production of propylene from methanolmy.fit.edu/~akurdi2012/plant design 2/design project/final...

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

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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

[email protected].

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

_________________________

mailto:[email protected]



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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

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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.

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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.

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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).

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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).

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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)

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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.

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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).

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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.

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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).

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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.

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Figure 3: Process Flow Diagram, PFD

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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



Pressure (bar) 2.70 2.50 1.60 6.10


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


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

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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


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



Pressure (bar) 5.49 5.29 11.65 11.45


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


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

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Pressure (bar) 25.20 25.00 25.32 25.00


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


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


Temperature (oC) -21 75 59.39 117.71

Pressure (bar) 25.00 25.24 25.00 25.34


Flowrate (kg/hr) 6,596 90,728 56,865 33,863

Flowrate (kmol/hr) 235.00 1,977.82 1,351.00 626.82


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

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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


Power = 17.1 kW Power = 48.0 kW

74% efficient 82% efficient

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Table 3-2: Equipment Summary (continued)

Towers

T-101 T-102


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

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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


Vertical Horizontal

L/D = 3.0 L/D = 3.0

V = 1956 m3 V = 63 m3

V-103 V-104


Horizontal Horizontal

L/D = 3.0 L/D = 3.0

V = 27 m3 V = 79 m3

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Compressors

C-101 C-102


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

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Heat Exchangers (continued)

E-103 E-104

A = 394 m2 A = 2,767 m2


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


Process stream in pipes Process stream in pipes

Q = 70,956 MJ/hr Q = 52,871 MJ/hr

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Heat Exchangers (continued)

E-109 E-110

A = 1,329 m2 Area = 1,634 m2


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

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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

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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.

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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.

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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.

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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.

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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

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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

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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))

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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.

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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

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http://www.eia.gov/forecasts/aeo/section_prices.cfm#natgas

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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

http://www.mindat.org/min-2779.html

http://www.vfpfire.com/systems-deluge.php

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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

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Pump (P-103)



the equipment





Pump (P-104)



the equipment





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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).

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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


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.



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o Aspen Plus calculates the area and heat duty for this heat exchanger: Area = 133

m2 and Q = 106,285 MJ/hr


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




m2 and Q = 95,553 MJ/hr



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 Cooling water flowrate was determined using sensitivity analysis by varying the

flowrate of cooling water (input) and monitoring the temperature of the output

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stream from the shell side such that it comes out at 45 oC (Turton,2012). The

cooling water flowrate = 1.1*107 kg/hr


2,767 m2 and Q = 742,877 MJ/hr


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 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 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


m2 and Q = 45,728 MJ/hr

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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.






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




o It is desired to condense the feed in this heat exchanger to 75 oC (Lurgi, 2003)



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cooling water flowrate = 1.05*106 kg/hr


m2 and Q = 70,956 MJ/hr


o Condenser for distillation column (T-101)



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.






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m2 and Q = 52,871 MJ/hr


o Reboiler for distillation column (T-101)


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 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


1,329 m2 and Q = 25,176 MJ/hr

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o 50/50 Water-Ethylene glycol refrigerant (2 bar, -30 oC) in the shell side of the

heat exchanger (Engineering Toolbox)



(T-102), and the temperature was -20.5 oC.


(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 Refrigerant flowrate was determined using sensitivity analysis by varying the


stream from the tube side to find the minimum amount of refrigerant. The

refrigerant flowrate = 684,649 kg/hr


1,635 m2 and Q = 19,238 MJ/hr

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(Turton,2012)







(Turton,2012)





temperature is achieved with the minimum amount of steam. The low pressure

steam flowrate = 793 kg/hr



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(T-101), and the temperature was 59.4 oC.


(Turton,2012)







m2 and Q = 51,621 MJ/hr





(Turton,2012)

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(Turton,2012)

o Heat transfer coefficient “U” was assumed 150 Btu/hr-ft2-oF (Turton, 2012)




temperature is achieved with the minimum amount of steam. The low pressure

steam flowrate = 450 kg/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

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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)

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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.

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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

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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

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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.

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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

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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.

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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

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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:


= 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:


= 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



= 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)


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



= 2.2897 + 1.3604 × log10(2425) + (−0.103) × (log10(2425))2 = 5.72

Cpo = $ 522008.6



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



= 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



= 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.



= 4.8306 + (−0.8509) × log10(12.72) + 0.3187 × (log10(12.72))2 = 4.3

Cpo = $ 19035.1



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.



= 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



= 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



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



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

production of propylene from methanolmy.fit.edu/~akurdi2012/plant design 2/design project/final...

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