 350,000 350,000 350,000 350,000
Flowrate (kmol/hr) 10,923.11 10,923.11 13,689.45 13,689.45
Component Flowrates (kmol/hr)
Methanol 1,485.24 1,485.24 2.97 2.97
Water 4,718.94 4,718.94 10,901.26 10,901.26
Dimethyl-ether 4,718.94 4,718.94 18.88 18.88
Ethylene 0 0 235.23 235.23
Propylene 0 0 1,353.26 1,353.26
Butene 0 0 573.46 573.46
Pentene 0 0 136.85 136.85
Hexene 0 0 114.04 114.04
Heptene 0 0 138.20 138.20
Octene 0 0 215.30 215.30
AKA 15
Table 1-2: Stream Summary
Stream Number 9 10 11 12
Temperature (oC) 506 383 38 38
Pressure (bar) 5.90 5.70 5.50 5.49
Vapor Mole Fraction 1 0.61 0.21 1
Flowrate (kg/hr) 350,000 350,000 350,000 154,069
Flowrate (kmol/hr) 13,689.45 13,689.45 13,689.45 2,815.82
Component Flowrates (kmol/hr)
Methanol 2.97 2.97 2.97 0.17
Water 10,901.26 10,901.26 10,901.26 30.46
Dimethyl-ether 18.88 18.88 18.88 18.87
Ethylene 235.23 235.23 235.23 235.22
Propylene 1,353.26 1,353.26 1,353.26 1,353.24
Butene 573.46 573.46 573.46 573.46
Pentene 136.85 136.85 136.85 136.85
Hexene 114.04 114.04 114.04 114.04
Heptene 138.20 138.20 138.20 138.20
Octene 215.30 215.30 215.30 215.30
Stream Number 13 14 15 16
Temperature (oC) 38 104 136 132
Pressure (bar) 5.49 5.29 11.65 11.45
Vapor Mole Fraction 0 1 1 1
Flowrate (kg/hr) 195,932 154,069 154,069 154,069
Flowrate (kmol/hr) 10,873.63 2,815.82 2,815.82 2,815.82
Component Flowrates (kmol/hr)
Methanol 2.80 0.17 0.17 0.17
Water 10,870.80 30.46 30.46 30.46
Dimethyl-ether 4.39E-03 18.87 18.87 18.87
Ethylene 0.01 235.22 235.22 235.22
Propylene 0.02 1,353.24 1,353.24 1,353.24
Butene 8.66E-04 573.46 573.46 573.46
Pentene 1.34E-05 136.85 136.85 136.85
Hexene 6.58E-07 114.04 114.04 114.04
Heptene 4.23E-08 138.20 138.20 138.20
Octene 2.05E-09 215.30 215.30 215.30
AKA 16
Table 1-3: Stream Summary
Stream Number 17 18 19 20
Temperature (oC) 168 75 236 58
Pressure (bar) 25.20 25.00 25.32 25.00
Vapor Mole Fraction 1 0 0 0
Flowrate (kg/hr) 154,069 154,069 56,745 97,324
Flowrate (kmol/hr) 2,815.82 2,815.82 603.00 2,212.82
Component Flowrates (kmol/hr)
Methanol 0.17 0.17 0.17 0.01
Water 30.46 30.46 0.98 29.48
Dimethyl-ether 18.87 18.87 4.12E-04 18.87
Ethylene 235.22 235.22 2.19E-08 235.22
Propylene 1,353.24 1,353.24 1.62E-03 1,353.24
Butene 573.46 573.46 1.84 571.62
Pentene 136.85 136.85 132.48 4.37
Hexene 114.04 114.04 114.04 4.30E-07
Heptene 138.20 138.20 138.20 4.68E-14
Octene 215.30 215.30 215.30 9.43E-21
Stream Number 21 22 23 24
Temperature (oC) -21 75 59.39 117.71
Pressure (bar) 25.00 25.24 25.00 25.34
Vapor Mole Fraction 0 0 0 0
Flowrate (kg/hr) 6,596 90,728 56,865 33,863
Flowrate (kmol/hr) 235.00 1,977.82 1,351.00 626.82
Component Flowrates (kmol/hr)
Methanol 8.28E-16 0.01 6.83E-13 0.01
Water 1.20E-08 29.48 7.86E-05 29.48
Dimethyl-ether 3.97E-05 18.87 4.85 14.03
Ethylene 234.77 0.46 0.46 4.10E-13
Propylene 0.23 1,353.00 1,345.60 7.40
Butene 4.06E-06 571.62 0.10 571.52
Pentene 2.17E-12 4.37 4.90E-10 4.37
Hexene 1.79E-23 4.30E-07 3.95E-23 4.30E-07
Heptene 0 0 0 0
Octene 0 0 0 0
AKA 17
Table 2: Utility Summary
Equipment Number E-101 E-104 E-105 E-106
Utility Type LPS CW HPS CW
Amount of Utility (ton/hr) 188.7 10,953 18.6 18.1
Equipment Number E-107 E-108 E-109 E-111
Utility Type CW CW HPS LPS
Amount of Utility (ton/hr) 1,046 792.7 13.1 0.793
Equipment Number E-112 E-113 R-101 R-102
Utility Type CW LPS CW CW
Amount of Utility (ton/hr) 761 0.45 3,645 13,498
Equipment Number C-101 C-102 C-103 P-101
Utility Type Electricity Electricity Electricity Electricity
Amount of Utility (MW) 39.3 2.4 2.4 0.04
Equipment Number P-102 P-103 P-104
Utility Type Electricity Electricity Electricity
Amount of Utility (MW) 0.04 0.02 0.05
Table 3-1: Equipment Summary
Pumps
P-101 A/B P-102 A/B
Centrifugal / electric drive Centrifugal / electric drive
Carbon steel Carbon steel
Power = 36.4 kW Power = 37.6 kW
82% efficient 81% efficient
P-103 A/B P-104 A/B
Centrifugal / electric drive Centrifugal / electric drive
Carbon steel Carbon steel
Power = 17.1 kW Power = 48.0 kW
74% efficient 82% efficient
AKA 18
Table 3-2: Equipment Summary (continued)
Towers
T-101 T-102
Carbon steel Carbon steel
44 CS sieve trays plus reboiler and
condenser
33 CS sieve trays plus reboiler and
condenser
60% efficient trays 70% efficient trays
Total condenser (E-108) Total condenser (E-110)
Feed on tray 28 Feed on tray 11
Reflux ratio = 0.81 Reflux ratio = 8.97
2 ft tray spacing 2 ft tray spacing
Column height = 32.2 m Column height = 24.1 m
Diameter = 4.8 m Diameter = 2.2 m
T-103
Carbon steel
48 CS sieve trays plus reboiler and condenser
70% efficient trays
Total condenser (E-112)
Feed on tray 22
Reflux ratio = 2.62
2 ft tray spacing
Column height = 35.1 m
Diameter = 3.3 m
AKA 19
Table 3-3: Equipment Summary (continued)
Reactors
R-101 R-102
Carbon steel, Shell & Tube, Packed tubes,
Aluminum Oxide Catalyst
Carbon steel, Process vessel, Fixed
bed, HMOR Zeolite Catalyst
V = 102 m3 V = 164 m3
Length = 8 m, Tube Diameter = 0.09 m L/D = 3.0
2000 Tubes 100% filled with active catalyst
Tubes are 100% filled with active catalyst Q = -281,784 MJ/hr
Q = - 75785 MJ/hr
Vessels
V-101 V-102
Carbon steel Carbon steel
Vertical Horizontal
L/D = 3.0 L/D = 3.0
V = 1956 m3 V = 63 m3
V-103 V-104
Carbon steel Carbon steel
Horizontal Horizontal
L/D = 3.0 L/D = 3.0
V = 27 m3 V = 79 m3
AKA 20
Table 3-4: Equipment Summary (continued)
Compressors
C-101 C-102
Carbon steel Carbon steel
Reciprocating Reciprocating
Power = 39.3 MW Power = 2.4 MW
85% adiabatic efficiency 75% adiabatic efficiency
C-103
Carbon steel
Reciprocating
Power = 2.4 MW
75% adiabatic efficiency
Heat Exchangers
E-101 E-102
A = 1,650 m2 A = 133 m2
Floating head, carbon steel Floating head, carbon steel
Process stream in tubes Process stream in tubes & shell
Q = 447,992 MJ/hr Q = 106,285 MJ/hr
AKA 21
Table 3-5: Equipment Summary (continued)
Heat Exchangers (continued)
E-103 E-104
A = 394 m2 A = 2,767 m2
Floating head, carbon steel Floating head, carbon steel
Process stream in tubes & shell Process stream in tubes
Q = 95,553 MJ/hr Q = 742,877 MJ/hr
E-105 E-106
A = 136 m2 A = 4.2 m2
Floating head, carbon steel Double pipe, carbon steel
Process stream in tubes Process stream in pipes
Q = 45,728 MJ/hr Q = 1,231 MJ/hr
E-107 E-108
A = 298 m2 A = 572 m2
Floating head, carbon steel Floating head, carbon steel
Process stream in pipes Process stream in pipes
Q = 70,956 MJ/hr Q = 52,871 MJ/hr
AKA 22
Table 3-6: Equipment Summary (continued)
Heat Exchangers (continued)
E-109 E-110
A = 1,329 m2 Area = 1,634 m2
Floating head, carbon steel Floating head, carbon steel
Process stream in tubes Process Stream in Tubes
Q = 25,176 MJ/hr Q = 19,238 MJ/hr
E-111 E-112
Area = 25 m2 Area = 600 m2
Floating Head, carbon steel Floating head, carbon steel
Process Stream in Tubes Process Stream in Tubes
Q = 2,024 MJ/hr Q = 51,621 MJ/hr
E-113
Area = 21 m2
Floating head, carbon steel
Process Stream in Pipes
Q = 1,056 MJ/hr
AKA 23
Table 3-7: Equipment Summary (continued)
Storage Tanks
V-105 V-106
Methanol Storage Fuel gas Storage
Volume: 32784 m3 Volume: 1374 m3
Capacity for 3 days Capacity for 3 days
Vertical tank on concrete pad Vertical tank on concrete pad
V-107 V-108
Propylene Storage Gasoline Storage
Volume: 9331 m3 Volume: 10074 m3
Capacity for 3 days Capacity for 3 days
Vertical tank on concrete pad Vertical tank on concrete pad
V-109
LPG Storage
Volume: 5555 m3
Capacity for 3 days
Vertical tank on concrete pad
AKA 24
Process Design and Simulation
The plant was designed and simulated using Aspen Plus simulator from Aspen
Technology Inc., version 8.8. The Aspen Plus design simulation of the plant can be
seen in Figure 4, at the end of this section. In general, Aspen Plus and the heuristics
from Turton were used to find the sizing parameters needed for designing and
costing purposes of the plant equipment.
Prior to starting the simulation of the plant, a property method had to be
chosen. Due to the majority of the components used in the plant being nonpolar, the
Peng-Robinson Equation of State was chosen as the property method for this plant
design (Aspen Physical Property System, 2011).
The following is a summary of how each type of equipment in the plant was
designed and the assumptions made. Detailed design methods, assumptions, and
calculations for each unit can be found in appendix A.
Pumps Design
Pumps were designed and simulated using Aspen Plus via inputting the
desired pressure discharge. The pressure discharge was specified based on the
desired pressure for a certain stream. Aspen Plus calculates the break power and
pump efficiency.
AKA 25
Heat Exchanger Design
Heat exchangers were designed and simulated in Aspen Plus via inputting
exchanger specifications, pressure drop, and heat transfer coefficient “U”.
Heuristics were used in determining the pressure drop and heat transfer coefficient
“U” values, which varied based on the application of the heat exchanger and the
fluids passing through the shell and tube sides. From the inputted information,
Aspen Plus calculates the area and heat duty of the heat exchanger, which was then
used for designing and costing purposes.
Compressor Design
Compressors in the plant were designed and simulated in Aspen Plus via
inputting compressor type, discharge pressure, and the efficiency of the
compressor. The discharge pressure was determined based on the process, while
the type of compressor and the efficiency were obtained from heuristics. From the
inputted information, Aspen Plus is able to calculate the break horsepower, which
was then used for designing and costing purposes.
Flash Separator & Reflux Drum Design
The flash separator and reflux drums in the plant were designed and
simulated in Aspen Plus. The specified variables were pressure and duty. The
holdup time, the length to diameter ratio, and the orientation of the vessel were all
based on heuristics; from these values with the use of the volumetric flowrate the
volume of the vessel was calculated, which was then used for design and costing
purposes.
AKA 26
Distillation Column Design
The distillation columns were designed and simulated in Aspen Plus. The
amount of distillate/bottom product was specified, and based on the desired purity
the reflux ration was varied. Furthermore, the condenser pressure was determined
based on the process, with the appropriate pressure drop from heuristics. The
columns were then optimized following the optimization process (Whitlow, 2016) to
obtain the optimum number of stages, reflux ratio, and feed stage. The type of trays
were specified in Aspen Plus, which led to determining the column diameter and
number of passes. Finally, tray spacing, tray efficiency, and column height were
determined using heuristics. All of the previous variables were obtained and used
for designing and costing of purposes.
Reactor Design
The first reactor, DME reactor (R-101), was designed and simulated in Aspen
Plus. The reactor was designed to operate isothermally (Hong, 2008), and was
modeled as a heat exchanger with the feed flowing into the tube side, where the
aluminum oxide catalyst is packed (Lurgi, 2003), and cooling water in the shell side
(Farsi, 2010). The number of tubes, diameter of tube, and reactor length were all
obtained from literature (Farsi, 2010); they were used to find the volume of the
reactor for designing and costing purposes.
AKA 27
The second reactor, MTP reactor (R-102), was simulated in Aspen Plus. This
reactor was a challenge to simulate due to the lack of reaction kinetics. Several
attempts were made to accurately simulate this reactor, but the complexity was high
and it was hard to simulate on Aspen Plus (refer to appendix D for further
information about the different attempts tackled in designing this reactor). From the
stoichiometric study of the reaction outputs, it was concluded that some of the side
products were produced in very small amounts, compared to the major products,
and neglecting them is a safe assumption. This assumption was then made to simplify
the Aspen Plus simulation. This reactor is a fixed bed with mordenite zeolite catalyst,
and operates isothermally. The reactor volume was determined using the weighted
hourly space velocity “WHSV” (Moreno-Pirajan, 2013); heuristics were used to find
the length to diameter ratio, all of which were obtained and used for designing and
costing of purposes.
Storage Tank Design
Storage tanks were not simulated on Aspen Plus, yet they were designed
based on feed and product flowrate, and heuristics. Heuristics were used to
determine the holdup time and orientation of the tanks. Using the holdup time and
the flowrate, the volume of the tank can be determined, which was then used for
designing and costing purposes.
AKA 28
Figure 4: Aspen Plus view of Plant Design Simulation of the Process
AKA 29
Capital Cost
The following section discusses the capital cost of the plant using the
methodology discussed in the Analysis, Synthesis and Design of Chemical
Processes, by Turton (Turton, 2012); it consists of several sources and helpful
parameters in estimating the cost of process equipment. The capital cost is the sum
of the costs of all process units. It is important to note that the data used in the
calculations are based on a survey of equipment manufacturers that were taken in
the year of 2001; the average Chemical Engineering Capital Cost Index (CEPCI) was
used to account for inflation. In 2001, CEPCI value was 397; the CEPCI for the 2016
was provided by Dr. Whitlow as 605, since the data was last updated in 2010 (Turton,
2012). The index was used to update the total capital cost to estimate the 2016 value
of the plant. There were some assumptions made in the design, all of which are
mentioned in the following Table 4.
Table 4-1: Assumptions Made in Calculating the Cost of certain Equipment
Unit Assumptions
Over Design Factor A safety over design factor of 10 %
Heat Exchangers
Some heat exchangers found to have
capacity not within the range. The
capacity was forced to be within the
range by dividing by lowest possible
number of exchangers. The final cost
was multiplied also by the number of
exchangers.
AKA 30
Table 4-2: Assumptions Made in Calculating the Cost of certain Equipment
Unit Assumptions
Reactors
DME Reactor is modeled as a shell
and tube heat exchangers – floating
head
MTP Reactor is modeled as process
vessels
A cooling jacket for the MTP reactor
was accounted for as 25% of the cost
of the MTP reactor. The cost of the
jacket was added to the final price of
the reactor.
Towers & Tanks
All of Towers and Tanks were
modeled as Process Vessels (vertical).
The assumptions were made to
find 𝐹𝑃,𝑉𝑒𝑠𝑠𝑒𝑙, FM, B1, and B2.
I. Capital Cost Methodology
Purchased Equipment Cost
The following equation was used for calculating the purchased cost of the
equipment, assuming ambient operating pressure (Turton, 2012):
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
Cpo: Purchased cost
A: Capacity or size parameter for the equipment
K1, K2, and K3: Coefficients that depends on the type of the equipment, given
constants (Turton, 2012)
Purchased Equipment Cost (for capacities out of the range)
Some of the equipment were found to have capacity not within the range. The
capacity was forced to be within the range by assuming there were more than one
piece of the equipment. The final cost was multiplied by the number of equipment.
AKA 31
Pressure Factors for Process Vessels
FP,Vessel
(P + 1) ∗ D2[850 − 0.6(P + 1)]
+ 0.00315
0.0063
The previous equation was used to determine the pressure factors for Vessels
and Towers. P is the pressers in barg, and D is the diameter in meter. There are
three Towers; all of the towers operating at the same pressure but different
diameter. The values of FP,Vessels were found to effects the cost due to the high
pressure factors value.
Pressure Factor for other Process Equipment
The pressure factor, FP, for other equipment such as Pumps, Heat Exchangers,
Compressors, and Reactors in the plant was found using the following equation:
log10Fp = C1 + C2 log10(P) + C3 [log10(P)]2
P: Design pressure in barg
C1, C2, and C3: Coefficients can be found by the type of the equipment (Turton, 2012)
Material Factors for Heat Exchangers, Process Vessels, and Pumps:
The values of the material factors, FM, for heat exchangers, process vessels
and pumps were obtained from Turton (Turton, 2012).
AKA 32
Bare Module Factor for Heat Exchangers, Process Vessels, Pumps and
Compressor:
CBM = CpoFBM = Cp
o(B1 + B2FMFp)
CPo: Purchased Cost
FBM: Bare module factor
B1 and B2: given constants (Turton, 2012)
FM: Material Factor used to find the cost for different materials of construction.
Fp: Pressure Factor
Bare Module and Material Factors for the Remaining Process Equipment
The values of the Bare Module and Material Factors, FBM and FM for the
remaining equipment were obtained from Turton (Turton, 2012).
Bare Module Cost for Sieve Trays
In the case of sieve trays, the bare Module cost was calculated differently; the
value of CBM is obtained using the following equation:
CBM = CpoNFBMFq
CPo: Purchased Cost
N: if N ≤ 20: log10Fq = 0.4771 + 0.08516 ∗ log10(N) − 0.3473 [log10(N)]2
CEPCI Costing Correction to 2016
Ca = Cb(Aa
Ab)
Where Ca is the cost of the equipment in 2016, Cb is the cost of the equipment
in 2001, Aa is the CEPCI in 2016 given by Dr. Whitlow to be 605, as a fixed
assumption, and Ab is the CEPCI in 2001 which is 397.
AKA 33
II. Capital Results
Table 5-1: Cost of Each Piece of Equipment
Equipment Unit # Cost (2016 $)
Pump
P-101 A/B $57,000
P-102 A/B $58,100
P-103 A/B $39,600
P-104 A/B $66,000
Vessel
V-101 $15,300,000
V-102 $480,000
V-103 $210,000
V-104 $600,000
Heat Exchangers
E-101 $755,000
E-102 $95,000
E-103 $193,000
E-104 $1,300,000
E-105 $96,000
E-106 $10,000
E-107 $156,000
E-108 $266,000
E-109 $610,000
E-110 $750,000
E-111 $60,000
E-112 $278,000
E-113 $60,000
Compressors
C-101 A/B $29,320,000
C-102 A/B $3,550,000
C-103 A/B $3,530,000
Towers
T-101 $9,470,000
T-102 $900,000
T-103 $3,480,000
Sieve Trays
T-101 $680,000
T-102 $83,000
T-103 $270,000
Reactors R-101 $63,000
R-102 $653,000
AKA 34
Table 5-2: Cost of Each Piece of Equipment
Equipment Unit # Cost (2016 $)
Tanks
V-105 (Methanol Storage) $ 22,800,000
V-106 (Fuel Gas Storage) $ 1,600,000
V-107 (Propylene Storage) $ 6,600,000
V-108 (Gasoline Storage) $ 7,000,000
V-109 (LPG Storage) $ 3,900,000
The following Figure 5 shows the total costs of the equipment used in the plant.
Figure 5: Total Cost of Equipment
The total capital cost was found to be $ 175,400,000. From Figure 5, it can be
noticed that the storage tanks govern the majority of the capital cost with 36.2%, the
second large cost is for the compressors, occupying 31.6 % of the total capital cost,
and the least expensive cost is the pumps, which occupy 0.2% of the total capital
cost.
$220,000 ,
0.2%
$16,600,000
, 14.4%
$4,600,000 ,
4.0%
$36,400,000
, 31.6%
$15,000,000
, 13.0%
$700,000 ,
0.6%
$41,800,000
, 36.2%
Fixed Capital Cost
Pumps Vessels Heat Exchangers
Compressors Columns\Towers Reactors
AKA 35
Manufacturing Cost
After calculating the capital investment needed to build the plant, the
operational investment is next to be determined. There are three types of
manufacturing costs to take into account: Direct Manufacturing Costs, Fixed
Manufacturing Costs, and General Expenses.
Direct Costs are dependent on production rate, and it includes raw materials,
utilities, labor, waste treatment, supplies, maintenance, lab charges, and patents &
royalties. Fixed costs are independent of production rate, and it includes taxes &
insurance and plant overhead. Finally, general expenses are loosely tied to the
production rate, and it includes sales and marketing, research & development and
administrative costs.
The following sections shows the manufacturing cost for the production of
Propylene from Methanol plant, using the methodology from Turton (Turton, 2012).
There are values that need to be found first to calculate the cost of manufacturing
(COM); the first value is the fixed capital investment (FCI), Cost of Operating Labor
(COL), Cost of Utility (CUT), Cost of Waste Treatment (CWT), and Cost of Raw Materials
(CRM). The Following equation was used to calculate COM without Depreciation:
𝐶𝑂𝑀 = (0.18 × 𝐹𝐶𝐼) + (2.73 × 𝐶𝑂𝐿) + (1.23 × (𝐶𝑈𝑇 + 𝐶𝑊𝑇 + 𝐶𝑅𝑀))
AKA 36
I. Manufacturing Methodology
Operating Labor
Labor wages and total number of operators are needed to calculate the COL.
Annual labor wages were found to be $52,500 in 2014 (51-9011 Chemical Equipment
Operators and Tenders, 2015). The Annual labor wages in 2016 is estimated to be
$53,550 via assuming a 2 % increase from 2014. The total number of operators is
found by using the following equation:
𝑁𝑂𝐿 × # 𝑜𝑓 𝑜𝑝𝑒𝑟𝑎𝑡𝑜𝑟 ℎ𝑖𝑟𝑒𝑑 𝑓𝑜𝑟 𝑒𝑎𝑐ℎ 𝑜𝑝𝑒𝑟𝑎𝑡𝑜𝑟.
𝑁𝑂𝐿 = √6.29 + 31.7𝑃2 + 0.23𝑁𝑛𝑝
This equation represents the number of operators per shift. P is the number of
steps involving particulate solids handling. Nnp is the number of steps not involving
particulate solids handling. P will be zero because there are no solids that need
handling such as no transportation or Particulate removal. Nnp is the number of none
particulate process includes reactors, towers, compressor, and heat exchanger;
Pumps and vessels are not included (Turton, 2012). The number of operator hired for
each operators is found to be 4.3 and this number should be rounded to 4.5; using
the following equation, the cost of operating labor is found to be $963,900
COL = The Annual labor wages × Total number of operators
Cost of waste treatment
The only waste of the process is water. Using the mass flow rate and the
estimated price of treatment (0.041 $/MT )(Turton), the cost of waste treatment is
found to be $67,479 per year.
AKA 37
Utility Cost:
In this section, the expenses associated with electricity, cooling water,
process steam and many other utilities are accounted for. It is important to note that
the costs of utilities are dependent on both inflation and energy cost. The main
utilities needed in the plant are electricity, cooling water, high-pressure steam, low-
pressure steam, and refrigerant; these utilities are used in the plant in heat
exchangers, reactors, compressors, and pumps. Table 6 shows the total amount of
each utility needed in the plant annually, and the price and annual cost of each
utility.
Table 6: Price, Total Amount, and Cost Annually Needed for Utilities
Cost (2016 $) Total Amount Needed Cost ($/yr) in 2016
Electricity ($/kW-hr) 0.0718 367,045,401 $26,400,000
Cooling water ($/kg) 0.0000175 258,002,773,140 $4,500,000
High Pressure Steam
($/kg) 0.01459 266,292,516 $3,900,000
Low Pressure Steam
($/kg) 0.01348 1,595,698,056 $21,500,000
Refrigerant ($/GJ) 11.2671 16,175 $182,000
The total utility cost, from Table 4, is $56,482,000. It can be noticed from Table
6 that the cost of electricity occupies a large part of the total with 46.74 % of the cost.
The next largest utility is low-pressure steam with a total cost of $21,500,000, which is
38% of the total utility cost.
For the cost of electricity, the cost was linearly extrapolated using the data
found in the U.S. Energy Information Administration (Electric Power Monthly, 2015) ;
to find the estimated price in 2016. All the calculation that are involved in the
manufacturing cost of the plant are presented in appendix C. Due to the dependence
AKA 38
of cooling water on electricity, the change in cooling water cost and refrigerant cost
were based on the total annual increase in the cost of electricity between 2006 and
2015. Due to the dependence of high and low pressure steam on natural gas, high
pressure steam and low pressure steam change in cost were based on the total
annual increase between 2009 and 2015 for natural gas (Annual Energy Outlook
2015, 2015) (See appendix C for more details). The data presented in appendix C
shows the linear extrapolation that was used for the electricity cost.
Cost of raw Material
The raw materials in the plant are methanol, and the catalysts. The mass
flowrate of methanol was 350000 kg/hr. Moreover, the price of methanol was found
to be 235 $/ton (Argaam Petrochemical Index Loses 3.7 Pts as Polymers, 2015).
Therefore, the mass flowrate was converted in ton/year in order to find the final cost
in unit of $/year. The amount need for Aluminum Oxide catalyst was 402.9 ton/yr for
the first reactor (DME, R-101), the life time for the catalyst is ten years. The amount
needed for Mordenite Zeolite catalyst was 344.4 ton/yr for the second reactor (MTP,
R-102). The price of Aluminum Oxide Catalyst was found to be 1000 $/ton and the
Mordenite Zeolite Catalyst was found to be 120 $/ton (Aluminum Oxide price, 2016).
AKA 39
II. Manufacturing Results
The sample calculations for the cost of Manufacturing can be found in
appendix C. The following Table 7 shows a summary of the costs included in the
manufacturing of the plant.
Table 7: Summary of the Costs included in the Manufacturing Cost
Direct Manufacturing Costs $791,600,000
Direct Supervisory and Clerical Labor $174,000
Maintenance and Repairs $10,600,000
Fixed Manufacturing Cost $12,700,000
Local taxes and Insurance $5,700,000
Plant Overhead costs $7,000,000
Raw Materials $691,400,000
Utilities $58,000,000
Operating Labor $964,000
Waste treatment $103,000
Lab Charges $145,000
Patents and Royalties $28,700,000
Fixed Capital Investment $175,420,000
Cost of Manufacturing $956,200,000
General Expenses $154,730,000
Administration Costs $1,750,000
Distribution and Selling Costs $105,200,000
Research and Development $47,810,000
Form Table 7, the total manufacturing cost of the plant is $956,200,000 while
the fix capital investment is $175,420,000. The direct manufacturing cost is
$791,600,000; the fixed manufacturing cost is $12,700,000 and the general expenses
have a total of $154,730,000.
AKA 40
Table 8: Direct Costs Distribution
Maintenance and Repairs 10,500,000 1.32%
Raw Materials 691,300,000 87.74%
Utilities 56,400,000 7.16%
Operating Labor 963,000 0.12%
Waste treatment 103,000 0.01%
Lab Charges 145,000 0.02%
Patents and Royalties 28,700,000 3.63%
Table 8 above, shows the distribution of the direct costs between its elements;
it can be noticed that Raw material take a large part of the pie chart with an 87.74%
of the total direct cost. The second largest cost is for the utilities of the plant, which
occupies 7.16% of the pie chart. The third largest segment in the pie chart is Patents
and Royalties occupying 3.63% of the direct costs. Maintenance and Repairs occupy
1.32%, while operating labor, waste treatment, and lab charges occupy 0.12%,
0.01% and 0.02% respectively.
AKA 41
Profitability Analysis
I. Profitability Methodology:
Profitability of the plant was determined through several steps. The product
annual flow rate and the cost of the product were calculated in order to find the
revenue. A spreadsheet was used to calculate the profitability analysis and some
assumptions were made in the profitability calculation; sample calculations are
shown in appendix D (Refer to the Excel Spreadsheet for more details). The land cost
was assumed to be equal to 5 million dollars. In addition, the annual interest rate was
assumed to be 6%(Turton, 2012). The revenue was assumed to increases by 6%
annually based on products price trends (Dukandar, 2016), and the operation cost
by 2% annually. The tax rate was assumed to be 20% (Saudi Arabia: Tax System,
2016); while the working capital was assumed to be 15% of the fixed capital
investment. The construction period was assumed to be two years, with an expected
plant lifetime of ten years (Turton, 2012).
The most up to date prices for all products are shown in Table 9, below.
Propylene price is 1250 dollar per ton (Dukandar, 2016) .From Saudi Aramco,
liquefied petroleum gas (LPG) prices for September 2016 are 20-45$ /ton (Argaam
Petrochemical Index Loses 3.7 Pts as Polymers, 2015). The price of the ethylene is
1177 $/ton (Dukandar, 2016).The price of gasoline depends on the location, in Saudi
Arabia the gasoline is considered one of cheapest country comparing to other
countries, the price of gasoline is 0.3 $/L (Petrol Prices Across the World, 2016).
Boiler feed water is produced in this plant; the steam can be sold to other
neighboring plant for 2.54 $/MT (Turton, 2012).
AKA 42
Table 9: Annual Revenue Per Product
Products Amount Annually Price Total
Propylene (MT/yr) 477,668 1250 $/MT $597,100,000
Ethylene (MT/yr) 55,405 1000 $/MT $55,400,000
LPG (MT/yr) 284,446 45 $/MT $12,800,000
Boiler Feed Water (MT/yr) 1,741,230 2.45$/MT $4,300,000
Gasoline (L/yr) 1,175,309,856 0.3 $/L $352,600,000
II. Profitability Results:
Figure 6: Cumulative Future vs. Time
Based on Figure 6 above, it can be concluded that the breakeven point is
going to be at fifth year of operation. The revenue in the fifth year will be $
26,700,000. The discounted cash flow rate of return (DCFROR) was found to be at a
37.34% annually interest rate.
-$200
$0
$200
$400
$600
$800
$1,000
$1,200
$1,400
0 1 2 3 4 5 6 7 8 9 10 11 12
Ca
sh F
low
, M
illi
on
US
Do
lla
rs
Time (year)
Cash Flow
AKA 43
Sensitivity Analysis
Table 10: Sensitivity Analysis parameters, Breakeven point and Rate of Return
Parameter % Change Break Even Year Rate of Return
Base Case 0 5 37.3%
Buying Price of Methanol 10 7 26.6%
Buying Price of Methanol -10 4 46.7%
Selling Price of Propylene 15 4 47.7%
Selling Price of Propylene -15 8 24.5%
Selling Price of Ethylene 30 5 39.4%
Selling Price of Ethylene -30 6 35.2%
Selling Price of LPG 25 5 37.8%
Selling Price of LPG -25 5 36.9%
Selling Price of Gasoline 15 5 43.7%
Selling Price of Gasoline -15 7 30.1%
Fixed Capital Investment 40 6 30.3%
Fixed Capital Investment -40 4 48.7%
Aluminum Oxide Catalyst Cost 60 5 37.31%
Aluminum Oxide Catalyst Cost -60 5 37.38%
Mordenite Zeolite Catalyst Cost 50 5 37.34%
Mordenite Zeolite Catalyst Cost -50 5 37.34%
Utility cost 15 6 36.1%
Utility cost -15 5 38.6%
Interest Rate 50 6 37.3%
Interest Rate -50 5 37.3%
Tax Rate 20 5 36.5%
Tax Rate -20 5 38.1%
The table above shows some significant parameters that were varied to study
the return of investment and breakeven year changes. Based on the sensitivity
analysis, it was noticed that cost of raw materials (methanol) and selling price of
Propylene are very sensitive. When the buying price of methanol increases by 10%,
the breakeven year movers to the 7th year. Likewise, when the buying price of
methanol decreases by 10%, the breakeven year moves to the 4th year.
AKA 44
Based on the sensitivity analysis, when the selling price of propylene
increases by 15%, the breakeven year becomes on the 4th year likewise when
decreases by 15%, breakeven year moves to the 8th year.
Process Control
The process instrumentation diagram, PID, is shown in Figure 7. The diagram
displays the control scheme that is proposed for the plant. It should be noted that
only the important controls are shown in order to avoid complicity and to emphasis
important controlling areas.
As illustrated on heat exchanger E-101, it utilizes a feedback control loop that
determines the outlet feed temperature and maintains the cooling water stream; the
cooling water valve is to fail open. Moreover, a feedback control loop is used in all
the condensers that are using cooling water and the coolers that are used in the
multistage compressors as well, yet that is not shown in the P&ID to avoid complicity.
The feedback control loop, as shown on the condenser in Figure 8, is to monitor the
process stream’s outlet temperature and manipulate the cooling water valve to
adjust the feed temperature.
As illustrated on heat exchanger E-105, when the stream is heated up, a feed
backward loop is used to regulate the outlet temperature of the feed; the valve here
is to fail close. In addition, advance controlling is applied on the steam valve via
cascade to manipulate the steam flowrate to control the temperature. Likewise, this
control loop is to be used in all the distillation column reboilers.
On distillation columns and reflux drums, the level of the liquid is controlled
via feed backward loop that manipulates a fail open valve on the liquid outlet. On
AKA 45
distillation column condenser, the reflux ratio is manipulated by a ratio controller.
The ratio controller manipulates a fail open valve that controls the flow ratio of the
reflux and the product that goes to the storage tank.
In the flash separator V-101, the pressure is controlled using a fail open valve
adjusted on the gas phase stream outlet, and the liquid level is controlled by fail
closed valve adjusted on the liquid stream output. For advance controlling, the stage
temperatures in all the distillation columns are cascaded with the liquid output to
insure achieving purities. Figure 8 presents the detailed controlling on one of the
distillation columns, which is also applies to the rest. Finally, Figure 9 shows the
water jacket design on the MTP reactor with necessary controlling (Refer to
appendix E for the discussion on relief systems).
AKA 46
Figure 7: Process Instrumentation Diagram, PID
AKA 47
Figure 8: Distillation Column Process Instrumentation Diagram
AKA 48
Figure 9: MTP Reactor (R-102) with Water Jacket Process Instrumentation Diagram
AKA 49
Environmental and Safety Considerations
Based on the analysis that conducted and presented on Tables 2E - 5E in
appendix E, it was noticed that there are several considerations with regards to the
environment and safety impacts of our plant. The environmental and safety impacts
of the chemical plant could affect the employers as well as the people in the
surrounding area. These impacts can be in different forms such as gases vented to
the atmosphere, the catalyst waste, and the noise that is produced from the plant.
The main product and side products are flammable hydrocarbons. Based on
the design of the process, the distillation columns are operating under high pressure
to achieve the needed purities and reduce operating cost. Relief Systems should be
designed to prevent inadvertent release of gases from the distillation columns. In
addition, the plant operates under high pressure therefore, it might cause pipe
rupture which might lead to gas leaks. Pipe rupture, due to high pressure, can be
prevented via using proper thickness and insulate on the pipes, also to prevent
heating the surface in the summer time. The lifetime of the mordenite zeolite catalyst
is estimated to be one year (Delft, 2009). The waste catalyst is going to be disposed
because it has no significant effects on the environment and it is not hazardous waste
(Wurzel, 2006). It was assumed that the catalyst will be likely disposed in a landfill.
All unites in the chemical plant produce relatively high noise which can annoy
people around the plant therefore, perimeter barrier is an ideal option to reduce the
noise to minimum. The environmental precautions in the case of accidental releases
are to, carefully contain and stop the source of the spill, if safe to do so. Protect
bodies of water by diking, absorbents, or absorbent boom, if possible. Do not flush
AKA 50
down sewer or drainage systems, unless system is designed and permitted to
handle such material.
From Table (E2), in appendix E, there are several safety considerations in the
plant due to the presence of high temperature, high pressure, toxicity and
flammability of the chemicals in the plant. As safety procedures, a relief system
should be designed in the vessels to prevent rapture due to excessive pressure. Fail
open valves were placed in each vessel to insure that no overheat occurs. In
addition, water jacket was designed in the MTP reactor to control the exothermic
reactor temperature. The main product and the side products of the plant are
extremely flammable and toxic. Therefore, the storage areas should be handled with
extreme care. All the employers of the plant should follow the regulations and use all
the necessary personal protective equipment
Emergency planning is primarily for the protection of plant personnel and
people in nearby areas and the environment that could be affected by plant
problems. It should be considered early in the design and should be coordinated
with related agencies. Emergency planning includes tornado and storm shelters,
flood protection, earthquakes, proximity to public areas, and safe exit routes.
Designing relief-venting systems is important to ensure that flammable or
toxic gases are vented to a safe mannar. This will normally mean venting at a
sufficient height to ensure that the gases are dispersed without creating a hazard.
For highly toxic materials it may be necessary to provide a scrubber to absorb or
change the material; for instance, the provision of caustic scrubbers for heavy
AKA 51
hydrocarbons. If flammable materials have to be vented at frequent intervals, such
an example is in some refinery plants, flare stacks are used (VEP Fire System, n.d.).
A deluge system is a water mist system using open spray heads attached to a
piping system that is connected to a water supply through a valve that is opened by
means of a detection system installed in the same area as the spray heads. When the
valve opens, water flows into the piping system and discharges through all spray
heads attached to the system. Deluge systems are typically used for the protection of
machinery with flammable liquid fire hazard (VEP Fire System, n.d.)
AKA 52
References
Agiurre L, L. B. (n.d.). Unconformities as mineralogical breaks in the burial metamorphism
of the Andes. vommission on nutural zeolites, 3 and 4.
Aspen Physical Property System. (2011, March). Retrieved from Aspen Tch.
Aluminum Oxide price. (2016, 3 15). Retrieved from Alibab.com Global trade :
www.alibaba.com/showroom/aluminium-oxide-price.html
Annual Energy Outlook 2015. (2015, April 14). Retrieved from U.S. Energy
Information Administration:
http://www.eia.gov/forecasts/aeo/section_prices.cfm#natgas
Argaam Petrochemical Index Loses 3.7 Pts as Polymers. (2015, December 2).
Retrieved from ArgaamPlus.
Beychok, M. (2012, July 16). The Encylcopedia of Earth. Retrieved from Flare
stacks: http://www.eoearth.org/view/article/173693/
BP Statistical Review of World Energy [Review]. (2014).
CIEC Promoting Science at the University of York, York, UK, Polyproplyene. (2014,
January).
Chemical Equipment Operators and Tenders. (2015, December 2). Retrieved from
U.S. Bureau of Labor Statistics:
http://www.bls.gov/oes/current/oes519011.htm
Dukandar, K. N. (2016). Alternative On-purpose Production Methods for Propylene.
CB&I.
Electric Power Monthly. (2015, Agust). Retrieved from U.S. Energy Information
Administration: http://www.eia.gov/forecasts/aeo/section_prices.cfm#natgas
Engineering Toolbox. Ethylene Glycol Heat-Transfer Fluid. (n.d.). from
http://www.engineeringtoolbox.com/ethylene-glycol-d_146.html
AKA 53
Farsi, M., Jahanmiri, A., & Eslamloueyan, R. (2010). Modeling and Optimization of
MeOH to DME in Isothermal Fixed-bed Reactor. International Journal of
Chemical Reactor Engineering, 8(1).
Galadima, A., & Muraza, O. (2015). Recent Developments on Silicoaluminates and
Silicoaluminophosphates in the Methanol-to-Propylene Reaction: A Mini
Review. Industrial & Engineering Chemistry Research Ind. Eng. Chem. Res.,
54(18), 4891-4905.
Hadi, N., Niaei, A., Nabavi, R., Farzi, A., & Navaei Shirazi, M. (2014). Development of
a New Kinetic Model for Methanol to Propylene Process on Mn/H-ZSM-5
Catalyst. 28(1), 53-63.
Hong, S. (2008). Retrofit Design of Methanol-to-Propylene Process for the Changes in
Feedstock and Catalyst. Korea Advanced Institute of Science and Technology.
IHS, Chemical Economics Handbook, (2015, February).
Jasper, S., & El-Halwagi, M. (2015). A Techno-Economic Comparison between Two
Methanol-to-Propylene Processes. Processes, 3(3), 684-698.
Lurgi MTP Plant Process Description. (2003). CHE397Hotel
Liebner, W. (2005). Lurgi MTP Technology. Access Engineering.
Louvar, D. A. (n.d.). Chemical Process Safety. New Jersey: Prentice Hall International
Series in the Pjysical and Chemical Engineering Sciences.
Meyers, R. (2005) Handbook of Petrochemicals Production Processes. LURGI MTP®
TECHNOLOGY, AccessEngineering
Moreno-Pirajan, J. C., & Giraldo, L. (2013). Catalytic Conversion Process of
Methanol-To-Propylene (MTP) With Zeolites. Rasayan J. Chem., 6(3), 172-174.
AKA 54
Mindat.org. Mordenite, from http://www.mindat.org/min-2779.html
Guide for Chemical Spill Response Planning in Industerial Areas. Waste Management
American Chemical Society, Washington, DC 1995: American Chemical
Society's CEI/CCS.
Petrol Prices Across the World. (2016, March 19). Retrieved from Kshitij Consultancy
Services: http://www.kshitij.com/research/petrol.shtml
Palagacheva, Y. (2015, February 13). Market outlook: Saudi Arabia's big
petrochemical adventure.
Pilling, M. (2012). Be Smatrt About Column Design. Reaction and Separation.
Saudi Arabia: Tax System. (2016, April). Retrieved from Santander TradePortal:
https://en.santandertrade.com/establish-overseas/saudi-arabia/tax-system
Turton, Baillie, Whiting, Shaeiwitz & Bhattacharyya. (2012). Analysis, Synthesis and
Design of Chemical Processes.
Ullmann's Encyclopedia of Industrial Chemistry, Propylene Glycol &. (2014)
VFP Fire System. (n.d.). Retrieved from Delge Fire Sprinkler
System: http://www.vfpfire.com/systems-deluge.php
Wen, M., Ding, J., Wang, C., Li, Y., Zhao, G., Liu, Y., & Lu, Y. (2016). High-
performance SS-fiber@HZSM-5 core–shell catalyst for methanol-to-propylene:
A kinetic and modeling study. Microporous and Mesoporous Materials, 221,
187-196.
Whitlow, Jonathan. (2016). Using Aspen Plus for Column Sizing. Canvas.FIT.edu/Files
AKA 55
Appendix A: Equipment Design Methods, Calculations and Assumptions
The following appendix presents the detailed calculations, assumptions and
methods used in designing and simulating the process equipment.
The material of construction for ALL units was preliminarily assumed to be
carbon steel because of its low cost, mechanical and chemical properties.
Pumps (P-101, 102, 103, and 104)
Pumps were designed and simulated using Aspen Plus via inputting the
desired pressure discharge. The pressure discharge was specified based on the
desired pressure for a certain stream. Aspen Plus calculates the break power and
pump efficiency.
Pump (P-101)
o Pump for flowing the feed methanol to the process
o Discharge pressure of 2.35 bar was chosen to accommodate for pressure drop in
the equipment and result in a pressure of 1.6 bar in stream 7 (Hong, 2008)
o The pump is to be a centrifugal pump with an electric drive; The reason for this
selection is because they are most common type of pumps used, and they are the
best choice for low viscosity and high flowrate (Turton, 2012).
o From Aspen Plus, break power = 36.4 kW, and the efficiency is 82%
Pump (P-102)
o Reflux pump for Distillation Column (T-101)
o Discharge pressure of 28 bar was chosen to accommodate for pressure drop in
the equipment
AKA 56
o The pump is to be a centrifugal pump with an electric drive; The reason for this
selection is because they are most common type of pumps used, and they are the
best choice for low viscosity and high flowrate (Turton, 2012).
o From Aspen Plus, break power = 37.6 kW, and the efficiency is 81%
Pump (P-103)
o Reflux pump for Distillation Column (T-102)
o Discharge pressure of 28 bar was chosen to accommodate for pressure drop in
the equipment
o The pump is to be a centrifugal pump with an electric drive; The reason for this
selection is because they are most common type of pumps used, and they are the
best choice for low viscosity and high flowrate (Turton, 2012).
o From Aspen Plus, break power = 17.1 kW, and the efficiency is 74%
Pump (P-104)
o Reflux pump for Distillation Column (T-103)
o Discharge pressure of 28 bar was chosen to accommodate for pressure drop in
the equipment
o The pump is to be a centrifugal pump with an electric drive; The reason for this
selection is because they are most common type of pumps used, and they are the
best choice for low viscosity and high flowrate (Turton, 2012).
o From Aspen Plus, break power = 48.0 kW, and the efficiency is 82%
AKA 57
Heat Exchangers (E-101 to E-113)
The heat exchangers in the plant were all simulated using the “HeatX” block
in Aspen Plus. The Exchanger specifications, pressure drop and heat transfer
coefficient “U” were specified based on the application of the heat exchanger. The
pressure drop and the heat transfer coefficient “U” were both determined based on
the application of the heat exchanger and the fluids passing through the shell and
tube sides using the heuristics (Turton, 2012). From the inputted information, Aspen
Plus calculates the area and heat duty of the heat exchanger. In general, the amount
of utility needed was determined via the sensitivity analysis function in Aspen Plus,
the sensitivity conditions vary per unit and utility type.
A simplified version of the “HeatX” block is the “Heater” block. In some
cases, this block was used in the main body of the plant simulation for simplification
purposes, all of which were designed as “HeatX” blocks separately to determine the
amount of utility needed, area of heat exchange, and heat duty.
Heat exchangers E-101 through E-113 (except for E-106) have been designed
with a floating head construction. The reasoning behind this selection is that a
floating head construction can handle thermal expansion and allows easier access to
the inner and outer tubes for cleaning purposes, since the bundle can be removed
(Turton,2012). Heat exchanger E-106 was designed with a double pipe construction
due to the exchange area being relatively small (between 1 – 10 m2).
AKA 58
Heat Exchanger (E-101)
o Feed goes in tube side of the heat exchanger
o Low pressure steam (5 barg, 160 oC) in the shell side of the heat exchanger
(Turton,2012)
o It is desired to vaporize the feed in this heat exchanger (Hong, 2008); the feed
should come out at 97 oC
o Pressure drop of 0.1 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 200 Btu/hr-ft2-oF (Turton,2012)
o Low pressure steam flowrate was determined using sensitivity analysis by
varying the flowrate of low pressure steam (input) and monitoring the
temperature of the output stream from the shell side such that it comes out as
boiler feed water (115 oC) (Turton,2012). This boiler feed water stream is to be
sold or reused in the plant. The low pressure steam flowrate = 118,721 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area =
1,650 m2 and Q = 447,992 MJ/hr
Heat Exchanger (E-102)
o Process stream goes in both shell and tube side of the heat exchanger
o Process stream #3 goes in tube side and #4 out of tube side, while process
stream #9 goes in shell side and #10 out of shell side
o The purpose of this heat exchanger is to superheat stream #3 to 266 oC (Hong,
2008) and cool down stream #9.
o Pressure drop of 0.2 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
AKA 59
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 133
m2 and Q = 106,285 MJ/hr
Heat Exchanger (E-103)
o Process stream goes in both shell and tube side of the heat exchanger
o Process stream #5 goes in shell side and #6 out of shell side, while process
stream #8 goes in tube side and #9 out of tube side
o The purpose of this heat exchanger is to heat stream #5 to 420 oC (Hong, 2008)
and cool down stream #8
o Pressure drop of 0.2 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 60 Btu/hr-ft2-oF (Turton,2012)
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 394
m2 and Q = 95,553 MJ/hr
Heat Exchanger (E-104)
o Feed goes in tube side of the heat exchanger
o Cooling water (3 bar, 30 oC) in the shell side of the heat exchanger (Turton,2012)
o It is desired to partially condense the feed in this heat exchanger to knockout
water in the flash separator; a sensitivity analysis was conducted by varying the
temperature output of this heat exchanger and monitoring the water fraction in
the liquid phase; based on sensitivity analysis, the feed should come out at 38 oC
o Pressure drop of 0.2 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o Cooling water flowrate was determined using sensitivity analysis by varying the
flowrate of cooling water (input) and monitoring the temperature of the output
AKA 60
stream from the shell side such that it comes out at 45 oC (Turton,2012). The
cooling water flowrate = 1.1*107 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area =
2,767 m2 and Q = 742,877 MJ/hr
Heat Exchanger (E-105)
o Due to Aspen Plus limitations, stream 12 was not sent directly to compressor C-
102; the stream had to be preheated prior to entering the compressor C-102
using this heat exchanger
o Feed goes in tube side of the heat exchanger
o High pressure steam (41 barg, 254 oC) in the shell side of the heat exchanger
(Turton,2012) was used to minimize area of heat exchanger and flow of steam.
o It is desired to heat up the feed in this heat exchanger; due to simulation
limitations, the feed should come out at 104 oC
o Pressure drop of 0.2 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o High pressure steam flowrate was determined using sensitivity analysis by
varying the flowrate of high pressure steam (input) and monitoring the
temperature of the output stream from the shell side such that it comes out as
boiler feed water (115 oC) (Turton,2012). This boiler feed water stream is to be
sold or reused in the plant. The high pressure steam flowrate = 18,568 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 136
m2 and Q = 45,728 MJ/hr
AKA 61
Heat Exchanger (E-106)
o Feed goes in tube side of the heat exchanger
o Cooling water (3 bar, 30 oC) in the shell side of the heat exchanger (Turton,2012)
o It is desired to cool down the feed in this heat exchanger but not condense it; this
is such that there is no need for liquid knockout prior to the second stage of
compression; the feed is to be cooled down to 132 oC, based on sensitivity
analysis of varying temperature of the feed output and monitoring the vapor
fraction.
o Pressure drop of 0.2 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o Cooling water flowrate was determined using sensitivity analysis by varying the
flowrate of cooling water (input) and monitoring the temperature of the output
stream from the shell side such that it comes out at 45 oC (Turton,2012). The
cooling water flowrate = 18,145 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 4.2
m2 and Q = 1,231 MJ/hr
Heat Exchanger (E-107)
o Feed goes in tube side of the heat exchanger
o Cooling water (3 bar, 30 oC) in the shell side of the heat exchanger (Turton,2012)
o It is desired to condense the feed in this heat exchanger to 75 oC (Lurgi, 2003)
o Pressure drop of 0.2 bar in both shell and tube side was assumed (Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
AKA 62
o Cooling water flowrate was determined using sensitivity analysis by varying the
flowrate of cooling water (input) and monitoring the temperature of the output
stream from the shell side such that it comes out at 45 oC (Turton,2012). The
cooling water flowrate = 1.05*106 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 298
m2 and Q = 70,956 MJ/hr
Heat Exchanger (E-108)
o Condenser for distillation column (T-101)
o Feed goes in tube side of the heat exchanger
o Cooling water (3 bar, 30 oC) in the shell side of the heat exchanger (Turton,2012)
o It is desired to condense the vapor rising from the top tray in this heat exchanger;
the desired temperature was obtained from the profile of the distillation column
(T-101), and the temperature was 58.3 oC.
o Pressure drop of 0.2 bar in shell and 0.007 bar in tube side was assumed
(Turton,2012); the tube side pressure drop is equal to the pressure drop between
the trays of the distillation column, since tray 1 in aspen represents the condenser
and tray 2 is the top try of the column.
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o Cooling water flowrate was determined using sensitivity analysis by varying the
flowrate of cooling water (input) and monitoring the temperature of the output
stream from the shell side such that it comes out at 45 oC (Turton,2012). The
cooling water flowrate = 792,672 kg/hr
AKA 63
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 572
m2 and Q = 52,871 MJ/hr
Heat Exchanger (E-109)
o Reboiler for distillation column (T-101)
o Feed goes in tube side of the heat exchanger
o High pressure steam (41 barg, 254 oC) in the shell side of the heat exchanger
(Turton,2012)
o It is desired to re-boil the liquid dropping from the bottom tray in this heat
exchanger; the desired temperature was obtained from the profile of the
distillation column (T-101), and the temperature was 236.4 oC.
o Pressure drop of 0.1 bar in shell side was assumed (Turton,2012); and a pressure
buildup 0.007 bar was assumed in the tube side, since the pressure drops from
the reboiler pressure to the condenser pressure by 0.007 bar per tray
(Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o High pressure steam flowrate was determined using sensitivity analysis by
varying the flowrate of high pressure steam (input) and monitoring the
temperature of the output stream from the tube side such that the desired
temperature is achieved with the minimum amount of steam. The high pressure
steam flowrate = 13,133 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area =
1,329 m2 and Q = 25,176 MJ/hr
AKA 64
Heat Exchanger (E-110)
o Condenser for distillation column (T-102)
o Feed goes in tube side of the heat exchanger
o 50/50 Water-Ethylene glycol refrigerant (2 bar, -30 oC) in the shell side of the
heat exchanger (Engineering Toolbox)
o It is desired to condense the vapor rising from the top tray in this heat exchanger;
the desired temperature was obtained from the profile of the distillation column
(T-102), and the temperature was -20.5 oC.
o Pressure drop of 0.2 bar in shell and 0.007 bar in tube side was assumed
(Turton,2012); the tube side pressure drop is equal to the pressure drop between
the trays of the distillation column, since tray 1 in aspen represents the condenser
and tray 2 is the top try of the column.
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o Refrigerant flowrate was determined using sensitivity analysis by varying the
flowrate of cooling water (input) and monitoring the temperature of the output
stream from the tube side to find the minimum amount of refrigerant. The
refrigerant flowrate = 684,649 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area =
1,635 m2 and Q = 19,238 MJ/hr
AKA 65
Heat Exchanger (E-111)
o Reboiler for distillation column (T-102)
o Feed goes in tube side of the heat exchanger
o Low pressure steam (5 barg, 160 oC) in the shell side of the heat exchanger
(Turton,2012)
o It is desired to re-boil the liquid dropping from the bottom tray in this heat
exchanger; the desired temperature was obtained from the profile of the
distillation column (T-102), and the temperature was 74.9 oC.
o Pressure drop of 0.1 bar in shell side was assumed (Turton,2012); and a pressure
buildup 0.007 bar was assumed in the tube side, since the pressure drops from
the reboiler pressure to the condenser pressure by 0.007 bar per tray
(Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton,2012)
o Low pressure steam flowrate was determined using sensitivity analysis by
varying the flowrate of low pressure steam (input) and monitoring the
temperature of the output stream from the tube side such that the desired
temperature is achieved with the minimum amount of steam. The low pressure
steam flowrate = 793 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 25
m2 and Q = 2,024 MJ/hr
AKA 66
Heat Exchanger (E-112)
o Condenser for distillation column (T-103)
o Feed goes in tube side of the heat exchanger
o Cooling water (3 bar, 30 oC) in the shell side of the heat exchanger (Turton,2012)
o It is desired to condense the vapor rising from the top tray in this heat exchanger;
the desired temperature was obtained from the profile of the distillation column
(T-101), and the temperature was 59.4 oC.
o Pressure drop of 0.2 bar in shell and 0.007 bar in tube side was assumed
(Turton,2012)
o Heat transfer coefficient “U” was assumed 200 Btu/hr-ft2-oF (Turton,2012)
o Cooling water flowrate was determined using sensitivity analysis by varying the
flowrate of cooling water (input) and monitoring the temperature of the output
stream from the shell side such that it comes out at 45 oC (Turton,2012). The
cooling water flowrate = 761,092 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 600
m2 and Q = 51,621 MJ/hr
Heat Exchanger (E-113)
o Reboiler for distillation column (T-103)
o Feed goes in tube side of the heat exchanger
o Low pressure steam (5 barg, 160 oC) in the shell side of the heat exchanger
(Turton,2012)
AKA 67
o It is desired to re-boil the liquid dropping from the bottom tray in this heat
exchanger; the desired temperature was obtained from the profile of the
distillation column (T-103), and the temperature was 117.7 oC.
o Pressure drop of 0.1 bar in shell side was assumed (Turton,2012); and a pressure
buildup 0.007 bar was assumed in the tube side, since the pressure drops from
the reboiler pressure to the condenser pressure by 0.007 bar per tray
(Turton,2012)
o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton, 2012)
o Low pressure steam flowrate was determined using sensitivity analysis by
varying the flowrate of low pressure steam (input) and monitoring the
temperature of the output stream from the tube side such that the desired
temperature is achieved with the minimum amount of steam. The low pressure
steam flowrate = 450 kg/hr
o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 21
m2 and Q = 1,056 MJ/hr
Reactor
DME Reactor (R-101)
The reactor was simulated in Aspen Plus using an “REquil” block. The
reaction inputted was as follow (Farsi, 2010):
2𝐶𝐻3𝑂𝐻 ⟺ 𝐶𝐻3𝑂𝐶𝐻3 + 𝐻2𝑂
The reactor operates isothermally at 300 oC (Hong, 2008), and is modeled as a heat
exchanger with the feed going into the tube side, where the aluminum oxide catalyst
is packed (Lurgi, 2003), and cooling water in the shell side to maintain the reactor
AKA 68
temperature (Farsi, 2010). A pressure drop of 0.35 bar was assumed to reach to the
desired pressure in stream 7. Furthermore, the number of tubes, diameter of tube,
and reactor length were all obtained from literature (Farsi, 2010); they were used to
find the volume of the reactor for designing and costing purposes.
o Number of tubes = 2000
o Tube diameter = 0.09 m
o Length of reactor = 8 m
o Volume of reactor = 2000 × 8 ×𝜋(0.09)2
4= 102 𝑚3
o Form Aspen Plus, the estimated amount of cooling water needed in the shell
side of this reactor is 3.65*106 kg/hr
MTP Reactor (R-102)
The reactor was simulated in Aspen Plus using “RStoic” block since the
reaction kinetics were unavailable. This reactor was a challenge to simulate due to
the unavailability of reaction kinetics, yet several attempts to accurately simulate it
were attempted (refer to appendix D for further information about the different
attempts tackled in designing this reactor). After several attempts, an assumption
had to be made to simplify the simulation. This assumption was that only the major
reactions are happening in this reactor and those reactions are the ones producing
the major product. The products are Ethylene, Propylene, Butene, Pentene, Hexene,
Heptene, Octene, and Water. Due to the limitations of Aspen Plus, two “RStoic”
blocks in series (R-102A and R-102B) were used to model the single MTP reactor.
The reactions inputted into the first reactor (R-102A) are as follow:
General Form: 𝑛𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶𝑛𝐻2𝑛 + 𝑛𝐻2𝑂 𝑓𝑜𝑟 𝑛 = 2, … ,8 (Meyers, 2005)
AKA 69
Table A1: List of Reactions inputted in MTP Reactor (R-102A)
n Reaction
2 2𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶2𝐻4 + 2𝐻2𝑂
3 3𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶3𝐻6 + 3𝐻2𝑂
4 4𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶4𝐻8 + 4𝐻2𝑂
5 5𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶5𝐻10 + 5𝐻2𝑂
6 6𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶6𝐻12 + 6𝐻2𝑂
7 7𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶7𝐻14 + 7𝐻2𝑂
8 8𝐶𝐻3𝑂𝐶𝐻3 → 2𝐶8𝐻16 + 8𝐻2𝑂
The reactions inputted into the second reactor (R-102B) are as follow:
General Form: 𝑛𝐶𝐻3𝑂𝐻 → 𝐶𝑛𝐻2𝑛 + 𝑛𝐻2𝑂 𝑓𝑜𝑟 𝑛 = 2, … ,8 (Hadi, 2014)
Table A2: List of Reactions inputted in MTP Reactor (R-102B)
n Reaction
2 2𝐶𝐻3𝑂𝐻 → 𝐶2𝐻4 + 2𝐻2𝑂
3 3𝐶𝐻3𝑂𝐻 → 𝐶3𝐻6 + 3𝐻2𝑂
4 4𝐶𝐻3𝑂𝐻 → 𝐶4𝐻8 + 4𝐻2𝑂
5 5𝐶𝐻3𝑂𝐻 → 𝐶5𝐻10 + 5𝐻2𝑂
6 6𝐶𝐻3𝑂𝐻 → 𝐶6𝐻12 + 6𝐻2𝑂
7 7𝐶𝐻3𝑂𝐻 → 𝐶7𝐻14 + 7𝐻2𝑂
8 8𝐶𝐻3𝑂𝐻 → 𝐶8𝐻16 + 8𝐻2𝑂
The reactor is designed as a fixed bed process vessel with mordenite zeolite,
HMOR, as the catalyst. The reactor operates isothermally at 452 oC (Hong, 2008). A
cooling jacket, with cooling water flowing in it, was planned to maintain the
temperature of the reactor. A pressure drop of 0.9 bar was assumed across the MTP
reactor, 0.45 bar in both reactors R-102A & R-102B, to reach the desired output
pressure of 1.6 bar in the exiting stream, stream 7. The volume of the reactor was
determined using the weighted hourly space velocity “WHSV” (Moreno-Pirajan,
2013); then using the heuristics from Turton, the length to diameter ratio was chosen,
and both values were calculated.
AKA 70
o WHSV = 1 hr-1 = 𝑀𝑎𝑠𝑠 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒 𝑜𝑓 𝑓𝑒𝑒𝑑
𝑊𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡=
350,000𝑘𝑔
ℎ𝑟
𝑊𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡
→ 𝑊𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 = 350,000 𝑘𝑔
o Mordenite zeolite density = 2135 kg/m3 (Mindat.org)
o Volume of reactor = 𝑊𝑒𝑖𝑔ℎ𝑡 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡
𝐷𝑒𝑛𝑠𝑖𝑡𝑦 𝑜𝑓 𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡=
350,000 𝑘𝑔
2135 𝑘𝑔/𝑚3= 164 𝑚3
o L/d ratio = 3, in other words L = 3 d (Turton, 2012)
o Diameter of reactor: 164 𝑚3 = (3 𝑑) ×𝜋(𝑑)2
4→ 𝑑 = 4.1𝑚
o Length of reactor: L = 3 d = 12.3 m
o Form Aspen Plus, the estimated amount of cooling water needed in the cooling
jacket is 13.5*106 kg/hr
Compressors (C-101, 102, and 103)
The compressors in the plant were simulated using the “Comp” block in Aspen Plus.
Based on Turton’s heuristics, the chosen compressor type for all compressors is
isentropic. All compressors are reciprocating compressors simply because the
required head is so high than an undesirable large number of stages needed
(Turton,2012). Isentropic type of compressor was selected because the compression
is taking place with no flow of heat energy either into or out of the gas (Turton,2012).
The discharge pressure and the efficiency for the compressors were determined
based on the process and heuristics from Turton, respectively. The efficiency in C-
101 is 85% because the compression ratio is roughly 3.8. The efficiency in C-102 and
C-103 is 75 % because the compression ratio is roughly 2.2 (Turton,2012). Aspen
Plus is then able to calculate the break horsepower that can be then used in costing
AKA 71
the compressors. The table below shows all the compressors with their
corresponding parameters.
Table A3: design compressors and parameters
Compressor
Inlet
Pressure
(bar)
Discharge
Pressure
(bar)
Pressure
Ratio
Break
horsepower
(kW)
Efficiency
C-101 1.6 6.1 3.8 39,299 85 %
C-102 5.3 11.6 2.2 2,425 75 %
C-103 11.5 25.2 2.2 2,400 75 %
Flash Separator Design (V-101)
The flash separator in the plant was designed and simulated in Aspen Plus.
The input of the flash separator V-101, stream (11) was cooled down prior to enter
the flash separator at 38 oC (100 oF) and 5.5 bar. It was assumed that the flash duty is
zero. In addition, it was assumed that the flash operates at 5.5 bar similar the feed
pressure to achieve the desired purities. Water was knocked out with purity of
(99.97 mole faction). Waste water goes to nearby treatment facility for further
purification due to the methanol contamination. All hydrocarbons are leaving the
flash separator from the top. The flash separator was designed and simulated as
vertical based on heuristics in Turton.
o Assuming holdup time for half full = 5 min
o Flow rate in = 195565 𝐿
𝑚𝑖𝑛
o Assuming 𝐿
𝑑= 3
195565 𝐿
𝑚𝑖𝑛 ×
1 𝑚3
1000 𝐿 = 195.567
𝑚3
𝑚𝑖𝑛
Volume of the vessel = 195.567 𝑚3
𝑚𝑖𝑛 × 5 min × 2 = 1955.65 𝑚3
AKA 72
Volume = 𝜋 𝑑2
4 × 𝐿 =
𝜋 𝑑2
4 × 3𝑑 = 1955.65 𝑚3
From the volume of the flash: diameter = 9.39784 m and height = 28.1935 m
Distillation Columns (T-101, 102, and 103)
The distillation columns in the plant were simulated using the “RadFrac” block
in Aspen Plus. The amount of distillate or bottom product was determined based on
the amount of product in the feed going into the distillation column. The reflux ratio
was varied and determined using Aspen to obtain the desired purity. The optimum
number of stages, feed stage, and reflux ratio were then determined using the
optimization method from (Whitlow, 2016). Furthermore, the condenser pressure
was determined from the process; however, the pressure drop was determined from
Turton’s heuristics. The type of trays used in all distillation columns are sieve trays,
mainly because they have higher entrainment than other types of trays (Pilling,
2012). From inputting the type of tray in Aspen Plus and estimating a number of
passes, an estimated dimeter for the column can be obtained. Using the maximum
liquid flowrate and the estimated diameter from Aspen Plus in Figure #13.7 from the
Koch Flexitray Design Manual (Whitlow, 2016), the number of passes can be
confirmed, thus the exact diameter of the column can be obtained. The tray spacing,
tray efficiency, and height of the distillation columns were determined using
Turton’s heuristics as well. The following is the sample calculation of determining the
number of passes in the first distillation column (T101).
o Maximum liquid flow rate found to be 11574.9 gal/min on stage 32
o Figure 13.7 from the Koch Flexitray Design Manual (Whitlow, 2016) was
used to determine the number of passes.
AKA 73
In the sizing and rating option, the number of passes and type of trays were
the inputs. The column diameter found to be 4.8m using Aspen Plus. After that, it was
assumed that 2 foot tray spacing in distillation columns. Additional 20 % of the total
height was added to the distillation column. The height of the column was estimated
by L = 1.2 (NT – 2) × 2 where NT is the number of trays. Below is a sample calculation
of T-101 height.
L = 1.2 (NT – 2) × 2
L = 1.2 (46 – 2) × 2
L = 1.5.6 ft = 32.2 m
The volume of the distillation column was calculated using the following
formula: Volume = 𝜋 𝑑2
4 × 𝐿
The following is sample calculation of the volume for the first distillation
column T-101:
Volume = 𝜋 (4.81𝑚)2
4 × 32.2 𝑚 = 585 m3
There are some specific procedures to finalize the column design such as
minimizing the capital cost by reducing the theoretical number of stages. In
addition, the reflux ratio and reboiler duty minimize the operating cost. The
following is a detailed explanation of distillation columns optimization using T-102 as
example. The optimization study was similar for the rest of the columns.
1. Varying the number of stages and monitoring the reflux ratio
The following is example of T-102
AKA 74
Table A4: Finding the Optimum Reflux Ratio
Number of Stages Feed Stage Reflux Ratio
20 10 21.606
30 15 11.902
50 25 8.600
60 30 8.023
70 35 7.690
75 36 7.531
85 38 7.367
95 41 7.319
It could be noticed that the reflux ratio is almost stable at 7.319. Based on the
heuristics in Turton the economical optimal reflux ratio is 1.2 higher than the found
value. The minimum reflux ration is found to be 8.782
2. Varying the number of stages to minimum
This step basically was done by reducing the number of stages until aspen
crashes and then calculate the economic optimum number of stages which is nearly
twice the minimum value. The flowing is an example of the case study on T-102.
Table A5: Finding the Optimum Number of Stages
Number of Stages Feed Stages Reflux Ratio
17 8 51.083
18 9 33.119
25 12 13.838
35 11 8.970
The minimum number of stages found to be 17. The theoretical number of
stages is roughly 35 stage.
3. Varying the feed stage and monitoring the heat of reboiler and condenser
The feed stage was varied to find the minimum reboiler and condenser heat
duty. The following table shows the study for T-102.
AKA 75
Table A6: Finding the Optimum Feed Stage
Feed Stage Reboiler Heat (MW) Condenser Heat (MW) Reflux Ratio
10 6.2235 -5.38955 9.054
11* 6.17844 -5.3449 8.970
12 6.21465 -5.38067 9.038
13 6.29864 -5.46319 9.190
*Optimum
The optimum feed tray found to be on stage 11 because it typically minimizes
the reboiler and conducer duty required.
Reflux Drums (V-102, 103, and 104)
Reflux drums were designed and simulated in Aspen Plus as “Flash2”. It was
assumed that no duty is taking place in the flash and the pressure in the flash is
nearly close to the stream pressure. In the main simulation part, the reflux drum is
not shown. Reflux drums are designed and simulated in the condensers modeling
part for all the distillation columns. The following assumptions were addressed when
designing the reflux drums.
o Horizontal vessels
o Half full holdup time = 5 min
o Assuming 𝐿
𝑑= 3
The volumes of the reflux drums were calculated as follow:
o 6270.1 𝐿
𝑚𝑖𝑛 ×
1 𝑚3
1000 𝐿 × 5 min × 2 = 72.7 m3
o Volume of reflux drums = 𝜋 𝑑2
4 × 𝐿 =
𝜋 𝑑2
4 × 3𝑑 = 72.7 m3
from the volume of the drum: diameter = 2.97 m; height = 8.93 m
AKA 76
The assumptions and calculations that were addressed above are also followed to
design and simulate all the reflux drums in the process.
Storage Tanks (V-105 to V-109)
Storage tanks were not simulated on Aspen Plus, yet they were designed based on
the product flowrate, and heuristics. Heuristics were used to determine the capacity
and orientation of the tanks. Using the capacity and the flow rate, the volume of the
tank can be determined, which was then used for designing and costing purposes.
o Capacity = 3 days
o Storage tanks are operating on 3 bar
o Assuming 𝐿
𝑑= 3
Volume, diameter, and height were calculated as follow:
Flow rate of propylene = 2159.856 𝐿
𝑚𝑖𝑛
2159.856 𝐿
𝑚𝑖𝑛 ×
1 𝑚3
1000 𝐿 60 𝑚𝑖𝑛
1 ℎ𝑟 ×
24 ℎ𝑟
1 𝑑𝑎𝑦 × 3 days = 9330.6 m3
Volume of storage tank = 𝜋 𝑑2
4 × 𝐿 =
𝜋 𝑑2
4 × 3𝑑 = 9330.6 m3
From the volume of the tank: diameter = 18.1m and height = 36.2 m
The same calculations are applied for all the storage tanks. The table below shows
all the parameters for each storage tank.
Table A7: Storage Tanks Parameters
Equipment Volume(m3) Diameter(m) Height (m)
Methanol Storage Tank 32784 27.53 55
Propylene Storage Tank 9331 18.1 36.2
LPG Storage Tank 5555 13.3 39.9
Fuel Gas Storage Tank 1374 8.98 17.9
Gasoline Storage Tank 10074 18.5 37.15
AKA 77
Appendix B: Capital Cost Sample Calculations
Calculation for Pump P-101 A/B Bare Module Cost:
Table B1: Bare Module Pump Costing Coefficients
Centrifugal
Capacity
(kw)
Discharge
pressure
(bar)
K1 K2 K3 B1 B2 C1 C2 C3 FM
(Carbon Steel)
36.4 3.685 3.3892 0.0536 0.1538 1.89 1.35 0 0 0 1.6
From Table B1, Log (Cpo) and Log (Fp) can be calculated using the following
equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 3.3892 + 0.0536 × log10(36.4) + 0.1538 × (log10(36.4))2 = 3.85
Cpo = 7041.95
For the P<10, C1, C2, and C3 will be zero
log10Fp = C1 + C2 log10(P) + C3 [log10(P)]2 = 0
Fp = 1
CBM = Cpo(B1 + B2FMFp) = 7041.95 × (1.89 + (1.35 × 1.6 × 1)) = $28,520
Calculation for Vessel V-101:
Table B2: Bare Module Process Vessel Costing Coefficients
Process vessel Vertical
Diameter
(M)
Capacity
(𝑚3) K1 K2 K3 B1 B2
FM
(Carbon Steel)
Pressure
(barg)
9.398 488.92 3.4974 0.4485 0.1074 2.25 1.82 1.6 4.944
From Table B2, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 3.4974 + 0.4485 × log10(488.92) + 0.1074 × (log10(488.92))2 = 5.48
Cpo = 302160.1
FP,Vessel
(P + 1) ∗ D2[850 − 0.6(P + 1)]
+ 0.00315
0.0063=
(4.944 + 1) × 9.3982[850 − 0.6(4.944 + 1)]
+ 0.00315
0.0063= 5.7
CBM = Cpo(B1 + B2FMFp) = 302160.1 × (2.25 + (1.82 × 1 × 1)) = $3,835,378
AKA 78
Calculation for Heat Exchanger H-103:
Table B3: Bare Module Heat Exchanger Costing Coefficients
Floating Head
Capacity
(𝑚3) Q (
𝑀𝐽
ℎ) K1 K2 K3 B1 B2
FM
(Carbon Steel)
Pressure
(barg)
394.19 106285 4.8306 0.851 0.3187 1.63 1.66 1 5.61
From Table B3, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 4.8306 + (−0.851) × log10(394.19) + 0.3187 × (log10(394.19))2 = 4.769
Cpo = 58777
For the P<10, C1, C2, and C3 will be zero (Turton, 2012)
log10Fp = C1 + C2 log10(P) + C3 [log10(P)]2 = 0
Fp = 1
CBM = Cpo(B1 + B2FMFp) = 58777 × (1.63 + (1.66 × 1 × 1)) = $193,378
Calculation for Compressors C-102:
Table B4: Bare Module Compressor Costing Coefficients
Reciprocating
Power
(𝐾𝑊)
Pressure
(barg) K1 K2 K3
FBM
(Carbon
Steel)
2425 11.7 2.2897 1.3604 -0.103 3.4
From Table B4, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 2.2897 + 1.3604 × log10(2425) + (−0.103) × (log10(2425))2 = 5.72
Cpo = $ 522008.6
For the P<10, C1, C2, and C3 will be zero
log10Fp = C1 + C2 log10(P) + C3 [log10(P)]2 = 0
Fp = 1
CBM = Cpo(FBM) = 522008.6 × 3.4 = $ 1,774,829
AKA 79
Calculation for Column T-103:
Table B5: Bare Module Column Costing Coefficients
Tray Vertical
Towers
Diameter
(M)
Capacity
(𝑚3) K1 K2 K3 B1 B2
FM
(Carbon Steel)
Pressure
(barg)
3.30 300 3.4974 0.4485 0.1074 2.25 1.82 1 26.4
From Table B5, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 3.4974 + 0.4485 × log10(300) + 0.1074 × (log10(300))2 = 5.3
Cpo = $ 184885.871
FP,Vessel
(P + 1) ∗ D2[850 − 0.6(P + 1)]
+ 0.00315
0.0063=
(26.4 + 1) × 3.302[850 − 0.6(26.4 + 1)]
+ 0.00315
0.0063= 9.1
CBM = Cpo(B1 + B2FMFp) = 184885.871 × (2.25 + (1.82 × 1 × 13.1)) = $ 3,477,566
Calculation for Sieve Trays T-103:
Table B6: Bare Module Sieve Trays Costing Coefficients
Sieve Trays
Diameter
(M)
Capacity
(𝑚2) K1 K2 K3 N( # Trays)
𝐹𝐵𝑀
(Carbon Steel) Fq
3.30 8.53 2.9949 0.4465 0.3961 48 1 1
Where is:
N: Number of trays
The quantity factor for trays, Fq, for N ≥ 20: Fq = 1
From Table B6, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 2.9949 + 0.4465 × log10(8.53) + 0.3961 × (log10(8.53))2 = 3.75
Cpo = $ 5676.5
CBM = CpoNFBMFq = 5676.5 × 48 × 1 × 1 = $ 272,471
AKA 80
Calculation for Reactor R-101:
Table B7: Bare Module Reactor Costing Coefficients
Floating Head
Tube
Diameter(M)
N
(# of Tube )
Capacity
(𝑚2) K1 K2 K3 B1 B2 𝐹𝑀
Pressure
(barg)
0.09 2000 12.72 4.8306 -0.8509 0.3187 1.63 1.66 1 1.87
Note: This reactor was design as heat exchanger.
From Table B7, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 4.8306 + (−0.8509) × log10(12.72) + 0.3187 × (log10(12.72))2 = 4.3
Cpo = $ 19035.1
For the P<10, C1, C2, and C3 will be zero
log10Fp = C1 + C2 log10(P) + C3 [log10(P)]2 = 0
Fp = 1
CBM = Cpo(B1 + B2FMFp) = 19035.1 × (1.63 + (1.66 × 1 × 1)) = $ 62,625
Calculation for Reactor R-102:
Table B8: Bare Module Reactor Costing Coefficients
Process Vessel Vertical
Diameter
(M)
Capacity
(𝑚3) K1 K2 K3 B1 B2
FM
(Carbon Steel)
Pressure
(barg)
4.1 164 3.4974 0.4485 0.1074 2.25 1.82 1 0.66
Note: This reactor was design as Process Vessel.
From Table B8, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 3.4974 + 0.4485 × log10(164) + 0.1074 × (log10(164))2 = 5.02
Cpo = $ 104138.9
FP,Vessel
(P + 1) ∗ D2[850 − 0.6(P + 1)]
+ 0.00315
0.0063=
(0.66 + 1) × 4.12[850 − 0.6(0.66 + 1)]
+ 0.00315
0.0063= 1.138
CBM = Cpo(B1 + B2FMFp) = 184885.871 × (2.25 + (1.82 × 1 × 1.38)) = $ 562,578
AKA 81
Calculation for Storage Tank V-107:
Table B9: Bare Module Tank Costing Coefficients
Tanks Vertical
Towers
Diameter
(M)
Capacity
(𝑚3) K1 K2 K3 B1 B2
FM
(Carbon Steel)
Pressure
(barg)
18.1 9331 5.9567 -0.7585 0.1749 2.25 1.82 1 2.25
From Table B9, Log (Cpo) and Fp can be calculated using the following equations:
log10Cpo = K1 + K2 log10(A) + K3 [log10(A)]2
= 5.9567 + (−0.7585) × log10(9331) + 0.1749 × (log10(9331))2 = 5.7
Cpo = $ 503476
FP,Vessel
(P + 1) ∗ D2[850 − 0.6(P + 1)]
+ 0.00315
0.0063=
(2.25 + 1) × 18.12[850 − 0.6(2.25 + 1)]
+ 0.00315
0.0063= 5.9
CBM = Cpo(B1 + B2FMFp) = 503476 × (2.25 + (1.82 × 1 × 5.9)) = $ 6,557,733
AKA 82
Appendix C: Manufacturing Cost Sample Calculations
Operating labor Cost Calculation:
NOL = √6.29 + 31.7P2 + 0.23Nnp
P= 0
Nnp= 22
NOL = √6.29 + 31.702 + 0.23 × 22 = 3.369 ≈ 4
NO = 4.5 × NOL
= 4.5 × 4 = 18
COL = NO × AW
AW= 53,500
𝐶𝑂𝐿 = 18 × $ 53,500 = $ 963,900
As shown in Table C1, there are 52 weeks in one year, 3 weeks for vacation, 8
hours in each shift, and 5 shift per week. To calculate the total hours per year, the
number of week should be 49 by subtracting 52 from 3. Calculating Total hour is by
Multiplying 8 × 5 × 49 = 1960 and total hour/year is 24 × 350 = 8400. The number of
operator hired for each operators is 8400
1960 = 4.3 and this number is rounded to 4.5.
Table C1: Operating labor cost variable
Number of week in year 52
Number of weeks for vacation 3
Number of shift per week 5
Hours each shift 8
Total hours /year 1960
Total hours of operation 8400
Number of operator hired for each
operator 4.5
NOL 4
Annual mean Wages in 2016 53,550
Total number of operators 20
COL $1,071,000
AKA 83
Calculation for Waste Water Cost:
CWT = Flow rate (MT
YR) × Price (
$
MT)
Flowrate = 1645837.2 MT
YR
Price = 0.0625 $
MT
CWT = 1645837.2 × 0.0625
Calculation for utility Cost
Calculation for Electricity Cost
The price of Electricity is 0.0718 $
KW−hr
The total amount of electricity needed is 367,045,401 KW−hr
yr
CElectricity = amount of eletricity needed × price of eletricity
CElectricity = 367,045,401 × 0.0718 = 26,535,860 $
𝑦𝑟
Calculation for Cooling water cost
The price of cooling water is 0.0000175 $
kg
The total amount of Cooling Water needed is 258,002,773,140 kg
yr
CCW = amount of CW needed × price of CW
CCW = 258,002,773,140 × 0.0000175 = 4,502,697 $
𝑦𝑟
Calculation for High Pressure Steam cost:
The price of HPS is 0.01459 $
kg
The total amount of HPS is 266,292,516 kg
yr
CHPS = amount of HPS needed × price of HPS
CHPS = 266,292,516 × 0.01459 = 3,885,247 $
𝑦𝑟
AKA 84
Calculation for Low Pressure Steam
The price of LPS is 0.01348 $
kg
The total amount of LPS is 1,595,689,056 kg
yr
CLPS = amount of LPS needed × price of LPS
CLPS = 1,595,689,056 × 0.01348 = 21,516,967$
𝑦𝑟
Calculation for Refrigerant Cost:
The price of Refrigerant Duty is 11.2671 $
GJ
The total amount of Refrigerant Duty is 161,597 GJ
yr
CRefrigerant = amount of Refrigerant duty needed × price of Refrigerant Duty
CRefrigerant = 161,597 × 11.2671 = 1,820,726 $
𝑦𝑟
Total Utility cost calculation:
CUT= C Electricity + CCW + CHPS + CLPS + C Refrigerant
CUT= 26,535,860 + 4,502,697 + 3,885,247 + 21,516,967 + 1,820,726
CUT= $ 58,079,297
Calculation For Raw Material Cost:
Feed of the process:
The Amount of Methanol needed is 2640000 MT
yr
The price of Methanol is 235 $
kg
CMethanol = amount of Methanol needed × price of Methanol
CMethanol = 2640000 × 23 = $690,900,000
R-101 ( Aluminum Oxide Catalyst)
The Mass of Aluminum Oxide Catalyst needed for the R-101 is 402.9 MT
The price of Aluminum Oxide catalyst is 1000 $
MT
AKA 85
CAl2O3 Catalyst = Mass of Al2O3 catalyst needed × price of Al2O3 Catalyst
CAl2O3 Catalyst = 402.9 × 1000 = $ 402,900
R-102 ( Mordenite Zeolite Catalyst)
The Mass of Mordenite Zeolite catalyst (HMOR) needed for the R-102 is 344.4 MT
The price of Mordenite Zeolite catalyst (HMOR) is 120 $
MT
CHMOR = Mass of HMOR catalyst needed × price of HMOR Catalyst
CHMOR Catalyst = 344.4 × 120 = $ 41.328
Total Row Material Cost
CRM = CMethanol + CAl2O3 Catalyst + CHMOR = $ 691,344,228
Calculation for Cost of Manufacturing (COM):
COM = 0.180 × FCI + 2.73 × COL + 1.23(CUT + CWT + CRM)
COM = 0.180 × $175,416,007 + 2.73 × $ 963,900 + 1.23($ 58,079,897 + $ 102,834
+ $691,344,228) = $ 965,123,749
Where FCI is the fixed capital cost, it was discussed in previous section (appendix B)
Direct Manufacturing Costs (DMC):
𝐷𝑀𝐶 = CRM + CWT + CUT + 1.33 × COL + 0.03 × 𝐶𝑂𝑀 + 0.069 × 𝐹𝐶𝐼
𝐷𝑀𝐶 = $ 791,595,762
Direct Supervisory and Clerical Labor:
Direct Supervisory and Clerical Labor = 0.18 × COL = $ 173,502
Maintenance and Repairs:
Maintenance and Repairs was calculated using the formula below.
Maintenance and Repairs = 0.06 × FCI = $ 10,524,960
Fixed Manufacturing Cost (FMC):
Fixed Manufacturing Cost = 0.708 × COL + 0.068 × FCI = $ 12,610,729
Local taxes and Insurance:
Local taxes and Insurance = 0.032 × FCI = $ 5,613,312
AKA 86
Plant Overhead costs:
Plant Overhead costs = 0.708 × COL + 0.036 × FCI = $ 6,997,417
Lab Charges:
Lab Charges = 0.15 × COL = $ 144,585
Patents and Royalties:
Patents and Royalties = 0.03 × COM = $ 28,683,712
General Expenses:
General Expenses = 0.177 × COL + 0.009 × FCI + 0.16 × (COM) = $ 154,729,154
Administration Costs:
Administration Costs = 0.177 × COL + 0.009 × ( FCI) = $ 1,749,354
Distribution and Selling Costs:
Distribution and Selling Costs = 0.11 × COM = $ 105,173,612
Research and Development:
Research and Development = 0.05 × (COM) = $ 47,806,187
Table C2: Electricity Price Extrapolation
year cent/kW-
h
2005 5.73
2006 6.16
2007 6.39
2008 6.96
2009 6.83
2010 6.77
2011 6.82
2012 6.67
2013 6.89
2014 7.01
2015 6.89
2016 7.18
y = 0.0919x - 178.09
R² = 0.5915
5.0
5.5
6.0
6.5
7.0
7.5
2004 2006 2008 2010 2012 2014 2016
Pri
ce in
Cen
ts/k
W-h
r
Year
Price of Electricity
AKA 87
The data in Table C2 was used to plot the correlation in the figure above and
extrapolate linearly to find the cost of electricity in 2016 (Electric Power Monthly,
2015).
Table C3: Natural Price Data (Annual Energy Outlook 2015, 2015)
Natural Gas
year Price ($/MBtu)
2005 10.08
2006 7.58
2007 7.64
2008 9.53
2009 4.21
2010 4.61
2011 4.13
2012 2.79
2013 3.73
2014 4.37
2015 369.00%
Total Annual Change -51.32%
Annual Change -5.70%
The data in Table C3 was used to estimate the cost of natural gas in 2016 via
extrapolating the data from the annual energy outlook.
AKA 88
Appendix D: Profitability Analysis Sample Calculations
Products Selling Price Calculation:
To get the revenue calculation, the products price need to be found and also the
amount annually that the plant will produce to get the total price.
Ethylene:
The amount annually of Ethylene produced is 55405.81 MT
yr , and the price of Ethylene
is 1000 $
MT
Selling price of Ethylene= amount annually of Ethylene produced × price of
ethylene
Selling price of Ethylene = 55405.81× 1000 = 55,405,812 $
MT
Propylene:
The amount annually of Propylene produced is 477668.52 MT
yr , and the price of
Propylene is 1250 $
MT
Selling price of Propylene = amount annually of Propylene produced × price of
Propylene
Selling price of Propylene = 477668.52 × 1250 = 597,085,650 $
MT
Gasoline:
The amount annually of Gasoline produced is 1175309856 L
yr , and the price of
Gasoline is 0.3 $
L
Selling price of Gasoline = amount annually of Gasoline produced × price of
Gasoline
AKA 89
Selling price of Gasoline = 1175309856 × 0.3 = 352,592,957 $
MT
LPG:
The amount annually of LPG produced is 284446 MT
yr , and the price of LPG is 45
$
MT
Selling price of LPG = amount annually of LPG produced × price of LPG
Selling price of LPG = 284446 × 45 = 12,800,070 $
MT
Boiling Feed Water (BFW):
The amount annually of BFW produced is 1741230 MT
yr , and the price of BFW is 2.45
$
MT
Selling price of BFW = amount annually of BFW produced × price of BFW
Selling price of BFW = 284446 × 45 = 4,266,015 $
MT
Revenue per Year Calculation:
Revenue = ∑ after all products sold
Revenue = $1,022,150,504
Depreciation Calculation (Year 3):
Depreciation = 10 % × Fixed Capital Cost(FCI)
Depreciation = 10 % × $ 175,416,007 = $ 17,541,601
Taxes (Year 3):
Taxes = (Revenue − Operating Cost − Depreciation) × Tax rate
Taxes = ( $ 1,022,150,504 − $ 956,123,749 − $ 17,541,601) × 0.20 = $9,697,031
Net Profit (Year 3):
Net Profit = (Revenue − Operating Cost − Depreciation − Taxes)
= ($ 1,022,150,504 − $ 956,123,749 − $ 17,541,601 − $ 9,697,031 ) = $ 38,788,123
AKA 90
Annual and Monthly Distributed Cash Flow Calculation (Year 3):
Annual Distributed Cash Flow = (Revenue − Operating Cost − Taxes)
Annual Distributed Cash Flow = ($ 1,022,150,504 − $ 956,123,749 − $ 9,697,031 )
Annual Distributed Cash Flow = $ 56,329,724
Monthly Distributed Cash Flow =Annual Distributed Cash Flow
12
Monthly Distributed Cash Flow =$ 56,329,724
12= $ 4,694,144
Present Worth Discrete Cash Flow (P) calculation in year 3:
P =A [
(1 + i)12−1 + Li(1 + i)12 ]
(1 + i)12×(n−1)=
$ 4,694,144 [(1 + 0.005)12−1 + (−26,312,401)
i(1 + i)12 ]
(1 + i)12×(n−1)
P =$ 4,694,144 [
(1 + 0.005)12−1 + (−26,312,401)i(1 + 0.005)12 ]
(1 + i)12×(3−1)= $ 25,043,952
Where A is the monthly distributed cash flow, i is the Annual or monthly interest
rate, and n is the annual or monthly period. L is Lump sum revenue; L will be zero
for all years except on year 3, which is the first year of operation (Turton, 2012)
Future Worth Discrete Cash Flow (F) calculation in year 3:
F = P[(1 + i)n] = $ 25,043,952 × (1 + 0.005)144 = $ 51,358,905
AKA 91
Appendix E: Literature Review
Figure 1E: Propylene Downstream Uses
The figure above shows the downstream uses of Propylene. It can be seen that
propylene is a raw material to many other intermediate products that go into the
production of daily used products. In addition, it can be concluded that propylene is
a material that has high importance.
AKA 92
Figure 2E: Global Propylene Consumption (IHS, 2015)
The chart above presents the world consumption of propylene. It can be
noticed that the propylene is roughly 64% used in the production of poly propylene,
7% in the production of propylene oxide; and 6% in the production of acrylonitrile.
These three derivatives of propylene are the main markets for propylene.
AKA 93
Figure 3E: Propylene and Ethylene Price Trends in the Middle East
The trend above shows the price of propylene in Middle east in the last 16
years. It can be noticed that the price is increasing with time. Based on the figure
above, it can be considered that the site location in Saudi Arabia is reasonable due
to the demand of propylene.
AKA 94
Figure 4E: Process Routes to Producing Propylene (Jasper, 2015)
The figure above presents the routes of producing propylene using different
raw materials. There are several alternative routes to produce methanol in the
industry. There are three main sources that produce propylene such as Crude oil,
Natural gas, and coal Heavy oil. In this case, Natural gas was selected to be the
source of producing propylene. The plant is located in Saudi Arabia, Jubail industrial
city. There are several reasons behind choosing this location such as that Saudi
Arabia is the 6th Largest in Natural Gas Reserves and also the 9th Largest Producer of
Natural Gas. Another advantage is that the cooperate tax is relatively low.
MTP Reactor Reaction Study
Based on the literature, the reactions were found to be 19 reactions (Wen,
2016), yet the kinetic data was not available. A stoichiometric study was performed
to theoretically model the reactor using the selectivity data from the novelty
(Moreno-Pirajan, 2013). From the stoichiometric study, the amounts for the alkanes
AKA 95
Figure (5E): Fire triangle
and aromatics were found to be negligible relative to the other alkenes produced.
From this conclusion, an assumption was made to neglect those reactions to simplify
the reactor simulation.
Relief system
There are several safety Precautions that need to be considered in any chemical
plant. Equipment failure or operation error can cause increase in the process
pressure beyond the safe level. In the case that pressure increases beyond the safe
level in a distillation column, tank, reactors and pipelines, it could result in rupture
in the units which lead to the release of toxic or flammable chemicals. Designing a
relief system is a significant procedure to insure plant safety. There are several steps
to install a relief system around the plant.
1- Install safety valves in the relieving locations
2- Choosing the relief type
3- Developing relief scenarios
4- Determining the worst case scenario and sizing the valves
5- Design the relief system
Table 1E: Ignition Sources of Major Fires (Louvar)
Source Present of Accident
Electrical 23
Smoking 18
Friction 10
Overheated Materials 8
Hot Surfaces 7
Burner Flames 7
Others 27
AKA 96
Fire and explosion
Fires and explosions occur when the triangle of fire is completed as shown in
figure (1). Both fire and explosion can be prevented by removing any leg from the
fire triangle. In the design, the fuel is mainly Propylene, methanol and gasoline, the
oxidizer is oxygen and the ignition sources could be sparks, flames, static electricity
and heat from hot surface. The ignition sources of major fires are shown in table (1).
It can be observed that the major sources of ignitions are electrical, smoking and
others. These sources of ignition can easily be controlled by adopting stringent
safety rules and following training guideline.
Plant Environment
The geographical location of the final plant can have a strong influence on the
success of an industrial venture because it is located in Jubail industrial area in Saudi
Arabia. An ideal location is where the cost of the product is kept to minimum, with a
large market share, the least risk and the maximum social gain. There is only one
waste in the process which is waste water. The waste is literally pure(99.97% mole
fraction) but it contains methanol which is flammable and toxic. It was assumed that
the waste water will be send to a neighbor treatment facility for further processes
due to the methanol contamination. The high pressure steam and the low pressure
steam outlet temperatures are kept to 115 oC. Economical wise, the product of the
heating units is manipulated to be boiler feed water to reduce the cost.
AKA 97
Identification of Hazards
Physical Hazards
Vibration and noise are examples of physical hazards. As a factor within the
environment that can harm the body even without necessarily physical touching. A
physical hazard arises when use of a chemical is potentially dangerous. For
example, to the possibility of explosion, fire or violent reaction.
Health Hazards
In today ‘s environment there are a number of potential health hazards that
you need to be aware of, and control properly, to help reduce the risk to your health
and the health of people around you. For example, the air we breathe can contain
emissions from motor vehicles, industry, heating and commercial sources, as well as
household fuels. Air pollution can be harmful to human health, particularly in those
people who are already vulnerable because of their age or existing health
problems.
Permissible Exposure Limits
We are exposed to all kinds of goods and materials daily. Different
substances involve different risks. The risk of fire or explosion may be present at the
same time as the danger of being exposed to poisoning or suffocation. The
permissible exposure limit (PEL) is the time-weighted average threshold limit a
person working an 8 hour shift can be exposed to a chemical without suffering any ill
effects (51-9011 Chemical Equipment Operators and Tenders, 2015).
AKA 98
Safe Handling
Handling and storage of Propylene and all the side products is an issue that
must be not to be forgotten or not to deal with it in a proper way. However, in
handling propylene and all the side products, it is recommended to keep it away
from fire, sparks and heated surfaces. Also no smoking near areas where material is
stored or handled. The product should only be stored and handled in areas with
intrinsically safe electrical classification. Based on literature, the only emission on
the process is the catalyst regeneration gas, which basically consists of nitrogen-
diluted air with a somewhat elevated CO2 content. It catalyst regeneration gases are
vented to the atmosphere because the amount is not significant.
AKA 99
Table 2E: Hazards and Safety Practices of Chemicals
Chemical Physical Hazards Health Hazards Safe Handling Controlling
Propylene
At room temperature
and atmospheric
pressure, it is a
colorless
Flammable gas
relatively nontoxic gas
Propylene is
nontoxic
Contact with the
liquid phase or
with the cold gas
escaping from
cylinder may
cause frostbite
Cylinders should
be stored and
used in dry, well
- ventilated areas
away from
sources of heat
or ignition.
Do not store with
oxidizers
In the case of
leakage, shut off
all ignition
sources and
ventilate the
area
Gasoline
Extremely flammable
gas
Contact may cause
eye, skin and
mucous
membrane
irritation
Harmful if
absorbed through
the skin
Inhalation may
cause irritation
Keep away from
flame, sparks,
excessive
temperatures
and open flame
Use approved
vented
containers
Keep containers
closed and
clearly labeled
In the case of
inhalation,
remove person
to fresh air. If
person is not
breathing,
ensure an open
airway and
provide
artificial
respiration
AKA 100
Table 3E: Hazards and Safety Practices of Chemicals
Chemical Physical Hazards Health Hazards Safe Handling Controlling
LPG
Extremely flammable
gas
Contains gas under
pressure
may explode if heated
Exposure could
cause irritation but
only minor
residual injury
even if no
treatment is given.
Keep away from
heat, sparks,
open flames or
hot surfaces
Store in a well-
ventilated place
where
temperature does
not exceed 125 oF
Leaking gas fire:
Do not
extinguish,
unless leak can
be stopped
safely
In the case of
fire, Evacuate all
personnel from
the danger area
Ethylene
Extremely flammable
gas
May form explosive
mixtures with air
Could explode if
heated
Central nervous
system depression,
difficulty breathing
Store and handle
in accordance
with all current
regulations and
standards.
Protect from
physical damage.
Store in a cool,
dry place.
EYE CONTACT:
Contact with
liquid:
Immediately
flush eyes with
plenty of water
for at least 15
minutes.
INGESTION: If a
large amount is
swallowed, get
medical
attention.
AKA 101
Table 4E: Hazards and Safety Practices of Chemicals
Chemical Physical Hazards Health Hazards Safe Handling Controlling
Methanol
Very Flammable
Stable in normal
Conditions
Explode at normal
Temperature
Hazardous in
case of skin
contact
(irritant), of eye
contact
(irritant), of
ingestion, of
inhalation
Slightly
hazardous in
case of skin
contact
(permeator)
Severe over-
exposure can
result in death.
Keep locked
up
Keep away
from heat
Keep away
from sources
of ignition
Ground all
equipment
containing
material
Do not ingest
Do not
breathe
gas/fumes/
vapor/spray
Store in a
segregated
and approved
area
Provide exhaust
ventilation or other
engineering
controls to keep the
airborne
concentrations of
vapors below their
respective
threshold limit
value
AKA 102
Risk Assessment
Risk assessment, in this context, is a tool used in risk management to help
understand risks and inform the selection and prioritization of prevention and
control strategies. With risk assessment, risks can be ranked on a relative scale and
technical/organizational/policy options can be evaluated, so that results can be
maximized in terms of increased safety. This helps in the choice of options. Risk
assessment also provides information to policymakers to help them develop risk
acceptability or tolerability criteria against which different objectives or
programmers can be assessed.
The following table shows the risk assessment of chemical plants.
AKA 103
Table (5E): Risk Assessment of Chemical Plant (Louvar)
Risk Assessment
What is
the hazard
Who could
be
harmed
Existing
Procedures
Needed
actions
How the assessment could
be transferred to an action
whom when
Chemicals
Hazards
Staff who
works in
the lab.
Getting
skin
problems
or
irritation
to eyes
All staff
wears PPE.
Special
chemicals
put in
shelves and
stored
properly.
Staff are
trained in
the risks.
Remind staff to
report any
health
problem.
Remind staff to
clean gloves
and wear PPE
Supervisor Every day
Electrical
Hazards
Electrical
operators.
Electrical
shocks.
Faulty
electrical
equipment
Insulating
electrical
wires. Staff
trained in
electrical
safety
Remind staff to
check any
electrical
equipment
before using
it.
Supervisor
During
installation
preventive
maintenance
Valves
handling
Operators.
Valves
may leak
and
release
chemicals.
Operators
wear PPE.
Valves are
coated with
insulated
materials
Remind staff to
wear PPE.
Check valve
before
handling it.
Safety
manager
During
operation
AKA 104
Table (6E): Environmental impact assessment of methanol to propylene plant
Components
Type
Composition
(Mole Fraction)
Reservoir
KSA Regulations
(Industry, n.d) Actions Needed
Product Side
Product
Storage
Tank
Treatment
Unit
Propylene Polymer grade of
Propylene 99.6%
Royal Decree
No. 38/ Dated
16.06.1427 -
12/7/2006
Article # 4
Article # 6
Article # 9
All Instructions
are presented in
the MSDS
Gasoline
Pentene 21.9 %
Hexane 18.9 %
Heptene 22.9 %
Octene 35.7 %
LPG
Butene 91.2 %
Ethylene
Ethylene 99.9 %
Methanol Methanol
Royal Decree
No. 38/ Dated
16.06.1427 -
12/7/2006
Waste Water Water 99.9 %
Royal Decree
No. 38/ Dated
16.06.1427 -
12/7/2006
AKA 105
Appendix F: Project Timeline
Table 1F: Project Tasks Performed During Spring 2016
Week # 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15
Proposal Review
Literature Review
Simulation
Optimization
Capital Cost
Operating Cost
Poster Design
Profitability Analysis
Sensitivity Analysis
Final Presentation
Final Report