my.fit.edumy.fit.edu/~akurdi2012/Plant Design 2/Design Project/Final Repor… · Web viewmy.fit.edu

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

AlmakhmariApril 27, 2016

Dr. Jonathan WhitlowFlorida Institute of TechnologyDepartment of Chemical Engineering150 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

_________________________

AKA 1

mailto:[email protected]



Nasser Almakhmari

_________________________

Table of ContentLetter of Transmittal DONE 1Table of Contents 2Executive Summary DONE 3Introduction DONE 4Process Description DONE

Process Flow Diagram, PFD DONEStream Table DONEUtilities Table DONEEquipment Tables DONE

Process Design and Simulation DONECapital Cost DONEManufacturing Cost DONEProfitability Analysis DONESensitivity Analysis Needs to be UpdatedProcess Control DONE

Process Instrumentation Diagram, PID DONEEnvironmental and Safety Consideration DONEReferences Needs to be alphabetizedAppendix A: Equipment Design Methods, Calculations and Assumptions DONEAppendix B: Capital Cost Sample Calculations DONEAppendix C: Manufacturing Cost Sample Calculations DONEAppendix D: Profitability Analysis Sample Calculations DONEAppendix E: Literature Review Needs work + in body referencing Appendix F: Project Timeline DONE

AKA 2

Executive Summary

AKA 3

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)(See Appendix E

for more information about the different routes of producing propylene)

(Jasper, 2015).

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

AKA 4

production rate of 480,000 metric tons of propylene annually. The design

and simulation of the plant was conducted using Aspen Plus V8.8.

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??. The

kingdom has recently started looking into diversifying its sources of income

and wants to satisfy local petrochemical demand??. 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 D for more information

AKA 5

about mordenite zeolite and the experimental results as well as the uses of

the by-products).

Polypropylene, propylene oxide, and acrylonitrile are the most

common chemical derivatives from propylene?. 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?. Propylene oxide goes into the production of propylene glycol which

is used as antifreeze for cars, deicing of aircrafts, and goes into making

cosmetics?. Acrylonitrile goes into the production of acrylic fibers, which are

used in clothing and goes into the production of paints and adhesives? (See

Appendix E for more information about the products of propylene).

AKA 6

The demand for propylene increases annually and continues to be

driven primarily by developments in the polypropylene industry followed by

the propylene oxide and the acrylonitrile industries (IHS, 2015). Figure 1

shows the increasing world demand and estimates that the demand for the

year 2020 will reach 100 million tons (Galadima, 2015).

Figure 1: Historical and Expected Propylene Worldwide demand (Galadima,

2015)

AKA 7

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 E. 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

AKA 8

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.

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

AKA 9

reactor (R-101), methanol is converted into dimethyl-ether and water via the

following equilibrium reaction (Farsi, 2010):

2CH3 OH⟺CH 3O CH3+H 2 O

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). 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):

nCH 3O CH3 →2Cn H 2n+n H 2O n=2 ,…,8

nCH 3OH →Cn H2 n+n H2O 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;

AKA 10

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. 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

AKA 11

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). The third distillation column

has 48 stages with the feed entering on the 22th stage. In this column,

propylene is separated from the hydrocarbon mixture as the distillate

product and leaves the column at a purity of 99.6%. The remaining

hydrocarbon mixture, mainly butene in this particular process, leaves the

process as liquid petroleum gas, LPG, as the bottoms product at a purity of

91.2%. In general, all product streams (19, 21, 23, and 24) are sent to their

designated storage tanks to be sold.

AKA 12

AKA 13

Figure 3: Process Flow Diagram, PFD

AKA 14

Table 1-1: Stream Summary

Stream Number 1 2 3 4Temperature (oC) 45 45 97 266Pressure (bar) 1.00 3.35 3.25 3.05Vapor Mole Fraction 0 0 1 1Flowrate (kg/hr) 350,000 350,000 350,000 350,000Flowrate (kmol/hr) 10,923.

1110,923.

1110,923.

1110,923.

11Component Flowrates (kmol/hr)

Methanol 10,923.11

10,923.11

10,923.11

10,923.11

Water 0 0 0 0Dimethyl-ether 0 0 0 0Ethylene 0 0 0 0Propylene 0 0 0 0Butene 0 0 0 0Pentene 0 0 0 0Hexene 0 0 0 0Heptene 0 0 0 0Octene 0 0 0 0


1110,923.

1113,689.

4513,689.


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 0 235.23 235.23Propylene 0 0 1,353.2

61,353.2

6

AKA 15

Butene 0 0 573.46 573.46Pentene 0 0 136.85 136.85Hexene 0 0 114.04 114.04Heptene 0 0 138.20 138.20Octene 0 0 215.30 215.30


Stream Number 9 10 11 12Temperature (oC) 506 383 38 38Pressure (bar) 5.90 5.70 5.50 5.49Vapor Mole Fraction 1 0.61 0.21 1Flowrate (kg/hr) 350,000 350,000 350,000 154,069Flowrate (kmol/hr) 13,689.

4513,689.

4513,689.

452,815.8


Methanol 2.97 2.97 2.97 0.17Water 10,901.

2610,901.

2610,901.

26 30.46Dimethyl-ether 18.88 18.88 18.88 18.87Ethylene 235.23 235.23 235.23 235.22Propylene 1,353.2

61,353.2

61,353.2

61,353.2

4Butene 573.46 573.46 573.46 573.46Pentene 136.85 136.85 136.85 136.85Hexene 114.04 114.04 114.04 114.04Heptene 138.20 138.20 138.20 138.20Octene 215.30 215.30 215.30 215.30


632,815.8

22,815.8

22,815.8


Methanol 2.80 0.17 0.17 0.17Water 10,870. 30.46 30.46 30.46

AKA 16

80Dimethyl-ether 4.39E-

03 18.87 18.87 18.87Ethylene 0.01 235.22 235.22 235.22Propylene 0.02 1,353.2

41,353.2

41,353.2

4Butene 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


Stream Number 17 18 19 20Temperature (oC) 168 75 236 58Pressure (bar) 25.20 25.00 25.32 25.00Vapor Mole Fraction 1 0 0 0Flowrate (kg/hr) 154,069 154,069 56,745 97,324Flowrate (kmol/hr) 2,815.8

22,815.8

2 603.00 2,212.82

Component Flowrates (kmol/hr)

Methanol 0.17 0.17 0.17 0.01Water 30.46 30.46 0.98 29.48Dimethyl-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 1.84 571.62Pentene 136.85 136.85 132.48 4.37Hexene 114.04 114.04 114.04 4.30E-

07Heptene 138.20 138.20 138.20 4.68E-

14Octene 215.30 215.30 215.30 9.43E-

21

AKA 17

Stream Number 21 22 23 24Temperature (oC) -21 75 59.39 117.71Pressure (bar) 25.00 25.24 25.00 25.34Vapor Mole Fraction 0 0 0 0Flowrate (kg/hr) 6,596 90,728 56,865 33,863Flowrate (kmol/hr) 235.00 1,977.8

21,351.0

0 626.82Component Flowrates (kmol/hr)

Methanol 8.28E-16 0.01 6.83E-

13 0.01

Water 1.20E-08 29.48 7.86E-

05 29.48

Dimethyl-ether 3.97E-05 18.87 4.85 14.03

Ethylene 234.77 0.46 0.46 4.10E-13

Propylene 0.23 1,353.00

1,345.60 7.40

Butene 4.06E-06 571.62 0.10 571.52

Pentene 2.17E-12 4.37 4.90E-

10 4.37

Hexene 1.79E-23

4.30E-07

3.95E-23

4.30E-07

Heptene 0 0 0 0Octene 0 0 0 0

Table 2: Utility Summary

Equipment Number E-101 E-104 E-105 E-106Utility Type LPS CW HPS CWAmount of Utility (ton/hr) 188.7 10,953 18.6 18.1Equipment Number E-107 E-108 E-109 E-111Utility Type CW CW HPS LPSAmount of Utility (ton/hr) 1,046 792.7 13.1 0.793Equipment Number E-112 E-113 R-101 R-102Utility Type CW LPS CW CWAmount of Utility (ton/hr) 761 0.45 3,645 13,498Equipment Number C-101 C-102 C-103 P-101Utility Type Electricity Electricity Electricity Electricity

AKA 18

Amount of Utility (MW) 39.3 2.4 2.4 0.04Equipment Number P-102 P-103 P-104Utility Type Electricity Electricity ElectricityAmount 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 driveCarbon steel Carbon steelPower = 36.4 kW Power = 37.6 kW82% efficient 81% efficientP-103 A/B P-104 A/B

Centrifugal / electric drive Centrifugal / electric driveCarbon steel Carbon steelPower = 17.1 kW Power = 48.0 kW74% efficient 82% efficient

Table 3-2: Equipment Summary (continued)

TowersT-101 T-102Carbon steel Carbon steel44 CS sieve trays plus reboiler and condenser

33 CS sieve trays plus reboiler and condenser

60% efficient trays 70% efficient traysTotal condenser (E-108) Total condenser (E-110)Feed on tray 28 Feed on tray 11

AKA 19

Reflux ratio = 0.81 Reflux ratio = 8.972 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 steel48 CS sieve trays plus reboiler and condenser70% efficient traysTotal condenser (E-112)Feed on tray 22Reflux ratio = 2.622 ft tray spacing Column height = 35.1 m Diameter = 3.3 m


ReactorsR-101 R-102Carbon steel, Shell & Tube, Packed tubes, Aluminum Oxide Catalyst

Carbon steel, Process vessel, Fixed bed, HMOR Zeolite Catalyst

V = 102 m3 V = 164 m3Length = 8 m, Tube Diameter = 0.09 m L/D = 3.02000 Tubes 100% filled with active catalystTubes are 100% filled with active Q = -281,784 MJ/hr

AKA 20

catalystQ = - 75785 MJ/hrVesselsV-101 V-102Carbon steel Carbon steelVertical Horizontal L/D = 3.0 L/D = 3.0V = 1956 m3 V = 63 m3V-103 V-104Carbon steel Carbon steelHorizontal Horizontal L/D = 3.0 L/D = 3.0V = 27 m3 V = 79 m3


Compressors

C-101 C-102

Carbon steel Carbon steelReciprocating ReciprocatingPower = 39.3 MW Power = 2.4 MW

AKA 21

85% adiabatic efficiency 75% adiabatic efficiencyC-103

Carbon steelReciprocatingPower = 2.4 MW75% adiabatic efficiencyHeat Exchangers

E-101 E-102

A = 1,650 m2 A = 133 m2

Floating head, carbon steel Floating head, carbon steelProcess stream in tubes Process stream in tubes & shellQ = 447,992 MJ/hr Q = 106,285 MJ/hr


Heat Exchangers (continued)

E-103 E-104

A = 394 m2 A = 2,767 m2

Floating head, carbon steel Floating head, carbon steelProcess stream in tubes & shell Process stream in tubes

AKA 22

Q = 95,553 MJ/hr Q = 742,877 MJ/hrE-105 E-106

A = 136 m2 A = 4.2 m2

Floating head, carbon steel Double pipe, carbon steelProcess stream in tubes Process stream in pipesQ = 45,728 MJ/hr Q = 1,231 MJ/hrE-107 E-108

A = 298 m2 A = 572 m2

Floating head, carbon steel Floating head, carbon steelProcess stream in pipes Process stream in pipesQ = 70,956 MJ/hr Q = 52,871 MJ/hr


Heat Exchangers (continued)

E-109 E-110

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

Floating head, carbon steel Floating head, carbon steelProcess stream in tubes Process Stream in Tubes

AKA 23

Q = 25,176 MJ/hr Q = 19,238 MJ/hrE-111 E-112

Area = 25 m2 Area = 600 m2

Floating Head, carbon steel Floating head, carbon steelProcess Stream in Tubes Process Stream in TubesQ = 2,024 MJ/hr Q = 51,621 MJ/hrE-113

Area = 21 m2

Floating head, carbon steelProcess Stream in PipesQ = 1,056 MJ/hr


Storage Tanks

V-105 V-106

Methanol Storage Fuel gas StorageVolume: 32784 m3 Volume: 1374 m3

Capacity for 3 days Capacity for 3 days

AKA 24

Vertical tank on concrete pad Vertical tank on concrete padV-107 V-108

Propylene Storage Gasoline StorageVolume: 9331 m3 Volume: 10074 m3

Capacity for 3 days Capacity for 3 daysVertical tank on concrete pad Vertical tank on concrete padV-109

LPG StorageVolume: 5555 m3

Capacity for 3 daysVertical tank on concrete pad

Process Design and Simulation (references + in text citation)

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.

AKA 25

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(1N) (See Appendix D for more information about

the process of selection of the property method).

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.

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

AKA 26

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.

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

AKA 27

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.

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

AKA 28

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 29

Figure 4: Aspen Plus view of Plant Design Simulation of the Process

AKA 30

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(1); 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 1.

Table #: Assumptions Made in Calculating the Cost of certain Equipment

Unit AssumptionsOver 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 31

Table () (Cont.): 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.

Vessels

All of Towers were modeled as Process Vessels (vertical).

The assumptions were made to findFP, Vessel, FM, B1, and B2.

I. Capital Methodology

Purchased Equipment Cost

The following equation was used for calculating the purchased cost of

the equipment, assuming ambient operating pressure (Turton, 2012):

log10 Cpo=K1+ K2 log10 ( A )+K3 [ log10 ( A ) ]2

C po: Purchased cost

A: Capacity or size parameter for the equipmentK1, 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

AKA 32

were more than one piece of the equipment. The final cost was multiplied by

the number of equipment.

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:

log10 Fp=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 33

Bare Module Factor for Heat Exchangers, Process Vessels, Pumps and

Compressor:

CBM=C po FBM=C p

o (B1+B2 FM Fp)

CPo: Purchased CostFBM: Bare module factorB1 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=C po N FBM Fq

CPo: Purchased CostN: number of traysFBM: Bare module factorFq: Quantity factor for trays The quantity factor for trays, Fq, for N ≥ 20: Fq = 1

Using this equation if N ≤ 20: log10 Fq=0.4771+0.08516∗log10 (N )−0.3473 [ log10 ( N ) ]2

CEPCI Costing Correction to 2016

Ca=Cb(Aa

Ab)

AKA 34

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.

II. Capital Results

Table (): Cost of Each Piece of Equipment

Equipment Unit # Cost (2016 $)

PumpP-101 A/B $57,000P-102 A/B $58,100P-103 A/B $39,600P-104 A/B $66,000

Vessel V-101 $15,300,000V-102 $480,000V-103 $210,000V-104 $600,000

Heat Exchangers

E-101 $755,000E-102 $95,000E-103 $193,000E-104 $1,300,000E-105 $96,000E-106 $10,000E-107 $156,000E-108 $266,000E-109 $610,000E-110 $750,000E-111 $60,000E-112 $278,000E-113 $60,000

CompressorsC-101 A/B $29,320,000C-102 A/B $3,550,000C-103 A/B $3,530,000

TowersT-101 $9,470,000T-102 $900,000T-103 $3,480,000

Sieve Trays T-101 $680,000 AKA 35

T-102 $83,000T-103 $270,000

Reactors R-101 $63,000R-102 $653,000

Table ()(Cont): Cost of Each Piece of Equipment

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 () shows the total costs of the equipment used in the

plant.

AKA 36

220000; 0.2% 16600000; 14.4%

4600000; 4.0%

36400000; 31.6%15000000; 13.0%

700000; 0.6%

41800000; 36.2%

Fixed Capital Cost

Pumps Vessels Heat Exchangers CompressorsColumns\Towers Reactors Tanks

Figure (): Total Cost of Equipment where are tanks?

The total capital cost was found to be $ 175,400,000. From figure 3, 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.

Manufacturing Cost

AKA 37

After calculating the capital investment needed to build the production

of Propylene from Methanol plant, the operational investment is 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 costs 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 stated in Turton

(Turton, 2012).

There are values that need to be found first to calculate the COM. The

First value is FCI (fixed capital cost), COL (Cost of operating labor), CUT (Cost of

Utility), CWT (Cost of wastewater), and CRM (Cost of raw material). The

Following equation was used to calculate COM without Depreciation:

COM=(0.18× FCI )+(2.73 ×COL)+(1.23 ×(CUT+CWT+CRM))

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 52500$ in 2014 (51-9011

AKA 38

Chemical Equipment Operators and Tenders, 2015). The Annual labor wages

in 2016 is $53550 by assuming 2 percent increasing from 2014. To calculate

the total number of operators is by using this equation

NOL×¿ of operator hired for eachoperator .

NOL=√6.29+31.7 P2+0.23N np

The previous 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; Pump and Vessel 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 will be $963,900

C OL=The Annual labor wages ×Total number of operators

Cost of waste treatment

Wastewater: the only waste of the process is water. Using the mass

flow rate. The price found to be equal to 0.041 $/MT (Turton, 2012). The cost

of waste treatment is equal to $67,479 per year.

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 cost of utilities are dependent on both inflation and energy

AKA 39

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 the heat exchangers, reactors, compressors, and pumps.

Table 4 shows the total amount of each utility needed in the plant annually,

and the price and annual cost of each utility.

Table 4: 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 $26,400,000Cooling water ($/kg) 0.0000175 258,002,773,140 $4,500,000High Pressure Steam

($/kg) 0.01459 266,292,516 $3,900,000Low Pressure Steam

($/kg) 0.01348 1,595,698,056 $21,500,000Refrigerant ($/GJ) 11.2671 16,175 $182,000

The total utility cost, from Table 4, is $58,100,000. It can be noticed

from Table 4 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 Methanol to Propylene plant are

presented in Appendix C. For cooling water, high pressure steam, low-

pressure steam, and refrigerant. Due to the dependence of cooling water on

AKA 40

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). The data presented in

Appendix C shows the linear extrapolation that was used the electricity cost

and for the natural gas.

Cost of raw Material

The raw materials in the methanol to propylene production are

Methanol, and catalyst. 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), 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). 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).

II. Manufacturing Results

AKA 41

The sample calculations for the cost of Manufacturing can be found in

Appendix C. The following Table () shows a summary of the costs included in

the manufacturing of the plant.

Table (): Summary of the Costs included in the Manufacturing Cost

Direct Manufacturing Costs $791,600,000Direct Supervisory and Clerical

Labor $173,000Maintenance and Repairs $10,500,000Fixed Manufacturing Cost $12,600,000Local taxes and Insurance $5,600,000

Plant Overhead costs $6,900,000Raw Materials $691,300,000

Utilities $56,400,000Operating Labor $963,000Waste treatment $103,000

Lab Charges $145,000Patents and Royalties $28,700,000

Fixed Capital Investment $175,400,000Cost of Manufacturing $956,100,000

General Expenses $154,700,000Administration Costs $1,800,000

Distribution and Selling Costs $105,100,000Research and Development $47,800,000

Form Table 5, the total manufacturing cost of the plant is

$953,618,411 while the fix capital investment is $172,935,343. The direct

manufacturing cost is $789,675,600; the fixed manufacturing cost is

$12,442,044 and the general expenses have a total of $154,305,974.

Table (6): Direct Costs Distribution

Maintenance and Repairs

10,500,000 1.32%

Raw Materials 691,300,0 87.74%

AKA 42

00

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 (6) 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.

Profitability Analysis

I. Profitability Methodology: Profitability of the methanol to propylene 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; details are shown in Appendix D. The land cost

was assumed to be equal to 5 million dollars. In addition, the annual interest

AKA 43

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(Turton, 2012). The construction

period was assumed to be two years, with an expected plant lifetime of ten

years (Turton, 2012).

Table 7: Annual revenue

ProductsAmount

Annually (MT/yr) Price ($/MT) TotalPropylene 477668.52 1250 $597,085,650Ethylene 55405.812 1000 $55,405,812

LPG 284446.008 45 $12,800,070Boiler Feed

Water 1741230.54 2.45 $4,266,015 (L/yr) Cost ($/L)

Gasoline 1175309856 0.3 $352,592,957

The most up to date prices for all products are shown in table (7).

Propylene price is 1250 dollar per MT (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 $/MT (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

AKA 44

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)

II. Profitability Results:

0 1 2 3 4 5 6 7 8 9 10 11 12-$200,000,000

$0

$200,000,000

$400,000,000

$600,000,000

$800,000,000

$1,000,000,000

$1,200,000,000

$1,400,000,000

Cash Flow

Time (year)

Cash

Flo

w, M

illio

n US

Dol

lars

Figure 2: Cumulative Future vs. Time

Based on Figure 2 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.

AKA 45

Sensitivity Analysis

AKA 46

Process Control

The process instrumentation diagram, PID, is shown in Figure #. 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 #, 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

AKA 47

liquid outlet. On 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 #

presents the detailed controlling on one of the distillation columns, which is

also applies to the rest. Finally, Figure # shows the water jacket design on

the MTP reactor with necessary controlling (Refer to appendix D for the

discussion on Relief Systems).

AKA 48

Figure #: Process Instrumentation Diagram, PID

AKA 49

AKA 50

Figure #: Distillation Column Process Instrumentation Diagram

Figure #: MTP Reactor (R-102) with Water Jacket Process Instrumentation Diagram AKA 51

Environmental and Safety Considerations

Based on the analysis that presented on table 2 and 3, it could be

noticed that there are several consideration with regards to the environment

and safety impacts. The environmental and safety impacts of the chemical

plant could affect the employers as well as the people in the surrounding

area. These impacts could be resulted in different forms such as gases

vented to the atmosphere, the catalyst waste and the noise that could be

produced from the plant units. The main product of the plant and also the

side products are flammable hydrocarbons. Based on the simulation of the

process, the distillation columns are operating under high pressure to

achieve the needed purities. In the distillation columns, relief system was

designed to prevent inadvertent release of gases from the distillation

columns. In addition to that, the plant operates under high pressure

therefore, it might cause a pipe rupture which might lead to escape gas to

the environment. Pipe rupture due to high pressure can be prevented by

using proper thickness and insulate the pipes to prevent heating the surface

in the summer time. The lifetime of the mordenite zeolite catalyst is one

year. The waste catalyst is going to be disposed because it has no significant

effects on the environment and it is nonhazardous waste [N1]. It was assumed

that the catalyst will be likely disposed in the landfill. All unites in the

chemical plant are emanating 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

AKA 52

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 down sewer or drainage systems, unless system is designed and

permitted to handle such material.

As shown in table (N2), there are several safety consideration in the plant

due to the presence of high temperature, high pressure, toxicity of the

chemicals, and the 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 I the plant should follow the regulations and

use the 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 the existing site emergency plan. Emergency planning

includes tornado and storm shelters, flood protection, earthquakes, proximity

to public areas, and safe exit routes. It also includes planning for the effect

that an emergency in the "new process" would have on other plants, and the

effect that an emergency in another plant would have on the new process.

AKA 53

The effects of potential spills on waterways and aquifers should be

considered [N5].

Designing relief-venting systems is important to ensure that flammable

or toxic gases are vented to a safe location. 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 and “kill” the material; for instance, the provision of

caustic scrubbers for heavy hydrocarbons. If flammable materials have to be

vented at frequent intervals; as, for example, in some refinery operations,

flare stacks are used [N6].

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 [N7].

References Needs to be alphabetized + add Nasser and Khalid

1 Chemical process safety, 2nd ed, daniel a. crowl, joseph f. louvar.

(n.d.).

2 - http://www.iza-online.org/natural/Datasheets/Mordenite/mordenite.htm

3-

https://www.fas.org/sgp/crs/homesec/R43070.pdf

AKA 54

https://www.fas.org/sgp/crs/homesec/R43070.pdf%204

http://www.iza-online.org/natural/Datasheets/Mordenite/mordenite.htm

4 -

http://mci.gov.sa/en/LawsRegulations/SystemsAndRegulations/chemicals/

Pages/regulation.aspx

7- Deluge Suppression System. (n.d.). Retrieved April 23, 2016, from

http://www.marioff.com/fire-protection/hi-fogr-suppression-system-

types/deluge-suppression-system

Meyers, R. (2005) Handbook of Petrochemicals Production Processes. LURGI

MTP® TECHNOLOGY, AccessEngineering

Lurgi MTP Plant Process Description. (2003). CHE397Hotel

Farsi, M., Jahanmiri, A., & Eslamloueyan, R. (2010). Modeling and

Optimization of MeOH to DME in Isothermal Fixed-bed Reactor.

International Journal of Chemical Reactor Engineering, 8(1).

Galadima, A., & Muraza, O. (2015). Recent Developments on Silicoaluminates

and Silicoaluminophosphates in the Methanol-to-Propylene Reaction: A

Mini Review. Industrial & Engineering Chemistry Research Ind. Eng.

Chem. Res., 54(18), 4891-4905.

Hadi, N., Niaei, A., Nabavi, R., Farzi, A., & Navaei Shirazi, M. (2014).

Development of a New Kinetic Model for Methanol to Propylene Process

on Mn/H-ZSM-5 Catalyst. 28(1), 53-63.

Hong, S. (2008). Retrofit Design of Methanol-to-Propylene Process for the

Changes in Feedstock and Catalyst. Korea Advanced Institute of

Science and Technology.

AKA 55

http://mci.gov.sa/en/LawsRegulations/SystemsAndRegulations/chemicals/Pages/regulation.aspx

http://mci.gov.sa/en/LawsRegulations/SystemsAndRegulations/chemicals/Pages/regulation.aspx

https://www.fas.org/sgp/crs/homesec/R43070.pdf%204

IHS, Chemical Economics Handbook, Propylene. (2015, February). Retrieved

from https://www.ihs.com/products/propylene-chemical-economics-

handbook.html

Jasper, S., & El-Halwagi, M. (2015). A Techno-Economic Comparison between

Two Methanol-to-Propylene Processes. Processes, 3(3), 684-698.

Moreno-Pirajan, J. C., & Giraldo, L. (2013). Catalytic Conversion Process of

Methanol-To-Propylene (MTP) With Zeolites. Rasayan J. Chem., 6(3),

172-174.

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

Turton, Baillie, Whiting, Shaeiwitz & Bhattacharyya. (2012). Analysis,

Synthesis and Design of Chemical Processes. Check formatting

Engineering Toolbox. Ethylene Glycol Heat-Transfer Fluid. (n.d.). from

http://www.engineeringtoolbox.com/ethylene-glycol-d_146.html

Mindat.org. Mordenite, from http://www.mindat.org/min-2779.html

Works Cited51-9011 Chemical Equipment Operators and Tenders. (2015, December 2). Retrieved from U.S. Bureau

of Labor Statistics: http://www.bls.gov/oes/current/oes519011.htm

AKA 56

Aluminum Oxide price. (2016, 3 15). Retrieved from Alibab.com Global trade : www.alibaba.com/showroom/aluminium-oxide-price.html

Annual Energy Outlook 2015. (2015, April 14). Retrieved from U.S. Energy Information Administration: http://www.eia.gov/forecasts/aeo/section_prices.cfm#natgas

Argaam Petrochemical Index Loses 3.7 Pts as Polymers. (2015, December 2). Retrieved from ArgaamPlus.

Dukandar, K. N. (2016). Alternative On-purpose Production Methods for Propylene. CB&I.

Electric Power Monthly. (2015, Agust). Retrieved from U.S. Energy Information Administration: http://www.eia.gov/forecasts/aeo/section_prices.cfm#natgas

Petrol Prices Across the World. (2016, March 19). Retrieved from Kshitij Consultancy Services: http://www.kshitij.com/research/petrol.shtml

Saudi Arabia: Tax System. (2016, April). Retrieved from Santander TradePortal: https://en.santandertrade.com/establish-overseas/saudi-arabia/tax-system

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)

AKA 57

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




2012).


Pump (P-103)




AKA 58




2012).


Pump (P-104)







2012).


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

AKA 59

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).

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)

AKA 60

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 cooldown stream #9.


(Turton,2012)



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

AKA 61

o The purpose of this heat exchanger is to heat stream #5 to 420 oC (Hong,

2008) and cooldown stream #8


(Turton,2012)



Area = 394 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


(Turton,2012)


o Cooling water flowrate was determined using sensitivity analysis by

varying the flowrate of cooling water (input) and monitoring the


out at 45 oC (Turton,2012). The cooling water flowrate = 1.1*107 kg/hr

AKA 62


Area = 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


(Turton,2012)


o High pressure steam flowrate was determined using sensitivity analysis

by varying the flowrate of high pressure steam (input) and monitoring the


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


Area = 136 m2 and Q = 45,728 MJ/hr

AKA 63




(Turton,2012)

o It is desired to cooldown 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.


(Turton,2012)





out at 45 oC (Turton,2012). The cooling water flowrate = 18,145 kg/hr


Area = 4.2 m2 and Q = 1,231 MJ/hr




(Turton,2012)

AKA 64

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

2003)


(Turton,2012)





out at 45 oC (Turton,2012). The cooling water flowrate = 1.05*106 kg/hr


Area = 298 m2 and Q = 70,956 MJ/hr


o Condenser for distillation column (T-101)



(Turton,2012)

o It is desired to condense the vapor rising from the top tray in this heat

exchanger; the desired temperature was obtained from the profile of the

distillation column (T-101), and the temperature was 58.3 oC.

o Pressure drop of 0.2 bar in shell and 0.007 bar in tube side was assumed

(Turton,2012); the tube side pressure drop is equal to the pressure drop

between the trays of the distillation column, since tray 1 in aspen

represents the condenser and tray 2 is the top try of the column.

AKA 65







Area = 572 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


o It is desired to re-boil the liquid dropping from the bottom tray in this heat



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

AKA 66

desired temperature is achieved with the minimum amount of steam. The

high pressure steam flowrate = 13,133 kg/hr


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




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

the heat exchanger (Engineering Toolbox)



distillation column (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 flowrate of cooling water (input) and monitoring the temperature of

the output stream from the tube side to find the minimum amount of

refrigerant. The refrigerant flowrate = 684,649 kg/hr


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

AKA 67


















low pressure steam flowrate = 793 kg/hr


Area = 25 m2 and Q = 2,024 MJ/hr




AKA 68


(Turton,2012)





(Turton,2012)







Area = 600 m2 and Q = 51,621 MJ/hr









AKA 69





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





low pressure steam flowrate = 450 kg/hr


Area = 21 m2 and Q = 1,056 MJ/hr

Reactor

DME Reactor (R-101)

The reactor was simulated in Aspen Plus using an “REquil” block. The

reaction inputted was as follow (Farsi, 2010):

2CH3 OH⟺CH 3O CH3+H 2 O

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 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

AKA 70

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=102m3

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: nCH 3O CH3 →2Cn H2n+n H2O for n=2,…,8 (Meyers, 2005)

AKA 71

Table #: List of Reactions inputted in MTP Reactor (R-102A)

n Reaction2 2CH3 OCH 3 →2C2 H 4+2 H 2O3 3CH 3O CH3 →2 C3 H 6+3 H 2O4 4 CH3 OCH 3 →2C4 H 8+4 H 2 O5 5CH3 O CH3 →2 C5 H 10+5 H 2O6 6CH 3O CH3 →2C6 H12+6 H 2O7 7CH 3O CH3 →2C7 H14+7 H2 O8 8CH 3O CH3 →2C8 H 16+8 H 2 O

The reactions inputted into the second reactor (R-102B) are as follow:

General Form: nCH 3OH →Cn H 2n+n H 2O for n=2 ,…,8 (Hadi, 2014)

Table #: List of Reactions inputted in MTP Reactor (R-102B)

n Reaction2 2CH3 OH →C2 H4+2 H2O3 3CH 3OH →C3 H 6+3 H2O4 4 CH3 OH →C4 H8+4 H 2O5 5CH3 OH →C5 H 10+5H 2 O6 6CH 3OH →C6 H 12+6 H2O7 7CH 3OH →C7 H 14+7 H 2O8 8CH 3OH →C8 H 16+8 H 2O

The reactor is designed as a fixed bed process vessel with mordenite zeolite,

HMOR, as the catalyst. The reactor operates isothermally at 452 oC (Hong,

2008). A cooling jacket, with cooling water flowing in it, was planned to

maintain the temperature of the reactor. A pressure drop of 0.9 bar was

assumed across the MTP reactor, 0.45 bar in both reactors R-102A & R-102B,

to reach the desired output pressure of 1.6 bar in the exiting stream, stream

7. The volume of the reactor was determined using the weighted hourly

space velocity “WHSV” (Moreno-Pirajan, 2013); then using the heuristics

from Turton, the length to diameter ratio was chosen, and both values were

calculated.

AKA 72

o WHSV = 1 hr-1 = Mass flowrateof feedWeight of catalyst

=350,000 kg

hrWeight of catalyst

→Weight of catalyst=350,000 kg

o Mordenite zeolite density = 2135 kg/m3 (Mindat.org)

o Volume of reactor = Weight of catalystDensity of catalyst

= 350,000 kg2135 kg /m3 =164 m3

o L/d ratio = 3, in other words L = 3 d (Turton, 2012)

o Diameter of reactor: 164 m3=(3d ) × 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 (Heuristics). 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 (Heuristics). 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(Heuristics) Aspen Plus is then able to calculate

AKA 73

the break horsepower that can be then used in costing the compressors. The

table below shows all the compressors with their corresponding parameters.

Table (!): design compressors and parameters

Compressor

Inlet Pressure

(bar)

Discharge Pressure

(bar)Pressure

RatioBreak

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 Lmin

o Assuming Ld=3

AKA 74

195565 Lmin × 1m3

1000 L = 195.567 m3

min

Volume of the vessel = 195.567 m3

min × 5 min × 2 = 1955.65 m3

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 (2N).

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, the number of passes

can be confirmed, thus the exact diameter of the column can be obtained.

AKA 75

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 was used to determine the number of passes.

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× 32.2m = 585 m3

AKA 76

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

Number

of Stages

Feed

Stage

Reflux

Ratio

20 10 21.60

6

30 15 11.90

2

50 25 8.600

60 30 8.023

70 35 7.690

75 36 7.531

85 38 7.367

95 41 7.319

AKA 77

It could be noticed that the reflux ratio is almost stable at 7.319. Based

on the heuristics in table 13 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.

Number

of Stages

Feed

Stages

Reflux

Ratio

17 8

51.08

3

18 9

33.11

9

25 12

13.83

8

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

AKA 78

The feed stage was varied to find the minimum reboiler and condenser

heat duty. The following table shows the study for T-102.

Feed

Stage

Reboiler

Heat (MW)

Condenser

Heat (MW)

Reflu

x Ratio

10 6.2235 -5.38955

9.05

4

11* 6.17844 -5.3449

8.97

0

12 6.21465 -5.38067

9.03

8

13 6.29864 -5.46319

9.19

0

*Optimum

The optimum feed tray found to be on stage 11 because it typically

minimize 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

AKA 79

o Half full holdup time = 5 min

o Assuming Ld=3

The volumes of the reflux drums was calculated as follow:

o 6270.1 Lmin × 1m3

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

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 Ld=3

Volume, diameter, and height were calculated as follow:

Flow rate of propylene = 2159.856 Lmin

2159.856 Lmin × 1m3

1000 L 60min

1hr × 24hr1day × 3 days = 9330.6 m3

AKA 80

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 (1): 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 39.9Fuel Gas Storage Tank 1374 8.98 17.9Gasoline Storage Tank 10074 18.5 37.15

Appendix B: Capital Cost Sample Calculations

1B) Calculation for Pump P-101 A/B Bare Module Cost:

Table (): Bare Module Pump Costing Coefficients

Centrifugal

Capacity (kw)

Discharge

pressure(bar)

K1 K2 K3 B1 B2 C1 C2 C3FM

(Carbon Steel)

36.4 3.685 3.3892

0.0536

0.1538 1.89 1.3

5 0 0 0 1.6

From table 1, Log (Cpo) and Log (Fp) can be calculated using the following equations:

log10 C po=K1+ K2 log10 ( A )+K 3 [ log10 ( A ) ]2=3.3892+0.0536 × log10 (36.4 )+0.1538 ×¿¿

C po=7041.95

For the P<10, C1, C2, and C3 will be zero

log10 Fp=C1+C2 log10 ( P )+C3 [ log10 ( P ) ]2=0

AKA 81

F p=1

CBM=C po ( B1+B2 F M F p )=7041.95 × (1.89+ (1.35 × 1.6× 1 ) )=$28,520

2B) Calculation for Vessel V-101:

Table (): 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 1.8

2 1.6 4.944

From table 2, Log (Cpo) and Fp can be calculated using the following equations:

log10 C po=K1+ K2 log10 ( A )+K 3 [ log10 ( A ) ]2=3.4974+0.4485× log10 (488.92 )+0.1074 × ¿¿

C po=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=C po (B1+B2 F M F p )=302160.1 × (2.25+(1.82 ×1×1 ) )=$3,835,378

3B) Calculation for Heat Exchanger H-103:

Table (): Bare Module Heat Exchanger Costing Coefficients

Floating Head

Capacity (m3)

Q ( MJh

)K1 K2 K3 B1 B2

FM(Carbon Steel)

Pressure(barg)

394.19 106285

4.8306

-0.85

10.318

7 1.63 1.66 1

5.61

AKA 82


log10 Cpo=K1+ K2 log10 ( A )+K3 [ log10 ( A ) ]2=4.8306+(−0.851)× log10 (394.19 )+0.3187 ×¿¿

C po=58777

For the P<10, C1, C2, and C3 will be zero (Turton,2012)


F p=1

CBM=C po (B1+B2 F M F p )=58777 × ( 1.63+ (1.66 ×1× 1 ) )=$ 193,378

4B) Calculation for Compressors C-102:Table (): 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


AKA 83

log10 C po=K1+ K2 log10 ( A )+K 3 [ log10 ( A ) ]2=2.2897+1.3604 ×log10 (2425 )+(−0.103)×¿¿

C po=$522008.6



F p=1

CBM=C po ( FBM )=522008.6× 3.4=$ 1,774,829

5B) Calculation for Column T-103:Table (): Bare Module Column Costing Coefficients

Tray Vertical Towers

Diameter (M)


FM(Carbon Steel)

Pressure(barg)

3.30 300 3.4974

0.4485

0.1074

2.25

1.82 1 26.4


log10 C po=K1+ K2 log10 ( A )+K 3 [ log10 ( A ) ]2=3.4974+0.4485× log10 (300 )+0.1074 ×¿¿

C po=$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=C po ( B1+B2 F M F p )=184885.871 × (2.25+(1.82 ×1×13.1 ) )=$3,477,566

6B) Calculation for Sieve Trays T-103:Table (): Bare Module Sieve Trays Costing Coefficients

Sieve Trays

AKA 84

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 traysThe quantity factor for trays, Fq, for N ≥ 20: Fq = 1


log10 C po=K1+ K2 log10 ( A )+K 3 [ log10 ( A ) ]2=2.9949+0.4465 × log10 (8.53 )+0.3961× ¿¿

C po=$5676.5

CBM=C po N FBM Fq=5676.5 × 48×1 ×1=$ 272,471

7B) Calculation for Reactor R-101:Table (): Bare Module Reactor Costing Coefficients

Floating HeadTube

Diameter(M)

N (# of Tube )

Capacity (m2) K1 K2 K3 B1 B2 FM

Pressure

(Barg)0.09 2000 12.72 4.830

6-

0.85090.318

7 1.63 1.66 1 1.87

Note: This reactor was design as heat exchanger.


log10 Cpo=K1+ K2 log10 ( A )+K3 [ log10 ( A ) ]2=4.8306+(−0.8509)× log10 (12.72 )+0.3187 ×¿¿

C po=$19035.1



AKA 85

F p=1

CBM=C po (B1+B2 F M F p )=19035.1 × (1.63+(1.66 ×1×1 ) )=$ 62,625

8B) Calculation for Reactor R-102:Table (): Bare Module Reactor Costing Coefficients

Process Vessel Vertical

Diameter (M)


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.


log10 C po=K1+ K2 log10 ( A )+K 3 [ log10 ( A ) ]2=3.4974+0.4485× log10 (164 )+0.1074 × ¿¿

C po=$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=C po ( B1+B2 F M F p )=184885.871 × (2.25+(1.82 ×1×1.38 ) )=$562,578

AKA 86

Appendix C: Manufacturing Cost Sample Calculations

Operating labor Cost Calculation:

NOL=√6.29+31.7P2+0.23 N np

P= 0

Nnp= 22

NOL=√6.29+31.7 02+0.23 ×22=3.369≈ 4

NO=4.5×NOL

¿4.5× 4=18

COL=NO × AW

AW= 53,500COL=18× $53,500=$ 963,900

As shown in Table (), 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 Multiplying8×5× 49=1960 and total hour/year is

24× 350=8400. The number of operator hired for each operators is 84001960= 4.3

and this number should be rounded to 4.5.

Table (): Operating labor cost variable

Number of week in year 52Number of weeks for vacation 3

Number of shift per week 5Hours each shift 8Total hours /year 1960

Total hours of operation 8400Number of operator hired for each

operator 4.5

AKA 87

NOL 4Annual mean Wages in 2016 53,550Total number of operators 20

COL $1,071,000

Calculation For Waste Water Cost:

CWT=Flow rate (MTYR )×Price( $

MT)

Flowrate = 1645837.2 MTYR

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−hryr

CElectricity=amount of eletricity needed × priceof 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 kgyr

CCW=amount of CW needed × price of CW

CCW=258,002,773,140× 0.0000175=4,502,697 $yr

AKA 88

Calculation for High Pressure Steam cost:

The price of HPS is 0.01459 $kg

The total amount of HPS is 266,292,516 kgyr

CHPS=amount of HPS needed × priceof HPS

CHPS=266,292,516× 0.01459=3,885,247 $yr

Calculation for Low Pressure Steam

The price of LPS is 0.01348 $kg

The total amount of LPS is 1,595,689,056 kgyr

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 GJyr

CRefrigerant=amount of Refrigerant duty needed× priceof 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:

AKA 89

Feed of the process:

The Amount of Methanol needed is 2640000 MTyr

The price of Methanol is 235 $kg

CMethanol=amount of Methanol needed× priceof 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

C Al2O3Catalyst=Mass of Al 2O 3catalyst needed × price of Al2O3Catalyst

C Al2O3Catalyst=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=Massof HMOR catalyst needed × price of HMORCatalyst

CHMOR Catalyst=344.4×120=$41.328

Total Row Material Cost CRM=CMethanol+C Al 2O 3Catalyst +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)

AKA 90

COM=$ 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∧Clerical Labor=0.18× COL

Direct Supervisory∧Clerical Labor=$173,502

Maintenance and Repairs:

Maintenance and Repairs was calculated using the formula below.

Maintenance∧Repairs=0.06× FCI

Maintenance∧Repairs=$10,524,960

Fixed Manufacturing Cost (FMC):

¿ Manufacturing Cost=0.708×COL+0.068× FCI

FMC=$12,610,729

Local taxes and Insurance:

Local taxes∧Insurance=0.032× FCI

Local taxes∧Insurance=$5,613,312

Plant Overhead costs:

Plant Overhead costs=0.708 ×COL+0.036 × FCI

Plant Overhead costs=$6,997,417

Lab Charges:

Lab Charges=0.15×COL

AKA 91

Lab Charges=$144,585

Patents and Royalties:

Patents∧Royalties=0.03×COM

Patents∧Royalties=$28,683,712

General Expenses:

General Expenses=0.177×COL+0.009× FCI+0.16 ×(COM)

General Expenses=$ 154,729,154

Administration Costs:

AdministrationCosts=0.177 ×COL+0.009×(FCI )

AdministrationCosts=$ 1,749,354

Distribution and Selling Costs:

Distribution∧SellingCosts=0.11×COM

Distribution∧SellingCosts=$105,173,612

Research and Development:

Research∧Development=0.05×(COM )

Research∧Development=$ 47,806,187

Table (): Electricity Price Extrapolation (3)

year cent/kW-h

2005 5.732006 6.162007 6.392008 6.962009 6.832010 6.772011 6.82

AKA 92

2012 6.672013 6.892014 7.012015 6.892016 7.18

The data in the table

was used to plot the

correlation in the

figure above and

extrapolate linearly to

find the cost of electricity in 2016.

Natural Price data (12)

Natural Gasyear Price ($/MBtu)2005 10.082006 7.582007 7.642008 9.532009 4.212010 4.612011 4.132012 2.792013 3.732014 4.372015 369.00%

Total Annual Change -51.32%Annual Change -5.70%

AKA 93

2004 2006 2008 2010 2012 2014 20165.0

5.5

6.0

6.5

7.0

7.5

f(x) = 0.0919090909090909 x − 178.09R² = 0.591464134435103

Price of Electricity

Year

Price

in C

ents

/kW

-hr

Appendix D: Profitability Analysis Sample Calculations

Appendix E: Literature Review

Figure #: Propylene Downstream Uses??

AKA 94

Figure #: Global Propylene Consumption (IHS, 2015)

Figure #: Propylene and Ethylene Price Trends in the Middle East??

AKA 95

Figure #: Process Routes to Producing Propylene (Jasper, 2015)

Discuss the stages we went through to designing the MTP reactor

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

AKA 96

3- Developing a relief scenarios

4- Determining the wort case scenario and sizing the valves

5- Design the relief system

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.

Source Present of Accident

Electrical 23Smoking 18Friction 10

Overheated Materials 8Hot Surfaces 7

Burner Flames 7Others 27

AKA 97

Figure (1): Fire triangle

Table (1): Ignition Sources of Major Fires [1]

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 aljubail 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 are manipulated to be boiler feed water to reduce the cost.

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

AKA 98

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 [2].

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 nitrogendiluted air with a somewhat elevated CO2

AKA 99

content. It catalyst regeneration gases are vented to the atmosphere

because the amount is not significant.

AKA 100

Table #: 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

In the case of inhalation, remove person to fresh air. If person is not breathing, ensure an open airway and provide artificial respiration

AKA 101

labeled

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 102

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 103

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.

Table (3): Risk Assessment of Chemical Plant [3]

Risk Assessment

What is the

hazard

Who could be harmed

ExistingProcedures

Needed actions

How the assessment could be transferred to

an actionwhom 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

Insulating electrical

wires. Staff trained in electrical

safety

Remind staff to

check any electrical

equipment before using

Supervisor

During installation preventive

maintenance

AKA 104

electrical equipme

ntit.

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 105

Table (3): Environmental impact assessment of methanol to propylene plant

component

TypeComposition

(Mole fractions)

ReservoirKSA Regulations

[4] Actions neededProduct

Side Product

Storage

TankTreatment

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

See Appendix A

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

See Appendix AWaste Water

Water 99.9 %

Royal Decree No. 38/ Dated 16.06.1427 -

AKA 106

12/7/2006See Appendix A

Appendix F: Project Timeline

Table #: Project Tasks Performed During Spring 2016

Week # 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15Proposal ReviewLiterature ReviewSimulationOptimizationCapital CostOperating CostPoster DesignProfitability AnalysisSensitivity AnalysisFinal PresentationFinal Report

AKA 107

Documents

my.fit.edumy.fit.edu/~akurdi2012/Plant Design 2/Design Project/Final Repor… · Web viewmy.fit.edu