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The University of Adelaide

CHEM ENG 3033: SIMULATION AND CONCEPT DESIGN

Cumene Production Plant Group:

7

Group members: Nguyen Mai Thanh Le Le Nha Trang Tran Nguyen Minh Phuc Truong The Nhut Nguyen Trung Hieu Nguyen

Advisor & Teacher: Steven Amos and Afshin Karami

a1746723 a1736508 a1746690 a1746724 a1746723

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Summary project Based on Dr Who Chemicals Ltd, the predicted rise of phenol-derived plasticizers demand lead to the requirement of phenol, which could be made from cumene. Then, a grassroots cumene have been designed in order to produce phenol feedstock from the alkylation reaction of benzene and propylene. A process flow diagram was set up by using relevant theories, mean-end analysis and software package. To satisfy the requirement, the process was designed to produce 100,000 metric ton cumene per year with 99.9% purity by weight. The calculation and HYSYY simulation lead to the optimal results the ideal volume and diameter of PFR reator would be 21.02 m3 and 1.6 m. Necessary operating conditions were specified: the temperature of inlet stream is set at 350oC, the pressure of 2000kPa and benzene/ propylene feed ratio of 1.4:1. The optimization of reactor resulted in the optimized temperature of 450oC at the isothermal conditions, the effects of fluidized bed and transalkylation reactor on the increased flow rate of desired product as well as overall process performance. In term of isothermal conditions, the optimization of reactor can be achieved at the inlet stream temperature of 420oC while an extra transalkylation reactor results in the growth of cumene recovery as well as reduction in the DIBP mass flow rate. Two distillation columns were used to separate benzene, cumen and DIPB. Base on simulation result, the first column is design to be 122 𝑓𝑡 height and 4.33 𝑓𝑡 in diameter with 27 theorical stages, with O’Connell correlation the actual number of stages is determined to be 60 stages. The second column which use to separated cumen and DIPB is design to be 131 𝑓𝑡 in height and 1.833 𝑓𝑡 in diameter with 27 theorical stages, using the same correlation as the first one the actual number of stages is 65 stages. Furthermore, after optimization process, the suggest temperature of the feed before entering the 1st column is 87oC with the reflux ratio 1. Condenser and reboiler are also found at 208.6 and 233.8 kPa, respectively to get the mole fraction 0.9659 of cumen in the bottom stream. Optimization of 2nd column needs the feed steam temperature 190oC with the pressure column is around 233.5 kPa and with the reflux ratio 0.66 in order to obtain the 99.91 %wt cumene. Preliminary design of the entire cumene production, heat exchanger plays a significant role in saving energy in terms of temperature hot and cold stream plus the utilities. A combination quite differences between the theoretical and practical heat exchanger operations is run in the flowsheet. From the result, the overall heat transfer coefficient in the flowsheet is 1803 [kJ/hm2-C] and the outlets of both tube and shell temperature meet the desired outcome to reach the optimization of the heat transfer machine with the fouling coefficient is 10 -6 and number of baffle segments is 130 mm. The Net Present Valued is considered to evaluate the economic possibility of plant in two case of pure propylene feed stream and prolylene feed stream contains 5% propane. It is clear that only the 95% pure propylene pathways yielded positive NPV valued, thus it is highly recommended that Dr. Who Chemicals uses this pathway for cumene production process. The cumene production process produce the 100,000 meter cubic ton per year, the cumene purity achieved is higher than the minimum expected purity is 99% with the mass flow rate 100,690 kg/hour. Mass flow rate of benzene and propylene feed steam 8920 k/h and 5360 kg/h, respectively.

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Table of Contents 1.

INTRODUCTION ...................................................................................................... 12

1.1.

Project Description ................................................................................................. 12

1.2.

Project Aim and Objective ..................................................................................... 12

2.

LITERATURE REVIEW .......................................................................................... 13

2.1.

Introduction of Cumene ......................................................................................... 13

2.2.

Chemical hazards ................................................................................................... 14

2.3.

Chemical properties ............................................................................................... 15

2.4.

Fluids Packages ...................................................................................................... 17

2.5.

Degree of freedom................................................................................................... 18

2.6. Catalyst ................................................................................................................... 18 2.6.1. Solid phosphoric acid ........................................................................................ 18 2.6.2. Aluminium chloride........................................................................................... 18 2.6.3. Zeolite ............................................................................................................... 18 2.7.

Plant location .......................................................................................................... 19

2.8.

Q-MAX™ Process Description for cumene production: ...................................... 20

3.

MEAN-END ANALYSIS ........................................................................................... 22

3.1. Mean-Ends Analysis ............................................................................................... 22 Step 1: Chemical reactions involved ............................................................................. 22 Step 2: Evaluation of alternative pathways .................................................................. 23 Step 3: Distribute the chemicals .................................................................................... 26 Step 4: Eliminate differences in composition ................................................................ 29 Step 5: Eliminate differences in temperature, pressure and phase ............................. 30 Step 6: Integrate tasks ................................................................................................... 32 3.2 Description of Process Plant Operating Condition .................................................... 33 3.3 4. 4.1.

The goals for overall optimization ......................................................................... 34 BASE CASE DESIGN ............................................................................................... 35 Reactor-parameter set-up ...................................................................................... 35

4.2. Summary of Mass and Energy Balance from HYSYS Simulation: ...................... 36 4.2.1. Material mass balance in the main equipment: ................................................... 36 4.2.2. Total energy released and from the process and required the process ................. 38 4.2.3. Summary the amount of cooling water used in the Cumene production process . 38

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5.

PROCESS DESIGN ................................................................................................... 39

5.1. Reactor design ........................................................................................................ 39 5.1.1. Gibbs Reactor .................................................................................................... 39 5.1.2. Reactor Sizing ................................................................................................... 41 5.2. Distillation column.................................................................................................. 43 5.2.1. Background ....................................................................................................... 43 5.2.2. Basic equipment and operation .......................................................................... 43 5.2.3. Basic Principles and Equations .......................................................................... 44 5.3. Heat Exchanger ...................................................................................................... 47 5.3.1. Heat transfer coefficient .................................................................................... 47 5.3.2. Pressure drops ................................................................................................... 47 5.3.3. Shellside design ................................................................................................. 47 5.3.4. Fouling factor .................................................................................................... 48 5.3.5. Tube layout patterns .......................................................................................... 48 5.3.6. Tube pitch ......................................................................................................... 49 5.3.7. Baffling ............................................................................................................. 49 5.3.8. Shellside stream analysis ................................................................................... 51 5.3.9. Mean temperature difference ............................................................................. 51 5.3.10. Temperature profile distortion ........................................................................... 52 5.3.11. Relevant Theory ................................................................................................ 52 6.

SUMMARY PROCESS OPTIMIZATION............................................................... 55

6.1. Reactor optimization .............................................................................................. 55 6.1.1. Reactor Performance - Isothermal Temp ............................................................ 55 6.1.3. Reactor Heat Transfer - Non-Isothermal Temp. ................................................. 57 6.1.4. Reactor Configuration - Fluidized Bed Simulation............................................. 58 6.1.5. Additional Reactor - Transalkylation Reactor .................................................... 60 6.1.6. Raw Material – Propylene ................................................................................. 62 6.2. Distillation Optimization: ....................................................................................... 63 6.2.1. Column T-101 – Column temperature................................................................ 64 6.2.2. Column T-102 – Column pressure ..................................................................... 66 6.2.3. Column T-101 – Reflux ratio ............................................................................. 68 6.3. Heat exchanger ...................................................................................................... 70 6.3.1. Input parameters in HYSYS .............................................................................. 70 6.3.1.1. Theoretical heat exchanger ............................................................................. 70 6.3.1.2. Practical heat exchanger ................................................................................ 71 6.3.2. Graph of heat exchanger performance - Fouling of tube..................................... 72 6.3.2.1. Tube fouling on overall U .............................................................................. 72 6.3.2.2. Tube fouling on pressure drop of shell and tube ............................................ 72 6.3.3. Graph of heat exchanger performance – Fouling of shell ................................... 73 6.3.3.1. Shell fouling on overall U .............................................................................. 73 6.3.3.2 Shell fouling on pressure drop of shell and tube ................................................ 73 6.3.4. Graph of heat exchanger performance – Number of baffle segments .................. 74 6.3.4.1. Number of baffle segments on overall U ........................................................ 74 6.3.4.2 Number of baffle segments on pressure drop of tube and shell .......................... 74

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6.3.5. 6.3.6. 7.

Analysis graph ................................................................................................... 75 Result ................................................................................................................ 76

ECONOMIC APPRAISAL ....................................................................................... 77

7.1.

Capital cost ............................................................................................................. 78

7.2.

Chemical engineering plant cost index (CEPCI) ................................................... 79

7.3.

Operating expenses (OPEX) .................................................................................. 79

7.4.

Equipment sizing and costs .................................................................................... 80

7.5.

Material of Construction ........................................................................................ 80

7.6.

Carbon Steel and Stainless Steel ............................................................................ 80

7.7.

Utilities .................................................................................................................... 81

7.8.

Operating Labour................................................................................................... 82

7.9. Raw materials and profits ...................................................................................... 82 Pathway 1 with 5% propane impurity .............................................................................. 82 Pathway 2 with more than 99% propylene purity ............................................................. 83 7.10. Net Present Value (NPV) ........................................................................................ 84 7.11. Payback Period ....................................................................................................... 84 8.

CONCLUSION .......................................................................................................... 85

9.

REFERENCES........................................................................................................... 86

10.

BIBLIOGRAPHY................................................................................................... 88

APPENDIX ........................................................................................................................ 89 Appendix 1 – Calculation procedure of reactor ............................................................... 89 Appendix 2 – Outline equipment ...................................................................................... 90 Appendix 3 – Calculation procedure for distillation column .......................................... 93 1. Splitter...................................................................................................................... 93 2. Short-cut column ...................................................................................................... 94 3. Rigorous distillation columns ................................................................................... 95 3.1. Column efficiency ............................................................................................. 95 3.2. The number of actual stages .............................................................................. 97 3.3. Height of column ............................................................................................... 97 3.4. Column diameter ............................................................................................... 98 3.5. Multi-pass Trays .............................................................................................. 102

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3.6. Pressure drop ................................................................................................... 103 3.7 Other dimensions ................................................................................................. 106 Appendix 4 – Calculation procedure for HEX ............................................................... 107 Appendix 5 – Economic Evaluation................................................................................ 113 Fixed Capital Investment ............................................................................................... 113 Labor Costs ..................................................................................................................... 97 Utilities............................................................................................................................ 99 Raw materials and profits ................................................................................................ 99 Summary of all costs ..................................................................................................... 100 Net Present Value calculation tables .............................................................................. 101 Appendix 5- HAZOP ....................................................................................................... 104 Appendix 6 – Date of meeting ......................................................................................... 124

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List of figures Figure 1 – Fluid Package ..................................................................................................... 17 Figure 2 – Q-MaxTM process for Cumene production .......................................................... 20 Figure 3 – The general flowsheet for cumene production process ........................................ 22 Figure 4 – Overall material balance for reactor .................................................................... 26 Figure 5 – Overall material balance for mixer...................................................................... 28 Figure 6 – Overall material balance for distillation column 2 (cumene column) ................... 28 Figure 7 – Overall material balance for separator ................................................................ 29 Figure 8 – Schematic representation of Eliminate differences in composition ...................... 30 Figure 9 – Schematic representation of eliminate differences in temperature, pressure and phase. .................................................................................................................................. 32 Figure 10 – Development of PFD for cumene production .................................................... 32 Figure 11 – Cumene manufacturing process simulation on HYSYS simulation ................... 33 Figure 12 – Gibbs reactor icon in HYSYS simulation.......................................................... 39 Figure 13 – Inlet feed stream information summary............................................................. 40 Figure 14 – The maximum conversion can be achieved from the Gibbs reactor ................... 40 Figure 15 – Plot of reactor volume again the main conversion rate for optimization process 41 Figure 16 – Plot of reactor volume again the main conversion rate for optimization process 42 Figure 17 – Continues distillation column ........................................................................... 43 Figure 18 – Types of tube layout pattern ............................................................................. 48 Figure 19 – Types of segmental baffe .................................................................................. 49 Figure 20 – Countercurrent flow and Cocurrent flow ........................................................... 51 Figure 21 – The effect of isothermal temperature on reactor conversion and main production fraction................................................................................................................................ 55 Figure 22 – The impact of Benzene/Propylene ratio on conversion and main production fraction................................................................................................................................ 56 Figure 23 – Conversion and main production fraction against outlet temperature ................ 57 Figure 24 – Cumene production with fluidized bed simulation ............................................ 58 Figure 25 – Cumene molar mole of the outlet stream without 10% bypass stream ............... 59 Figure 26 – Cumene molar mole of the outlet stream with 10% bypass stream .................... 59 Figure 27 – The parameters for reaction set ......................................................................... 60 Figure 28 – PFD for cumene manufacturing with transalkylation reactor ............................. 61 Figure 29 – Hydraulic plots of column T-101 ...................................................................... 63 Figure 30 – A plot of feed temperature and cumene mole fraction ....................................... 64 Figure 31 – A plot of feed temperature and reboiler/condenser duties .................................. 64 Figure 32 – A plot of condenser pressure and cumene mole fraction.................................... 66 Figure 33 – A plot of condenser pressure and condenser/reboiler duties .............................. 66 Figure 34 – A plot of condenser pressure and cumene mole fraction.................................... 68 Figure 35 – A plot of condenser pressure and condenser/reboiler duties .............................. 68 Figure 36 – Expected temperature of inlet and outlet of tube and shell (Amos. S., University of Adelaide) ........................................................................................................................ 70 Figure 37 – Tube fouling against overall U in theoretical and experimental heat exchanger . 72 Figure 38 – Fouling coefficient of tube in theoretical and practical heat exchanger .............. 72 Figure 39 – Shell fouling against overall U in theoretical and experimental heat exchanger 73 Figure 40 – Fouling coefficient of shell in theoretical and practical heat exchanger ............. 73 Figure 41 – Baffle spacing against overall U in theoretical and experimental heat exchanger ........................................................................................................................................... 74 Figure 42 – Baffle spacing in theoretical and practical heat exchanger ................................ 74 Figure 43 – Summary cost for three different type of energy ............................................... 81 Figure 44 – Operating cost for DIPB treatment and total utilities cost .................................. 81

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Figure 45 – PFD for cumene production .............................................................................. 90 Figure 46 – Design of splitter X-101 ................................................................................... 93 Figure 47 – Design of splitter X-101 ................................................................................... 93 Figure 48 – Flooding velocity determination plots for column T-100................................... 99 Figure 49 – Flooding velocity determination plots for column T-101................................. 100 Figure 50 – A plot of liquid flow rate (gal/min) and column diameter (ft) .......................... 102 Figure 51 – Overall Heat-Transfer Coefficients in Tubular Heat Exchangers ..................... 107 Figure 52 – Reynold number-Friction factor plot ............................................................... 111 Figure 53 – Reynold number-Friction factor plot ............................................................... 112

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List of Table Table 1 – Chemical hazards................................................................................................. 14 Table 2– Chemicals properties 1 ......................................................................................... 15 Table 3 – Chemicals properties 2......................................................................................... 15 Table 4 – Chemical properties ............................................................................................. 16 Table 5 – Table of degree of freedom .................................................................................. 18 Table 6 – Requirement of product composition in Cumene production process ................... 22 Table 7 – Catalyst properties given by Dr. Who Chemicals ................................................. 23 Table 8 – Physical and chemical properties of main substance............................................. 23 Table 9 – The price of raw material and profit of product (University of Adelaide 2018) ..... 25 Table 10 – Boiling point (oC) of the main subschemicals at 1 atm ....................................... 29 Table 11 – Active energy and constant value for kinetic reactions ....................................... 36 Table 12 – Summary the mass balance of benzene and propylene feed stream ..................... 36 Table 13 – Summary the mass balance of reactor ................................................................ 36 Table 14 – Summary the mass balance of separator ............................................................. 37 Table 15 – Summary the mass balance of distillation column .............................................. 37 Table 16 – Total energy required and from the process ........................................................ 38 Table 17 – Total energy released from the process .............................................................. 38 Table 18 – Summary of the amount of cooling water used in the Cumene production process ........................................................................................................................................... 38 Table 19 – Summary of reactor sizing results ...................................................................... 42 Table 20 – Comparison on Cumene performance between originated process and with additional reactor ................................................................................................................ 61 Table 21 – Comparison on Cumene performance between two pathways ............................ 62 Table 22 – Input parameters of distillation column .............................................................. 63 Table 23 – Assumption and calculation all parameters in appendix (x) ................................ 70 Table 24 – Auto sizing all parameters in the Rigorous model .............................................. 71 Table 25 – Tube fouling against overall U in theoretical and experimental heat exchanger .. 72 Table 26 – Theoretical and experimental results of heat exchanger ...................................... 76 Table 27 – Summary of costs .............................................................................................. 77 Table 28 – Cost study of project expenditure ....................................................................... 78 Table 29 – Material factors associated with different materials (Amos 2018) ...................... 78 Table 30 – Pressure factors associated with different pressures (Amos 2018) ...................... 79 Table 31 – Various costs of raw materials required for plant operation with 5% propane impurity in fee ..................................................................................................................... 83 Table 32 – Various the costs of raw materials required for plant operation with 99% propylene purity in feed. ..................................................................................................... 83 Table 33 - Net Present Value ............................................................................................... 84 Table 34 – Outline equipment ............................................................................................. 90 Table 35 – Determined result of pressure in distillation and bottom in column X-100 and column X-101 ..................................................................................................................... 94 Table 36 – Specifying the parameter for column T-102 and column T-103 .......................... 94 Table 37 – Summary the result of the output in column T-102 and column T-103 ............... 95 Table 38 – Temperature of each stage in column T-100 and column T-101 ......................... 95 Table 39 – Summary the mass density and surface tension in column T-100 and column T101 ...................................................................................................................................... 95 Table 40 – Summary the viscosity of each stages in column T-100 and column T-101 ........ 96 Table 41 – Determination K-value for each stage in column T-100 and column T-101 ........ 96 Table 42 – Determination of the column efficiency for column T-100 and column T-101 ... 97 Table 43 – Calculating the number of actual stages for column T-100 and column T-101.... 97

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Table 44 – Summary the height of column for column T-100 and column T-101................. 97 Table 45 – Parameter specification in column diameter calculation ..................................... 98 Table 46 – Summary the liquid flow rate for column T-100 and column T-101 ................. 102 Table 47 – Vapour velocity in 2 columns .......................................................................... 103 Table 48 – Summary the velocity of hole for column T-100 and column T-101 ................. 103 Table 49 – Summary the dry tray pressure drop for column T-100 and column T-101 ....... 103 Table 50 – Summary the weir height for column T-100 and column T-101 ....................... 103 Table 51 – Summary the equivalent head on tray for column T-100 and column T-101 ..... 104 Table 52 – Summary the pressure drop due to surface tension for column T-100 and column T-101 ................................................................................................................................ 105 Table 53 – Summary the total heat loss for column T-100 and column T-101 .................... 105 Table 54 – Summary the tray pressure drop for column T-100 and column T-101 ............. 105 Table 55 – Summary the result of other dimensions for column T-100 and column T-101 . 106 Table 56 – Fixed Capital Cost for design process using the propylene feed stream contains 5% propane ......................................................................................................................... 95 Table 57 – Fixed Capital Cost for design process using the propylene feed stream (99%) .... 96 Table 58 – Operating labour ................................................................................................ 97 Table 59 – Detail calculation on annual cost ........................................................................ 98 Table 60 – Summary of utilities required in the process with propylene feed stream contain 5% propane ......................................................................................................................... 99 Table 61 – Summary of utilities required in the process with propylene feed stream (99%) . 99 Table 62 – Estimation of raw material cost and Profits (Propylene feed stream contain 5% propane impurity) ................................................................................................................ 99 Table 63 – Estimation of raw material cost and Profits (Propylene feed stream (99%) ......... 99 Table 64 – Summary of all costs ....................................................................................... 100 Table 65 – Summary of NPV and other economic indicator for pathways 1....................... 102 Table 66 – Summary of NPV and other economic indicator for pathways ......................... 103

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Group allocation of work Group member’s name Trung Hieu Nguyen

Student’s number A1746723

The Nhut Nguyen

A1736508

Nguyen Minh Phuc Truong

A1746690

Le Nha Trang Tran

A1746724

Nguyen Mai Thanh Le

A1746589

Allocation of work Heat exchanger design and optimization Column design and optimization Degree of Freedoms Manual calculation of column sizing Manual calculation of heat exchanger sizing Summary for mass and energy balances Base case design, Mean ends analysis Reactor optimization Economic evaluation Reactor kinetics Base case design Mean ends analysis Column design and optimization HAZOP assessment Record meeting minutes Format and layout Heat exchanger design and optimization, Process synthesis Analysis environmental issues, Safety consideration Determine feasibility of plant location Analysis environmental issues Manual calculation of heat exchanger sizing Column design and optimization Safety consideration Determine feasibility of plant location Analysis environmental issues Manual calculation of column sizing

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1. Introduction 1.1. Project Description Dr Who Chemicals Ltd has recently predicted that the demand for plasticizers derived from phenol will increase. Phenol is majority produce from cumene, which is used primarily as a reactant. Since cumene is the main feed stock, the company would like to design and construct a grassroots cumene plant to satisfy the increased requirement of the component. The new plant should produce 100,000 metric ton of cumene per year from benzene and propylene.

1.2. Project Aim and Objective Considering the demand of the new plant, the design and optimization of the cumene production process was undertaken by using software package. The aim of the project: The cumene chemical plant is designed to obtain 100,000 metric tons of cumene per year by applying fundamental knowledge and theories of chemical engineering into cumene production project. Also, software Aspen Hysys v10 simulation is used to optimise an industrial process. The objectives of this projects are: -

Achieving 99.9%wt of cumene production.

-

Optimising, sizing and producing of the major equipment design, including reactor, columns and heat exchanger.

-

Estimating the interest rate and payment of undertaking of the project.

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2. Literature review 2.1.Introduction of Cumene Chemical names: Cumene; Isopropylbenzene Molecular formula: C9H12; C6H5CH(CH3)2 Molecular weight: 120.195 g/mol Density: 0.862 g/mL at 25oC Cumene (Isopropylbenzene) is an organic compound, which is an alkyl aromatic hydrocarbon. It is a volatile and colourless liquid that has a boiling point of 152oC at atmosphere pressure. Cumene is used for producing phenol, acetone and a-methyl styrene as a solvent and intermediate. It is also used as a catalyst for the production of acrylic and polyester resins. Short-term inhalation exposure to cumene may cause headaches, dizziness or unconsciousness in humans. Cumene also causes skin and eye irritant when direct contact. On an industrial scale, cumene is generally converted to cumene hydroperoxide, which is primarily used as an intermediate for the production of acetone and phenol. Cumene also used to manufacture bisfenol-A, polycarbonate, epoxy resins, nylon 6, etc. At present, the cumene production is one of the world’s five biggest large scale production.

14

2.2. Chemical hazards Table 1 – Chemical hazards Chemical Propylene

Propane

Benzene

Hazard description  Non-dangerous hazardous substance  Highly flammable  Explode under high pressure condition  Extremely flammable  Explosive (gas/air phase)  Hazardous substance  Volatile  Flammable

Cumene

 Hazardous substance  Highly flammable  Vapour explosion

DIPB

 Non-dangerous hazardous substance

Nitrogen

 Explosive (gas phase)

Health hazard  Dizziness  Drowsiness  Unconsciousness  Freezing burn

 Dizziness  Asphyxiation (high concentration)  Damage to central nervous system  Damage to bone marrow  Damage to immune system  Developing lymphatic  Hematopoietic cancers  Myelogenous leukemia  Lymphocytic leukemia  Headaches  Dizziness  Drowsiness  Slight incoordination  Unconsciousness  Skin irritant  Eye irritant  Skin irritation  Respiratory irritation  Drowsiness  Dizziness  Dizziness  Asphyxiation

Environment impact  Smog creation in atmosphere  Contaminate water and soil  Damage to organs of animals and plants  Causing genetic defects  Toxic to aquatic life  Clean-burning fuel

 Smog creation in atmosphere  Contaminate water and soil  Damage to organs of animals and plants  Causing genetic defects  Toxic to aquatic life

 Moderate chronic toxicity to aquatic life

 Toxicity to aquatic life

15

2.3.Chemical properties Table 2– Chemicals properties 1

Chemical formula IUPAC name Chemical structure

Molecular weight Appearance Odour Density Boiling point Melting point

Propylene C3H6 CH2CHCH3 prop-1-ene

Benzene C6H6

42.081 g/mole colourless gas odourless Gas: 1.81 kg/m3 Liquid: 613.9 kg/m3 -47.68 deg C -185.30 deg C

78.114 g/mole Colourless liquid Aromatic odour 0.8765 g/cm3

benzene

80°C 5.5 °C

Table 3 – Chemicals properties 2 Chemical formula IUPAC name Chemical structure

Molecular weight Appearance Odour Density Boiling point Melting point

Cumene 1. C9H12 2. C6H5CH(CH3)2 cumene

120.195 g/mol Colourless liquid Gasoline-like odour 0.862 g/cm3 152°C -96.9 °C

DIPB C12H18 1,4-di(propan-2-yl)benzene

162.276 g/mol liquid Odourless 0.8568 g/cm3 210.3 oC -17.1 oC

16

Table 4 – Chemical properties Propane Chemical formula3. C3H8 4. CH3CH2CH3 IUPAC name propane

Nitrogen N2

Carbon monoxide CO

Carbon dioxide CO2

molecular nitrogen

carbon monoxide

carbon dioxide

28.014 g/mol Colourless gas Colourless liquid odourless Gas: 1.2504g/L (0oC, 1013mbar Liquid: 0.808 3 kg/m -196°C -210.01°C

28.01 g/mol Colourless gas Odourless 1.14 g/cm3

44.009 g/mol Colourless gas Odourless 1.98 g/cm3

-191.5°C -205.02°C

-78.464°C -56.5°C

Chemical structure

Molecular weight Appearance Odour Density

44.097 g/mol Colourless gas Colourless liquid odourless 0.493 g/cm3

Boiling point Melting point

-42°C -189.7°C

17

2.4. Fluids Packages There are two fundamental consideration for choosing a Fluids Package, including operating conditions and specific system under consideration. For the fluids package of the project, the Peng Robinson model is suitable. The Peng Robinson model, which has temperature range is over -271oC (T 99.9𝑤𝑡% purity)

$1,120/1000kg

Propylene (>99.9 𝑤𝑡% purity)

$1,570/1000kg

Propylene ( 5mol% propane impurity)

$880/1000kg

Cumene

$1,434/1000kg

Propane (fuel gas)

$833/1000 kg

26

The gross profit calculations have indicated that pathway contains 5% propane seems to bring more profit than higher-impurity pathway. Therefore, the feed with 5% propane would be more likely to be used in this plant design as well as the need of one or more separators to separate the inert propane. In additional, the high temperature in the inlet feed stream and increasing reactor sizing would be required to achieve expected cumene conversion (Luyben 2014).

Step 3: Distribute the chemicals In order to carry out mass balance calculations, the assumptions were established to as follows: 1. The impure propylene feed (5 mol% propane) was chosen for calculations based on pathway 1. 2. Base on the HYSYS simulation results, the conversion rates of main and side reactions are 95.72% and 3.51% respectively. 3. The ratio of Benzene-to-Propylene is 1.5: 1. 4. The molar fraction of cumene and DIBP from the product stream leaving distillation column 2 are 99.99 wt% and 0.01 wt% respectively. 5. The recovery ratio of cumene from both distillation column is equal to 1. 6. The recovery ratio of benzene from partial separator column is equal to 1. 7. The operating period of plant is continuously 330 days. Overall material balance Overall material balance for reactor

Figure 4 – Overall material balance for reactor

27

The operating period of plant is continuously 330 days. Hence, the cumene molar flow produced 𝑡

𝑡

= 100,000 𝑦𝑒𝑎𝑟 = 100,000 𝑦𝑒𝑎𝑟 × Or =

𝑘𝑔 ℎ𝑟 𝑘𝑔 120.194 𝑚𝑜𝑙

12626.26

= 105.05

1000𝑘𝑔 1𝑡

1𝑦𝑒𝑎𝑟

× 330 𝑑𝑎𝑦𝑠 ×

1 𝑑𝑎𝑦𝑠 24 ℎ𝑟

𝑘𝑔

= 12626.26 ℎ𝑟

𝑘𝑚𝑜𝑙 ℎ𝑟

Base on the assumption, the number of mole DIPB in the cumene product stream is: 𝑘𝑚𝑜𝑙 105.05 ℎ𝑟 𝑘𝑚𝑜𝑙 = × 0.001 = 0.105 0.999 ℎ𝑟 Because the recovery ratio of cumene from both distillation column is equal to 1, hence the molar mole flow rate of cumene in product stream is equal to the cumene originated from the PFR reactor: = 105.05

𝑘𝑚𝑜𝑙 ℎ𝑟

From PFD base case, conversion of R1= 95.72 % and conversion of R2 = 3.51 % R1:

𝐶3 𝐻6 (𝑃) + 𝐶6 𝐻6 (𝐵) → 𝐶9 𝐻12 (𝐶) (1)

MW 42.08

78.11

120.19

Conversion of R1=95.72% = Conversion of R1=3.51% =

𝑛𝑅1𝑃 𝑛𝐹𝑃

𝑛𝑅2𝑃 𝑛𝐹𝑃

 𝑛𝑅1𝑃 = 0.9572 × 𝑛𝐹𝑃

 𝑛𝑅2𝑃 = 0.0351 × 𝑛𝐹𝑃

From equation (1): 𝑛𝑅1𝑃 = 𝑛𝑅1𝐶 From equation (2): 𝑛𝑅2𝑃 = 𝑛𝑅2𝐶 𝑛𝑅𝐶 = 𝑛𝑅1𝐶 − 𝑛𝑅2𝐶 = 𝑛𝑅1𝑃 − 𝑛𝑅2𝑃 = (0.9572 − 0.0351)𝑛𝐹𝑃 𝑛𝐹𝑃 =

𝑛𝑅𝐶 105.05 𝑘𝑚𝑜𝑙 = = 113.92 0.9572 − 0.0351 0.9572 − 0.0351 ℎ𝑟

The mass flow of propane occupied 5% of feed stream, the molar mole flow of propane can be estimated: 𝑛𝐹,𝑃𝐼 = 113.92 ×

44.1 0.05 𝑘𝑚𝑜𝑙 × = 6.28 42.08 0.95 ℎ𝑟

28

Overall material balance for mixer

Figure 5 – Overall material balance for mixer The benzene/ propylene ratio flows the assumpition is 1.5:1, hence: 𝑛𝑚𝑖𝑥𝑒𝑑 𝑏𝑒𝑛𝑧𝑒𝑛𝑒 = 1.5 × 𝑛𝐹,𝑃 = 1.5 × 113.92 = 170.9 𝑛𝑅1𝐵 = 𝑛𝑅1𝑃 = 95.72% × 𝑛𝐹𝑃 = 109.05

𝑘𝑚𝑜𝑙 ℎ𝑟

𝑘𝑚𝑜𝑙 ℎ𝑟

Because all benzene is separated from benzene distillation column 𝑛𝑟𝑒𝑐𝑦𝑐𝑙𝑒𝑑 𝑏𝑒𝑛𝑧𝑒𝑛𝑒 = 𝑛𝑚𝑖𝑥𝑒𝑑 𝑏𝑒𝑛𝑧𝑒𝑛𝑒 − 𝑛𝑅1𝐵 = 170.9

𝑘𝑚𝑜𝑙 𝑘𝑚𝑜𝑙 𝑘𝑚𝑜𝑙 − 109.05 = 61.85 ℎ𝑟 ℎ𝑟 ℎ𝑟

Overall material balance for distillation column 2 (cumene column)

Figure 6 – Overall material balance for distillation column 2 (cumene column) 𝑛𝑅2𝑃 = 0.0351 × 𝑛𝐹𝑃 = 0.0351 × 113.93 = 4

Molar flow of DIPB produced is 𝑛𝑅2 𝐷𝐼𝑃𝐵 = 𝑛𝑅2𝑃 = 4

𝑘𝑚𝑜𝑙 ℎ𝑟

𝑘𝑚𝑜𝑙 ℎ𝑟

29

Number mole of DIPB in DIPB stream

= 𝑛𝑅2 𝐷𝐼𝑃𝐵 − 𝐷𝐼𝑃𝐵 𝑖𝑛 𝑝𝑟𝑜𝑑𝑢𝑐𝑡 𝑠𝑡𝑒𝑎𝑚 = 4 − 0.105 = 3.885

𝑘𝑚𝑜𝑙 ℎ𝑟

Overall material balance for separator

Figure 7 – Overall material balance for separator Number of moles of propane in stream Fuel gas = 𝑛𝐹𝑃𝐼 = 6.28

𝑘𝑚𝑜𝑙 ℎ𝑟

Step 4: Eliminate differences in composition Table 10 – Boiling point (oC) of the main subschemicals at 1 atm

After completing the reaction process, the stream leaving the PFR reactor contains propylene and propane, cumene, DIPB, unreacted benzene. The main purpose of plan is to achieve the possible highest cumene purity, hence it is important to split all propane and propylene from the main product stream. An understanding on the boiling point of all components are show in the table above. According the property data, it is clear that propane and propylene have relatively low boiling point range from -42oC to - 48oC. Therefore, the separating process for propane and propylene is carried out firstly by a separator. This means the temperature is smaller than benzene boiling point (under 80.1°C) and atmospheric pressure condition is set for the separator operation in

30

order to separate propane and propylene from other composition. To be more specific, propane and propylene would go through the overhead vapor stream of separator while mixture of benzene, cumene and DIPB will introduce into the first distillation column. Base on the requirement of the plant and property data, unreacted benzene needs to be drawn off by the distillation column. The necessary condition for this process is that pressure of 1 atm and temperature is under 152°C. Under this condition, benzene would be removed into the top of the distillation column, and then turn back to recycle stream. On the other hand, the bottom stream continuous to feed for the next distillation column. In order to removing cumene from undesirable product, the operating temperature needs to be increase into the range from 152 oC to 210.3oC. As a result, the expected cumene stream with the purity of 99.9% could be achieved from the distillation column while DIBP will be the main product from the bottom stream.

Figure 8 – Schematic representation of Eliminate differences in composition

Step 5: Eliminate differences in temperature, pressure and phase The manufacturing process can be operated under assumption condition which is suggested by the Luyben literature review (2010) Following Luyben suggestion, benzene fresh stream is assumed to be set at 1 atm and 25 oC while the propylene feed stream should be set at 25oC and assumed under the liquid saturated condition. By defining all these conditions, the pressure of propylene stream is 1143 kPa.

31

The feed streams enter the reactor are suggested to be under operating pressure set in the range of 20 to 30 bar. The pressure of two feed streams are elevated to expected pressure of 25 bar by a pump. Once the pressure of both feed stream are achieved, a mixer is also utilized for mixing two feed stream. Moreover, Elsevier (2010) recommend that the inlet temperature needs to be heated to 350oC and thus the reaction could be taken in gas phase. To meet these requirements, heat exchanger system includes a set of tube and shell heat exchanger as well as the furnace can be used for vaporizing and heating the mixed feed stream. The reactor effluent temperature might be grown to 450oC while the pressure could be kept constant. In order to separate the propane and unreacted propylene from the mixed fluid, the temperature of product stream which enter the phase separator must be under benzene’s boiling point (80.1oC) and 1 atm. As a result, a cooled and valve would be installed for achieving these desired conditions before the separation process was carried out. Then, a pump was provided to increase the pressure of bottom product stream include benzene, DIPB and cumene before going through the first distillation column. The pressure of the first column with a partial reboiler and total condenser is 1 bar which lead to a temperature is 75oC. Under this temperature, benzene is vaporized and separated from the initial feed stream and then was sent back to fresh benzene stream through recycle. Due to the difference in physical condition, it is necessary to liquefy the recycled benzene stream before being combining with the fresh feed. A valve and cooler need to be installed. It is suggested that the operating pressure for second column (contain a partial reboiler and total condenser) is set with lower than first column pressure. Consequently, a valve was used to adjust the pressure back to 0.75 bar. After finishing these steps, it is highly possible to gain the high cumene purity of 99.99%.

32

Figure 9 – Schematic representation of eliminate differences in temperature, pressure and phase. Step 6: Integrate tasks

Figure 10 – Development of PFD for cumene production

33

3.2 Description of Process Plant Operating Condition A HYSYS simulation for cumene manufacturing process is operated to further explain the equipment, all the chemical component and operating conditions. Base on the chemical and physical characteristics that related to main ingredients provided in the literature review, the operating condition such as pressure and temperature was specified for the process. A process flow diagram (PFD) would simulated the cumene production manuscript which was shown in the Figure 11 below.

Figure 11 – Cumene manufacturing process simulation on HYSYS simulation In the Figure 11, the feed stream of propylene (95%) and fresh benzene was presented. All initial temperature of the feed stream will be set at the same room temperature of 25 oC. However, while the pressure of benzene stream is 1 atm, the pressure of the propylene stream need to be improved to 1143kPa to avoid the vaporization at room temperature. In the next step, both feed streams went through the pumb to increase the pressure of 25 bars before mixing together by a mixer (Mix-100). The outlet stream called Mix went to a heater to meet the required temperature of 200oC (heater will be replaced by a heat exchanger when the product stream is created). On the other hand, the ideal phase for the PFR reaction is vapor phase, hence the Mix stream needs to be preheated by a fire heater. The fuel streams for the fire heater are the air stream and 100% propylene stream to calcine the material up to 350 oC (Gera 2011, p. 523). After that, the stream came out the furnace would be fed for the reactor to create cumene and DIBP with support fo new catalyst. Skogestad (2011) suggested that the outlet temperature should be set at 357oC. The assumption for the reactor sizing at this stage is the volume of 60m3 and diameter of 2m. As a result, the conversion of main reaction gains 90.45%. By adjusting the reactor output temperature, the conversion for main reaction increase to 95.72% with higher temperature of 450oC. The pressure of outlet stream decrease slightly to 2492 kPa compared to

34

the initial pressure of 2500kPa. The outlet stream then flowed through a valve and a cooler to change into the liquid-vapor phase before being separated by a separator. The separating mechanism will base on the boiling point of unreacted propylene and propane which have the lowest boiling point from -42.1oC and -48oC respectively. The operating temperature of separator is 87oC. In fact, the overhead stream of separator also contains a considerable amount of cumene, benzene and DIBP even the temperature was set under the benzene boiling point of 80.01oC. As can be show in the Figure 11, all excessive propylene and benzene from the bottom stream of separator will be splinted by the distillation column 1. Base on the operation condition, the distillation column 1 has 27 stream with the main stream is stage 14, the pressure for the condenser and reboiler was 206.8 kPa and 233.8 kPa respectively. The overhead stream leaving the column contain almost molar mole of benzene with small amount of cumene and propylene will be recycled by feeding into second Mixer with the fresh benzene stream. The main product stream then is fed into the second distillation column which has 37 stages with main stage is stage 7. The pressure was set for condenser and rebuilder is 206.8 kPa and 243.8 kPa. The final product contains the cumene with purity of 99.93% and the side product stream (DIBP) have 99.99%. Ultimately, the PFD for the plant design work well following advises and operating information which is provided by Dr Who Chemicals. The conversion of cumene was 95.72% and the purity of cumene achieves 99.93% which higher than the minimum expected conversion of the process. 3.3 The goals for overall optimization It is important to know that the recovery of cumene molar flow from both separator and two distillation columns is actually lower than 1. Therefore, in order to meet the requirement of 100,000 tons cumene per year, the purpose of overall optimization process aim to get the higher main conversion rate with slight difference with the assumption and initial base case. All detail will be specified in sec 6: Summary of process optimization.

35

4. Base case design 4.1. Reactor-parameter set-up Benzene and propylene are two main reactants used for cumene manufacturing process. Base on the catalyst properties, the plug flow reactor and the vapour phase of the feed stream are required. The alkylation reaction associate to the cumene production process is presented below:

With 𝑟1 =

𝑘1 𝑐𝑐 𝑐𝑝 𝑚𝑜𝑙𝑒 (𝑔𝑐𝑎𝑡)(𝑠)

k1 = 3.5 × 104 × exp (

−24.9 ) 𝑅𝑇

where: r1= reaction rate k1= rate constant R= gas constant T= temperature (K) Dr Who Chemicals Ltd. presented a new type of catalyst in order to reduce side product (DIBP) as well as increase the recovery of main product (cumene). Through laboratory tests, a pressure range of 20 bar to 30 bar and a temperature range of 200oC to 500oC were specified. Moreover, the outlet temperature leaving the reactor should be operated under 500oC to avoid carbon decomposition into the catalyst.

With: 𝑟2 =

𝑘2 𝑐𝑐 𝑐𝑝 𝑚𝑜𝑙𝑒 (𝑔𝑐𝑎𝑡)(𝑠)

k2 = 2.9 × 106 × exp (

−35.08 ) 𝑅𝑇

36

In order to demonstrate the cumene process in HYSYS simulation, it is necessary to add unit conversion for those parameters into expected unit. The more detail calculation for converting unit of parameter can be referred in Appendix. The constant value A and activation energy E for both main and side reaction are shown in the table below Table 11 – Active energy and constant value for kinetic reactions Parameter

Main Reaction

Side Reaction

A (m3/kmol.sec)

2.8x107

2.32x109

E (kJ/mol)

104580

147336

4.2.Summary of Mass and Energy Balance from HYSYS Simulation: 4.2.1.

Material mass balance in the main equipment:

Mass balance equation: Total mass in + total mass out - total mass generated = 0 Table 12 – Summary the mass balance of benzene and propylene feed stream Feed stream name Fresh benzene Inlet stream

Recycle bezene Propylene

Outlet stream

Mix

Component Benzene Benzene Propane Propylene Propylene Propane Benzene Propane Propylene

Mass balance Total mass balance (kg/h) 8920 8919.8 3370.6 3487 104.7 11.7 4810.2 5063 253.2 12290.5 357.8 17470 4821.9

Table 13 – Summary the mass balance of reactor

Feed stream name Inlet stream

6

Outlet stream

7

Reactor (PFR-100) Component Mass balance Total mass balance (kg/h) Benzene 12290.5 Propane 17471 357.8 Propylene 4821.9 Cumene 12759.5 DIBP 602.6 Benzene 17471 3708.5 Propane 357.8 Propylene 42.2

37

Table 14 – Summary the mass balance of separator

Feed stream name

Inlet stream

16

Fuel gas Outlet stream Liquid

Separator (V-100) Component Mass balance Total mass balance (kg/h) Cumene 12759.5 Benzene 3708.5 Propene 42.2 17471 14-iP-BZ 602.6 Propane 357.8 Cumene 124.0 Benzene 330.4 Propene 30.6 740 14-iP-BZ 1.1 Propane 253.1 Cumene 12635.4 Benzene 3378.0 Propene 11.6 16731 14-iP-BZ 601.5 Propane 104.8

Table 15 – Summary the mass balance of distillation column Distillation Column T-101 Feed stream name

Inlet

Bottom stream

stream

stream

stream DIBP pruduct stream

Mass balance

Cumene

12635.4

Benzene

0.1

14-iP-BZ cumene product

outlet

Component

13237

601.5

Cumene

12635.4

Benzene

0.1

14-iP-BZ

11.2

Cumene

0.0

14-iP-BZ

Total mass balance (kg/h)

590.3

12647

590

38

4.2.2.

Total energy released and from the process and required the process

Calculation of the energy of all the equipment are showed in the equation below: Q= mc×cp× ∆𝑇 All energy released and from the process and required the process is shown in the tables below: Table 16 – Total energy required and from the process Total Energy required Equipment Name P-100 P-101

Pump Heat exchanger Distillation column Furnace

Energy stream Q1 Q2

Utility Type Electricity Electricity

Utility per unit 5.028 kW 12.71 kW

E-101

Energy consumption (kj/h) 1.81E+04 4.58E+04 1.02E+07

T-100 Reboiler T-101 Reboiler FH-100 Total Energy balance

Q6 Q9

LP Steam LP Steam Fuel gas+ air

1642 1683

kW 5.91E+06 kW 6.06E+06 6.42E+06 2.86E+07

Table 17 – Total energy released from the process Total Energy released Energy Equipment Name stream

Utility Type Cooling E-102 Q100 Water Cooler Cooling E-103 Q106 Water Cooling Water Distillation T-100 Condenser Q101 column Cooling T-101 Condenser Q8 Water Cooling Reactor PFR-100 Q3 Water Total Energy released

Utility per unit

Energy consumption (kj/h)

2303 kW

-8.29E+06

54.32 kW

-1.96E+05

878.5 kW

-3.16E+06

1715 kW

-6.17E+06

1874 kW

-6.75E+06 -2.46E+07

4.2.3. Summary the amount of cooling water used in the Cumene production process Table 18 – Summary of the amount of cooling water used in the Cumene production process E-102 E-103 PFR-100 T-100 Condenser T-101 Condenser

Q100 Q106 Q3 Q101 Q8

Cooling Water Cooling Water Cooling Water Cooling Water Cooling Water

3.96E+05 9350 3.23E+05 1.51E+05 2.95E+05

Kg/h Kg/h Kg/h Kg/h Kg/h

39

5. Process design 5.1. Reactor design 5.1.1. Gibbs Reactor A Gibbs reactor has been used to determine the maximum conversion limits without entering any equipment sizing information and reaction set. The Figure indicates the set-up stream for Gibb reactor which the inlet stream has the same properties as the inlet stream of PFR reactor.

Figure 12 – Gibbs reactor icon in HYSYS simulation Figure 12 shows the properties and important detail of inlet stream entering the Gibbs reactor, its properties are the same with the inlet stream going to the PFR reactor. Base on the data was constructed from the Gibbs reactor, the maximum conversion limit could be found is around 99.32%, which is calculated by the ratio of the reacted propene flow rate and the propene molar flow from feed stream. To be more specific, the maximum conversation calculation would be show in the Figure 13.

40

Figure 13 – Inlet feed stream information summary

Figure 14 – The maximum conversion can be achieved from the Gibbs reactor

41

5.1.2.

Reactor Sizing

Several case studies have been run in this plant aims to determine the optimal length, diameter as well as volume for the PFR reactor. The Figure 15 shows the case study which is used to evaluate the effect of PFR reactor volume on the cumene conversion. In this case study, the conversion of the main reaction and the volume of reactor are defined as dependent and independent variables respectively. The range of 1 to 200 m3 with step size is 1, 150 steps were set for this case study. Additionally, the optimal tube volume is considered base on expected conversion rate of cumene from actual base case. According to the plot, the conversion rate of main reaction achieves 95.88% with 21.02 m3 of tube. After that, the conversion rate gains the peak of 97 % and is stable with the increase of tube volume. Moreover, comparing with the maximum conversion limit is 99.32%, 95.88% is an acceptable result. Therefore, the optimal volume of reactor is 21.02 m3.

Figure 15 – Plot of reactor volume again the main conversion rate for optimization process Diameter of reactor is also the main factor contribute to cumene conversion improvement. Another case study was set with the independent variable is the tube diameter of reactor while cumene is the dependent variables. The range from 1 to 5 m with step size is 0.01 was chosen for tube diameter. The plot of tube diameter against the cumene conversion rate is shown in the Figure 21. According to plot, the conversion rate of main reaction achieves 95.88% at the diameter of 1.6 m. After this stage, the percentage of conversion rate remained unchanged. Therefore, the optimal diameter of reactor will be 1.6 m.

42

Figure 16 – Plot of reactor volume again the main conversion rate for optimization process In those case studies, the tube diameter and volume are chosen as independent variable so HYSYS will result the optimal length of reactor, tube packing void volume and wall thickness. After running necessary case studies, important result of reactor sizing properties is presented in the Table 19. Table 19 – Summary of reactor sizing results Reactor sizing properties Total Volume (m3) Tube Length(m) Tube diameter (m) Number of Tubes (m) Wall thickness (m) Tube packing void fraction Tube packing void volume (m3)

Results 21.02 10.45 1.6 1 0.005 0.5 10.51

43

5.2.Distillation column 5.2.1. Background Distillation is a physical process used for separate a mixture of two or more substances into its components fractions of desired purity, by base on the differences boiling points of components in the mixture (McCabe, Smith and Harriott 1993, p. 521). Distillation is the most widely separation process used in many industries, such as chemical, pharmaceutical and food industries. In this project, the distillation column is used for separate a mixture of cumene and the other components to obtain a high purity of cumene as product.

5.2.2.

Basic equipment and operation

Figure 17 – Continues distillation column The liquid leaving the top of the column is the light component, while the liquid leaving the bottom of the column is the heavy component. Liquid leaving the bottom of the column is split into a bottoms product and a fraction that is made available for boiling. The reboiler is employed to boil the portion of the bottom liquid that is not drawn off as product. The vapour produced flows up through the column and comes into intimate contact with the down flowing liquid. After the vapour reaches and leaves the top of the column, the condenser is encountered where heat is removed from the vapour to condensate it. The condensed liquid is split into two streams. One is the overhead product; the other liquid stream is called reflux and is returned to the top of the column to improve the separation.

44

5.2.3. Basic Principles and Equations According to Figure 5.2 A, the following equation presents the overall material balance of the column (Treybal 1980, p.363) 𝐹 = 𝐷+𝐵 And for a component balance: 𝐹𝑥𝐹 = 𝐷𝑥𝐷 + 𝐵𝑥𝐵 The reflux ratio (RD) relates the amount of distillate that returns to the column (Robert 1980, p.384) 𝐿

𝑅𝐷 = 𝐷 As reflux ratio increases, less stages are required but larger equipment are now needed to handle the increased reflux liquid and reboiled vapour load. Thus, the fixed cost initially decreased but eventually increase again when the reflux ratio approaches total reflux. The fixed cost this falls through a minimum and then rise again to infinity. As for the operating cost, it will continue to increase with increasing reflux ratio. Typically, the optimum reflux ratio is approximately 1.2 to 1.5 times R min. However, in the actual life, the reflux ration can reach more than 2 times the minimum reflux. The efficiency of plate can be determined by O’Connel correlation (Treybal 1980, p. 423) 𝐸𝑜 = 0.492 × (μL × 𝛼)−0.245 ± 10% The range of 40%-90% aqueous solutions is acceptable conditions for efficiency plate. The overall tray efficiency describes the ratio of the number of theoretical trays to the actual number of trays required for an entire column (McCabe Smith and Harriott 1993, p.565) : 𝑁𝑡ℎ𝑒𝑜𝑟𝑦

𝐸𝑜 = 𝑁

𝑎𝑐𝑡𝑢𝑎𝑙

Increasing the number of stages will improve the separation process. If keeping feed at top section of column, the bottoms products will be more pure, and lead to the condenser load increase, while if keeping feed at bottom section of column, the top products will be more pure and lead to the reboiler load increase. There are some heuristics for design the distillation column :the height of the column should not be higher than 175 ft. Futhermore, if the tower is higher than 190 ft, the smaller tray spacing should be considered in the design, the ration of height and diameter should be in range of 20 – 30. Moreover, in the case of the diamter column is between 1.5 – 4.5 ft, both of plate and packed might be used (Amos 2018).

45

Column diameter is determined based on the constraints imposed by flooding. The equation below calculates the flooding velocity: 𝑈𝑓 = 𝐶𝑠𝑏 × (

𝜎 0.2 𝜌𝐿 − 𝜌𝐺 0.5 ) × 𝐹𝑓 × 𝐹ℎ𝑎 × ( ) 20 𝜌𝐺

If the system is non-foaming  FF = 1  FF = 0.5-0.75

If the system is foaming

𝐴

If the ration of hole area and active are (𝐴ℎ ) ≥ 1  FHA = 1 𝑎

𝐴

𝐴

If the ration of hole area and active are 0.06 ≤ 𝐴ℎ ≤ 0.1  FHA = 5(𝐴ℎ ) + 0.5 𝑎

𝑎

Velocity of vapour flow is one of the factors affect to column operation, because of the condtions of vapour, either excessive or too low. 𝑈𝐺 = (0.75 𝑡𝑜 0.85)𝑈𝐹 The flow parameter can be determined: 𝐹𝐿𝐺 =

𝐿 𝜌𝐺 0.5 ( ) 𝐺 𝜌𝐿

Thereby, the equation determined column diamnter is: 𝐷𝑡 = (

4𝐺 𝐴 𝑈𝐺 𝜋. (1 − 𝐴𝑑 )𝜌𝐺 𝑇

)1/2

𝐴

when 𝐴𝑑 is the ratio between downcomer area and total area. If: 𝑇

𝐴



FLG ≤ 0.1  𝐴𝑑 = 0.1



0.1 ≤ FLG ≤ 1  𝐴𝑑 = 0.1 +



FLG ≥ 1 𝐴𝑑 = 0.2

𝑇

𝐴

𝑇

(𝐹𝐿𝐺−0.1) 9

𝐴

𝑇

The velocity of vapour is dependent on the column diameter. Due to the column capacity, the minimum required of vapour flow are determined by the weeping while the maximum allowed for vapour flow is determined by flooding. Therefore, the column will not run in satisfactory way, if the diameter of column is not sized properly. Not only will operational problems occur, the desired separation duties may not be achieved The pressure drop of tray are required to flow vapour in tower is in the range of 0.05 – 0.15 psi/tray (McCabe, Smith and Harriott 1993, p.562). The pressure drop across the plate can be divide into 2 parts, which are the friction loss in the holes and pressure drop, because of the holdup of the liquid on the plate. ℎ𝑡 = ℎ𝑑 + ℎ𝐼 + ℎ𝜎

46

The dry tray pressure drop hd though the holes can be modified orifice equation (McCabe, Smith and Harriott 1993, p.563): 0.165 2 44.41 ℎ𝑑 = 0.186 ( ) ×( ) 0.66 41.28 The equivalent head on tray hI are estimated: 2/3 𝑞𝑙 ℎ𝑙 = 𝜑𝑒 [ℎ𝑤 + 𝐶 ( )] 𝐿𝑤 𝜑𝑒

With

𝜑𝑒 is the effective relative froth density 0.91 )

𝜑𝑒 = 𝑒 (−4257𝑘𝑠

Ks is capacity parameter (ft/s) 𝑘𝑆 = 𝑢𝑎 × (𝜌

𝜌𝐺 𝐿 −𝜌𝐺

)

0.5

The pressure drop due to the surface tension (h) is bubble gas overcome the surface tension to emerge from tray perforation ℎ𝜎 =

6𝜎 𝑔𝜌𝐿 𝐷𝐵(max)

47

5.3.Heat Exchanger 5.3.1. Heat transfer coefficient The tube side heat-transfer coefficient is a function of the parameters, which including the Reynolds number, the Prandtl number and the tube diameter. These parameters can be divided into basic parameters, which are physical properties (viscosity, thermal conductivity, and specific heat), tube diameter and mass velocity. The alteration in liquid viscosity has the most dramatic influence on heat transfer coefficient, so this physical property is absolutely significant. For turbulent heat-transfer inside the tube, the fundamental equation is: 𝑁𝑢 = 0.027(𝑅𝑒)0.8 (𝑃𝑟)0.33 or ℎ𝐷

𝐷𝐺

𝑐𝜇

( 𝑘 ) = 0.027 ( μ ) 0.8 ( 𝑘 ) 0.33 The heat transfer coefficient is influenced by the viscosity when being a parameter of the Reynolds number or a parameter of Prandtl number. Mass velocity has a firm influence on the heat transfer coefficient.

5.3.2.

Pressure drops

The increase of mass velocity leads to the increase of pressure drop, which is more rapid than the increase of heat transfer coefficient The recommendable minimum of liquid velocity inside tubes is 1.0 m/s and the maximum is 2.5-3.0 m/s. Because of the present of erosion when the velocity is very high. However, the limitation of pressure drop is necessary to control the existence of erosive velocities. The tube diameters are generally used in the CPI are 3/8, 1/2, 5/8, 3/4, 1, 1¼, and 1½ in. 3/4 in. and 1 in. are most commonly used among of these diameters. For fouling services, the tube diameters which are smaller than ¾ should not be used. The variation of pressure drop distribution, which in the various heat exchangers for a given stream in a specific circuit, may be exist to obtain a good heat transfer.

5.3.3.

Shellside design

The shellside calculations are more complicated than the tubeside calculations. The reason is on the shellside, there are a flow stream, a principal cross flow steam and four leakage or bypass streams. The shell side stream analysis is determined by using various shellside flow arrangements, various tube layout patterns and baffling designs.

48

5.3.4.

Fouling factor

Heat transfer surfaces of the heat exchanger may be covered by the deposits in the flow systems, or may be corroded by the interaction between the fluids and the material used for construction of the heat exchanger. Therefore, an additional resistance to the heat flow is existed and reduce the performance of heat transfer. The overall effect is frequently described by a fouling factor (or fouling resistance), Rf. For such an increase in the overall heat transfer coefficient, the fouling factor must be included along with the other thermal resistance. To define the fouling factor, the values of U for both clean and dirty conditions in the heat exchanger is used as: Rf = 1/ Udirty – 1/ Uclean

5.3.5.

Tube layout patterns

Tube layout patterns have 4 different types, which are triangular (30 oC), rotated triangular (60oC), square (90oC), and rotated square (45oC).

Figure 18 – Types of tube layout pattern A triangular pattern or rotated triangular pattern will contain more tubes than a square pattern or rotated square pattern. A triangular pattern generates high turbulence and it obtains a high heat transfer coefficient. Therefore, using a triangular pitch is better for heat transfer and surface area per unit length. The square pitch, which has 45oC or 90oC is commonly needed for dirty shellside services.

49

5.3.6. Tube pitch Tube pitch is a shortest distance between two adjacent tubes. The minimum tube pitch should be preferred to employ, because the minimum of tube pitch leads to the smallest shell diameter for estimating number of tubes. Increasing tube pitch Using the smaller tube pitch lead to the smaller shell diameter and reducing cost The tube pitch is generally set at 1.25 times the tube outside diameter. For the conversion of pressure drop to heat transfer, the optimum tube-pitch to tube-diameter ratio: Turbulent flow: 1.25< (tube pitch)/(tube diameter) < 1.35 Laminar flow: (tube pitch)/(tube diameter) ≈ 1.4 Reducing the pressure drop by increasing the tube pitch is not highly recommended. Because the tube diameter increase when the tube pitch increase and lead to the cost increase.

5.3.7.

Baffling

Baffle The functions of baffles are supporting tubes, allowing a desirable velocity to be maintained for a shellside fluid and avoiding the failure of tubes, which is caused by the flow fluctuation. Baffles have two types, which are plate and rod. Plate baffles have single-segmental, double-segmental, and triple-segmental.

Figure 19 – Types of segmental baffe Baffle spacing

The centerline-to-centerline distance between adjacent baffles is the baffles spacing. The minimum baffle spacing is specified as one-fifth of the shell inside diameter (1/5 ID shell). The closer bundle will cause the insufficient bundle penetration by the shellside fluid and the difficult problem for mechanically outsides cleaning of the tubes. The maximum baffle spacing and shell inside diameter are of equal value. The predominantly longitudinal flow, which is less efficient than cross-flow, will exist when the baffle spacing is

50

higher. The higher baffle spacing also cause large unsupported tube spans and will lead to the tube failure of the exchanger due to flow-induced vibration. When decreasing the baffle spacing, the increasing rate of pressure drop is much faster than the increasing rate of heat-transfer coefficient. An optimum ratio of baffle spacing is commonly between 0.3 and 0.6 (0.3 < Lb > 0.6). This optimum ratio will lead to the highest efficiency of conversion of pressure drop to heat transfer. Decrease the baffle spacing can increase the pressure drop without a corresponding increase the heat-transfer coefficient. The resistance and the pressure drop of the main cross-flow path increase when the baffle spacing decrease. The leakage and bypass streams also increase until the balance of the pressure drops of all the streams is reached, due to the pressure drops of all five streams must be equal. Baffle cut The height of segment, which is cut in each baffle, is called baffle cut. The baffle cut permits the shellside fluid to flow across the baffle and is indicated as a percentage of the shell inside diameter. For shell-and-tube heat exchangers design, the baffles cut is one of important parameters and have less profound effect than the baffle spacing. The baffle cut can change between 15% and 45% of the diameter of shell inside. However, the baffle cuts between 20% and 35% is strongly recommended for using. The baffle cut below 20% lead to the increase of the shellside heat-transfer coefficient and the baffle cut beyond 35% lead to the decrease of the shellside pressure drop. Both of them usually lead to poor design. Reducing pressure drop by modifying baffle design In particular increasing in pressure drop across the heat exchanger if applying the same single pass shell and single segmental baffles, it may reduce such a significant scenario if handling other parameters related to tube and shell geometrical configurations.

51

5.3.8.

Shellside stream analysis

All five streams are flow parallel along paths of varying hydraulic resistance. Therefore, all streams begin and end at the inlet and outlet nozzles, and lead to the identical pressure drop of each stream. Because of the firmly dependence on the path resistances of the flow fraction, the stream analysis and the shellside performance will be influenced by varying any of the following construction parameters: • baffle spacing and baffle cut; • tube layout angle and tube pitch; • number of lanes in the flow direction and lane width; • clearance between the tube and the baffle hole; • clearance between the shell I.D. and the baffle; and • location of sealing strips andsealing rods.

5.3.9. Mean temperature difference Countercurrent flow is known as hot and cold stream flow in opposing directions across a tube wall. And cocurrent flow is known as hot and cold streams flow in the same direction.

Figure 20 – Countercurrent flow and Cocurrent flow

The countercurrent flow is usually preferred to concurrent flow due to the existing temperature cross, which is the outlet temperature of the cold stream is higher than the outlet temperature of the hot stream. A correction factor, Ft depends on the four terminal temperature and the shell style. The correction factor can be determined graphically provided the shell and tube configuration is known.

52

However, the overall heat transfer coefficient along the length of the shell is not influenced by the concept assumption of LMTD and Ft factor 5.3.10. Temperature profile distortion In general, the shellside stream is the cold fluid and the tubeside stream is the hot fluid. In this case, the temperature between the hot and the cold streams will be lower all along the length of the heat exchanger. Therefore, the mean temperature difference reduce and it is known as the temperature profile distortion (or correction) factor. The temperature profile distortion factor is important in many situation, the baffles are packed as close as possible to get the maximum shellside heat-transfer coefficient, pressure drop permitting.

5.3.11. Relevant Theory Heat duty: 𝑄 = 𝑚𝑐 (ℎ𝑜𝑢𝑡 − ℎ𝑖𝑛 ) Where 𝑚𝑐

: Mass flow rate of cold fluid (kg/h)

ℎ𝑜𝑢𝑡

: Specific enthalpy of input stream (kJ/kg)

ℎ𝑖𝑛

: Specific enthalpy of output steam (kJ/kg)

When the changes in the stream temperatures with distance through the heat exchanger (or with stream enthalpy) are linear, then the ∆𝑇𝑀 is a function only of the driving forces at the two ends of the heat exchanger, ∆𝑇2 𝑎𝑛𝑑 ∆𝑇1. Thus, log-mean temperature different is determined as: ∆𝑇𝐿𝑀 =

∆𝑇2 − ∆𝑇1 ∆𝑇 𝑙𝑛 ∆𝑇2 1

Where: ∆𝑇1 = 𝑇ℎ,1 − 𝑇𝑐,1 ∆𝑇2 = 𝑇ℎ,2 − 𝑇𝑐,2 The rate of heat transfer between two streams flowing through a heat exchanger is administered by: 𝑄 = 𝑈𝐴∆𝑇𝑚 Where Q: energy flow (W) U: overall heat-transfer coefficient (Wm-2 K-2) ∆𝑇𝑚 : mean temperature driving force (K)

53

The surface area of heat exchanger can be estimated by rearranging the equation: 𝐴=

Q 𝑈∆𝑇𝑚

The number of tubes can be calculated from: 𝐴 = 𝜋 × 𝑑𝑜 × 𝑁𝑡 × 𝐿 Where 𝑑𝑜 : Tube outside diameter 𝑁𝑡 : Number of tubes 𝐿 : Tube length Then rearranging the equation: 𝑁𝑡 =

𝐴 𝜋𝑑𝑜 𝐿

F correction factor The overall mean temperature difference ∆𝑇𝑚 is influenced by the multiple direction changes of two fluids. A correction factor, F, rises when the resulting ∆𝑇𝑚 , based on counter-current flow, is less than the log-mean temperature different is determined , ∆𝑇𝐿𝑀 .

𝑅=

𝑇ℎ𝑜𝑡,𝑖𝑛 − 𝑇ℎ𝑜𝑡,𝑜𝑢𝑡 𝑇𝑐𝑜𝑙𝑑,𝑜𝑢𝑡 − 𝑇𝑐𝑜𝑙𝑑,𝑖𝑛

𝑆=

𝑇𝑐𝑜𝑙𝑑,𝑜𝑢𝑡 − 𝑇𝑐𝑜𝑙𝑑,𝑖𝑛 𝑇ℎ𝑜𝑡,𝑖𝑛 − 𝑇ℎ𝑜𝑡,𝑜𝑢𝑡

Shell diameter, Ds 𝑃 ( 𝑡 )2 𝑑𝑜 𝐶𝐿 𝑑 ] 𝐷𝑠 = 0.637 × √ × 𝐴[ 𝑜 𝐶𝑇𝑃 𝐿 Where: CTP: tube constant for the incomplete coverage of the shell diameter by the tube CL : tube layout constant Having number of tubes and inside diameter, the area heat transfer can be calculated from the equation: 𝜋 × 𝑑𝑖2 𝐴𝑠 = 𝑁𝑡 × 4 The velocity of fluid in each tube pass or in shell is calculated by using the equation: 𝑚 𝑣= 𝜌 × 𝐴𝑠

54

Equivalent diameter can be determined as: 𝑑𝑒 =

1.10 2 (𝑝𝑡 − 0.917𝑑𝑜2 ) 𝑑𝑜

The Reynolds number can be obtained from: 𝑅𝑒 =

𝑢𝑡 × 𝑑𝑒 × 𝜌 𝜇

The pressure drops of tube side and shell side: 𝐿 𝜇 𝜌𝑢𝑡2 ∆𝑃𝑡 = 𝑁𝑝 [8𝑗𝑓 ( ) ( ) + 2.5] 𝑑𝑖 𝜇𝑤 2 ∆𝑃𝑠 = 8𝑗𝑓 (

𝐷𝑠 𝐿 𝜌𝑢𝑡2 𝜇 −0.14 )( )( )( ) 𝑑𝑒 𝑙 𝐵 2 𝜇𝑤

55

6. Summary process optimization 6.1.Reactor optimization 6.1.1.

Reactor Performance - Isothermal Temp

The purpose of this case study is to evaluate the effect of isothermal temperature condition on the reactor performance. In term of variable selection, conversion of main reaction and master component molar flow of DIPB (kmol/h), master component molar flow of cumene (kmol/h), were chosen as dependent variables. The temperature of inlet stream and the outlet stream will be defined as independent variables. The case study is run with initial temperature is set from 200oC to 500oC with step size of 20. This range is applied for both temperature of inlet and outlet stream. Base on the result of case studies, the result run with the same inlet and outlet temperature were chosen to make Figure 22 illustrates the impact of isothermal temperature on the overall performance and the main conversion rate

Figure 21 – The effect of isothermal temperature on reactor conversion and main production fraction

56

6.1.2.

Reactor Performance - Benzene/ Propylene Ratio

It is important to conduct the contribution of Benzene/ Propylene ratio on the conversion rate and selectivity. The master molar flow of cumene, DIPB, conversion of main reaction are dependent variables while independent variables is master component molar flow of benzene. The range of running is set to run from 114.1 to 1141 with 37 steps and step size of 28.5. The results were showed in the Figure 23.

Figure 22 – The impact of Benzene/Propylene ratio on conversion and main production fraction It is clear that when the Benzene/Propylene ratio run from 1 to 3, the percentage of conversion and cumene fraction in overall product have considerable growth from 91.47 to 99.76 %. At the Benzene/Propylene ratio is 1.4, the interception of conversion and cumene fraction are 95.88% and 97%. Therefore, the optimal value of Benzene/Propylene ratio is 1.4.

57

6.1.3.

Reactor Heat Transfer - Non-Isothermal Temp.

The purpose of this case study is to assess the influence of outlet temperature condition on the reactor performance when the inlet temperature fixed at 500 oC. In term of variable selection, conversion of main reaction and master component molar flow of DIBP (kmol/h), master component molar flow of cumene (kmol/h), were chosen as dependent variables. The temperature of the outlet stream will be defined as independent variables. The case study is run with initial temperature is set from 500oC to 1000oC with step size of 20. The results were shown in the Figure 24.

Figure 23 – Conversion and main production fraction against outlet temperature

58

6.1.4.

Reactor Configuration - Fluidized Bed Simulation

A case study is undertaken with a fluidized bed to evaluate the effect of 10% bypass on the overall process optimization achievement. The result needs to be compared with the process without bypass. A Tee equipment is uses to split the inlet stream into the reactor with 10% reactant. A new base case is also presented in Figure 25:

Figure 24 – Cumene production with fluidized bed simulation Base on HYSYS simulation results, it is obvious to make a comparison between the conversation rates of the outlet stream leaving the reactor in two case and then evaluate the effect of the bypass stream on the product requirements.

59

Figure 25 – Cumene molar mole of the outlet stream without 10% bypass stream

Figure 26 – Cumene molar mole of the outlet stream with 10% bypass stream Base on the PFD, it can be clearly seen that an extra 10% bypass stream leads to a poor conversion of the main reaction. As show in the Figure 26 and Figure 27 the molar flow of cumene without bypass is 106.15 kgmole/h while with the bypass it is 78.53 kgmole/h. It also means that there is a larger amout of propylene does not react with benzene. As a conclusion, the bypass stream is not useful for cumene production process.

60

6.1.5.

Additional Reactor - Transalkylation Reactor

To optimize the economic profit from the plant, it is required to decrease the side product DIPB or convert the side product DIBP back to the cumene. Therefore, an additional reactor known as the translkylation reactor is taken into account. Economic profit os one of the important criteria for assessing the feasibility of a project. Therefore, side products such as DIPB need to be used properly. One of the possible innovations is to convert DIPB back to cumene through a secondary reactor that performs translkylation a reaction: C6H6 Benzene

+

C12H18

→

DIPB

2 C9H12 Cumene

Nitin Kaistha (2011) has summarized the kinetic information for the transalkylation reaction: Rf = 2.529 × 108 exp(−100 000/RT)XBXD Reaction rate units for Rf kmol·m−3·s−1 while XB and XD is expressed for benzene and DIBP mole fraction respectively. From the equation, it is obvious to define the constant A is.529x108 m3/ (kmol.s) and activation energy is 100000 kJ/kmol, which is necessary for HYSYS simulation operation. The model the thermodynamic properties used is also Peng−Robinson equation of state.

Figure 27 – The parameters for reaction set

61

Due to the requirement of plant design, the operating condition such temperature is set by 200 oC while the pressure of 12 bar is specified. In order to achieve the conversion of 90%, the tube volume of 2.135 m3 and the diameter is 1.2 m. The outlet stream from distillation column T100 will be separated by a Tee and then a feed stream from this separation and DIBP from the final product stream would be fed for transalkylation reactor. Finally, extra cumene product goes through distillation column T-100 for recovering the benzene and cumene again. The detail flow diagram is structured as the Figure 29:

Figure 28 – PFD for cumene manufacturing with transalkylation reactor Base on the HYSYS simulation results and summarization in Table 20, although the conversion rate of cumene decreases, the total product can be recovered is higher than the initial process (112.4165 and 105.06 kmol/h respectively). As a result, transalkylation reactor is potentially recommended for cumene plant design. Table 20 – Comparison on Cumene performance between originated process and with additional reactor

62

6.1.6.

Raw Material – Propylene

Finding an appropriate component for the feed materials is important for a reliable and economical manner. There are two approaches including a pure propylene feed (higher than 99% purity) and impurity propylene feed with 5% propane. In this plant, both two feed streams were simulated with HYSYS software to evaluate the results which associate to overall performance (conversion for both main and side reaction) as well as economic possibility.

The Table 21 show the advantage of utilizing the impurity stream compares to the purer one. It is obvious that the product molar flow rate from impurity feed stream is much higher than the molar flow rate from the pure propylene feed stream (105.13 kmol/h and 100.2 kmol/h). More importantly, the cost of raw material for propylene feed with 5% propane is also acceptable. This is because the cost of more than 99% purity is $1570 which is much higher than just $880 from using of propylene with 5% propane (The University of Adelaide 2018). Another advantage of feed stream relates to the unreacted reactants. The propane which is recovered from the separator can be used or sold as fuel gas. The flow rate production of fuel gas is 283 kg/h and the cost for fuel gas is 630 per ton, which means that the extra profit from the design plant is $1,412,057 per year. Additionally, the desirable minimum molar flow rate cannot be achieved by the purity feed stream as well as increase the raw material cost, capital cost and utility cost relate to waste treatment. For more detail, the Section - economic appraisal will discusses this issue. As a result, the cumene plant design using propylene feed with 5% propane is more likely to be profitable and would achieve the product flow rate require Table 21 – Comparison on Cumene performance between two pathways Process with higher propylene feed stream transalkylation reactor

Process with propylene feed stream contains 5% propane reactor

Total cumene molar flow in the final product stream (kmol/h)

100.2

105.13

Conversion rate for side product (%)

6.097

3.241

Conversion rate for main product (%) Cumene purity in product stream (%)

93.9

95.88

99

99.93

63

6.2.Distillation Optimization: Table 22 – Input parameters of distillation column Parameters

Column T-101

Inlet temperature

190.4 [oC]

Condenser & Reboiler pressure

206.8 & 243.8 [kPa]

Summary specs Reflux ratio

0.66

Temperature of Reboiler

252.2 [oC]

Figure 29 – Hydraulic plots of column T-101

64

6.2.1. Column T-101 – Column temperature Case study is run for investigating the effective of temperature on the separation efficiency and duties of reboiler and condenser. In the variable selection, the independent variable is the feed temperature, whereas purity of cumene produce and condenser/reboiler duties are the dependent variable. Mass fraction of cumene in distillate reflects separation efficiency of the unit. The feed stream inlet temperature is in the range of 25 oC to 250 oC. The temperature of reboiler are stable at 252.2 OC and the reflux ratio are kept constant at 0.6603. Variables are plotted and the present of the unit can be analyzed from the obtained results. 0.9976

Mole fraction of Cumene

0.9974 0.9972 0.997 0.9968 0.9966 0.9964 0.9962 0.996 0.9958 0.9956 0.9954 0

50

100

150

200

250

Temperature (oC) Mole fraction of cumene on top

Figure 30 – A plot of feed temperature and cumene mole fraction

-6195000

10000000

-6200000

8000000 -6205000

6000000 -6210000 4000000 -6215000

2000000 0

-6220000 0

50

100

150

200

250

Temperature (oC) reboiler duties

condenser duties

Figure 31 – A plot of feed temperature and reboiler/condenser duties

Energy consumed for condener (KJ/h)

Energy consumed for reboiler (KJ/h)

12000000

65

In overall, the changing operation of the feed temperature affects to the purities of cumene product, while the energy consumption for reboiler and condenser are impacted. The figure 31 provided the temperature of inlet stream plotted with the purity of cumen product. It is observed that the value of cumene purities decrease slightly from 99.746% at 25 oC to 99.721% at 190 o

C, and then decline rapidly reach 99.568% purity of cumene in 250oC. It can be explained that

the increase of temperature lead to the increase of pressure in column and the less volatile components. As a result, the purity of the distillate may deteriorate. Furthermore, a plot of between feed temperature and the duty of reboiler and condense are presented in figure 32. The change of feed temperature has not considerable impact on condenser duty, which means that increasing nearly 20000 KJ/h from 25oC to 250 OC for the energy consumed in condenser is not negligible. Nevertheless, the duty of reboiler is impacted dramatically by the alterative of feed temperature. The consumption of energy in reboiler reduce sightly in the range of 25 oC to 190 oC, then sudden decrease over a half of energy consumption from 190oC to 193 oC (drop from 6.805016.e +006 to 2.191186.e+006). => From this case study, 190oC is suggestion inlet temperature for column to achieve the high purity of product.

66

6.2.2. Column T-102 – Column pressure In this case study, the column pressure is setting as the independent variable and the mass fraction of cumene, condenser/reboiler duties are dependent variable. The condenser pressure is in the range of 1 bar – 12 bar, which approximately 100kPa to 1200 kPa. The aim of this case study is to determine the performance of column based on pressure has been conducted. Variables are plotted and the present of the unit can be analyzed from the obtained results. 0.9975

Cumene mass fraction

0.997 0.9965 0.996 0.9955 0.995 0.9945

0.994 0.9935 0

200

400

600

800

1000

1200

1400

Condenser pressure (kPa)

Figure 32 – A plot of condenser pressure and cumene mole fraction 9000000

Energy consumped (kJ/h)

8000000 7000000 6000000 5000000 4000000 3000000 2000000 1000000 0 0

200

400

600

800

1000

1200

Condenser pressure (kPa) Condenser duty

Reboiler duty

Figure 33 – A plot of condenser pressure and condenser/reboiler duties

1400

67

Generally, the change in condenser pressure have effect on both the mass fraction of cumene and the condenser/reboiler duties. Regarding Figure 33 a plot of condenser pressure with mass fraction of cumene, there is a fluctuated from 100 kPa (1 bar) to 300 kPa (3 bar), peak a top at 99.96% (at 300 kPa) cumene purity, and then the percentage of cumene purity start to decline at the pressure of 300kPa. It can be explained that while increasing column pressure, the volatility of component also decreases, and results in decrease in separation efficiency. To achieve the desired purity of product with the high pressure, the more stages and the rise value of reflux ratio is required. Furthermore, the duties of condenser and reboiler are considered with the alternative column pressure from figure 34. The energy consumed in condenser reduces gradually, while reboiler duty growth steady from 2.233273.e+006 kJ/hr to 6.419022.e+006 kJ/hr in the range of 200kPa to 300kPa, then moderately increase. A possible explanation for that the risen of column pressure lead to the increase of the bubble point of mixture in column, so there is more requirement for reboiler duty. In case of more energy consumption for reboiler, the more operation cost is needed. => Thereby, from this case study, the pressure is recommended to run this process is approximately 233.5 kPa for best preformation of separation, saving energy and cost operation.

68

6.2.3.

Column T-101 – Reflux ratio

In this case study, the independence variable is reflux ratio with the value in the range of 0.1 – 1.0, when the efficiency of separation and duties of condenser/reboiler are consistently specified. The purpose of this case study is investigating on how the alternative reflux ratio in separation can affect to the efficiency of column and energy consumed of condenser/reboiler. Variables are plotted and the present of the unit can be analyzed from the obtained results. 0.9975

0.9965 0.996 0.9955 0.995 0.9945 0.994 0.9935 0

200

400

600

800

1000

1200

1400

Condenser pressure (kPa)

Figure 34 – A plot of condenser pressure and cumene mole fraction 9000000 8000000

Energy consumped (kJ/h)

Cumene mass fraction

0.997

7000000 6000000 5000000 4000000 3000000 2000000 1000000 0 0

200

400

600

800

1000

1200

Condenser pressure (kPa) Condenser duty

Reboiler duty

Figure 35 – A plot of condenser pressure and condenser/reboiler duties

1400

69

In general, the alteration reflux ratio has impact on product purity and energy consumed in condenser and reboiler. The mole fraction of cumene increases rapidly from 98.9%, and reach at 99.767 %, while reflux ratio equal to 1. As shown in figure 35, the higher reflux ratio greatly increases the purities of cumene, because while the higher number of reflux ratio, the less tray for separation is needed. Therefore, the amount of light key product in distillation increase. However, there are some withdraw if the value of reflux ratio is too high. First, the velocity of vapour through the tray rise and lead to flooding. Moreover, there will back-mixing of components between the lower and upper trays in column which decrease separation. Furthermore, it is clearly to see from figure 36 that the risen of reflux ration lead to more required energy for reboiler and decrease the energy consumed in condenser. The reboiler load increase to restore the set point temperature, so it causes to rise more operation cost. Therefore, the value of reflux ratio is chosen in this project is 0.66 which not only can achieve the high percentage of purities of cumene product, but also available with operation cost from consumed energy of reboiler.

70

6.3. Heat exchanger

Figure 36 – Expected temperature of inlet and outlet of tube and shell (Amos. S., University of Adelaide) 6.3.1. Input parameters in HYSYS 6.3.1.1. Theoretical heat exchanger Table 23 – Assumption and calculation all parameters in appendix (x) Parameters TEME type Allowable pressure drop

Location Values Overall heat exchanger A E Shell 140 [kPa ] Tube 140 [kPa] 1 1 1

Convert units L

Number of Shell Passes Number of Shells in Series Number of Shells in Parallel Tube Passes per Shell 1 Tube geometrical configuration Outside diameter ¾ [in] Length 20 [ft] Thickness 18 [BWG] Inside diameter 16.56 [mm] Fouling coefficient 0.000001 [C-h-m2/kJ] Shell geometrical configuration Inside diameter 322.3 [mm] Number of tubes 104 Tube Pitch 28.575 [mm] Tube Layout Angle Triangular (30 degrees) Fouling coefficient 0.000001 [C-h-m2/kJ] Shell Baffle Type Single Shell Baffle Orientation Horizontal Baffle Cut (%Height) [%] 25% Baffle Spacing [mm] 64 [mm]

-

19.05 [mm] 6.096 [m] 1.245 [mm] -

-

71

6.3.1.2. Practical heat exchanger Table 24 – Auto sizing all parameters in the Rigorous model Parameters TEME type Allowable pressure drop

Location Values Overall heat exchanger B E Shell 0.5-0.7 Kg/cm2 Tube 0.05-0.2 Kg/cm2 1 1 1

Convert units M

Number of Shell Passes Number of Shells in Series Number of Shells in Parallel Tube Passes per Shell 1 Tube geometrical configuration Outside diameter 19.05 [mm] Length 1.8 [m] Thickness 2.108 [mm] Inside diameter 14.834 [mm] Fouling coefficient 0.000001 [C-h-m2/kJ] Shell geometrical configuration 438.15 [mm] Number of tubes 203 Tube Pitch 23.81 [mm] Tube Layout Angle Triangular (30 degrees) Fouling coefficient 0.000001 [C-hm2/kJ] Shell Baffle Type Single Shell Baffle Orientation Horizontal Baffle Cut (%Height) [%] 25% Baffle Spacing [mm] 210 [mm] Inside diameter

68.65 kPa 19.6 kPa

-

-

-

72

6.3.2. Graph of heat exchanger performance - Fouling of tube 6.3.2.1. Tube fouling on overall U 2600

Overall U [kJ/h-m2-C]

2400 2200 2000 1800 1600 1400 1200 1000 0

0.00005

0.0001

0.00015

0.0002

0.00025

0.0003

0.00035

fouling coefficient [C-h-m2/kJ] theoretical tube fouling

practical tube fouling

Figure 37 – Tube fouling against overall U in theoretical and experimental heat exchanger 6.3.2.2. Tube fouling on pressure drop of shell and tube

197 196.5 196 195.5 195 0

0.0001

0.0002

0.0003

fouling coefficient [C-h-m2/kJ] shell pressure drop

tube pressure drop

pressure drop [kPa]

pressure drop [kPa]

197.5

5.15 5.1 5.05 5 4.95 4.9 4.85 4.8 0.0004

pressure drop [kPa]

198

5

0.2

4

0.15

3

0.1

2

0.05

1 0 0

0.0001

0.0002

0.0003

0 0.0004

fouling coefficient [C-h-m2/kJ] shell pressure drop

tube pressure drop

Figure 38 – Fouling coefficient of tube in theoretical and practical heat exchanger

pressure drop [kPa]

practical HEX

theoretical HEX

73

6.3.3. Graph of heat exchanger performance – Fouling of shell 6.3.3.1. Shell fouling on overall U 2400

Overall U [kJ/h-m2-C]

2200 2000 1800 1600 1400 1200 1000 0

0.00005

0.0001

0.00015

0.0002

0.00025

0.0003

0.00035

fouling coefficient [C-h-m2/kJ] theoretical shell fouling

practical shell fouling

Figure 39 – Shell fouling against overall U in theoretical and experimental heat exchanger

6.3.3.2 Shell fouling on pressure drop of shell and tube

197.5

5.05

197

5

196.5

4.95

196

4.9

195.5 0

0.0001

0.0002

0.0003

4.85 0.0004

fouling coefficient [C-h-m2/kJ] shell pressure drop

tube pressure drop

pressure drop [kPa]

5.1

pressure drop [kPa]

pressure drop [kPa]

198

5

0.2

4

0.15

3

0.1

2

0.05

1 0 0

0.0001

0.0002

0.0003

0 0.0004

fouling coefficient [C-h-m2/kJ] shell pressure drop

tube pressure drop

Figure 40 – Fouling coefficient of shell in theoretical and practical heat exchanger

pressure drop [kPa]

practical HEX

theoretical HEX

74

6.3.4. Graph of heat exchanger performance – Number of baffle segments 6.3.4.1. Number of baffle segments on overall U 1572.5

1999.4276

1571.5

Overall U [kJ/h-m2-C]

Overall U [kJ/h-m2-C]

1572

1571 1570.5 1570 1569.5 1569 1568.5 0

200

400

600

800

1999.4276 1200

1000

baffle spacing [mm] theoretical baffle spacing

practial baffle spacing

Figure 41 – Baffle spacing against overall U in theoretical and experimental heat exchanger 6.3.4.2 Number of baffle segments on pressure drop of tube and shell

5.012 5.0118 5.0116 5.0114 5.0112 5.011 5.0108

400 300 200 100 0 0

50

100

150

200

250

300

350

baffle spacing [mm] shell presure drop

tube pressure drop

30 25 20 15 10 5 0

0.2 0.15

0.1 0.05 0

200

400

600

800

1000

0 1200

baffle spacing [mm] shell pressure drop

Figure 42 – Baffle spacing in theoretical and practical heat exchanger

tube pressure drop

pressure drop [kPa]

500

pressure drop [kPa]

practical HEX pressure drop [kPa]

pressure drop [kPa]

theoretical HEX

75

6.3.5. Analysis graph The increase in fouling coefficient, in which increases viscosity resistance of fluid to influence the heat transfer scenario that generally makes the drop in overall heat transfer coefficient. However, with the rise in fouling coefficient in range between 0 to 0.0003 [C-h-m2/kJ] of tube and shell, the overall heat transfer coefficient decreases in the theoretical heat exchanger and remains constantly in the practical heat exchanger. Due to the stable value of overall heat transfer coefficient, the pressure drops of both tube and shell, 0.1467 and 4.04 are too low and the heat transfer per shell is not sufficiently high, 21.87 [m2], plus one tube pass per shell within the practical heat exchanger. In theory, the pressure drop of the tube increases while the fouling coefficient increases to 0.0003 [C-h-m2/kJ]. Decrease parameters from default to real values, including shell diameter (Ds), 738 [mm] to 322 [mm]; tube diameter (Dt), from 20 [mm] to 19.05 [mm], number of tubes, (Nt), from 160 to 104, thereby increases the flow velocity, hence, increases the tube side pressure drop. In contrast, the shell pressure drop declines along with the same fouling coefficient. Increasing the tube pitch ratio. Increase in tube pitch ratio reduces the cross flow velocity and, thereby, reduces the pressure drop. Moreover, the tube volume is 2 times smaller than the shell volume in each shell, 0.1155 and 0.2728 [m3], respectively. Moreover, the molar flowrate of cold fluid is 2 times higher than the molar flowrate of hot fluid, 280.1 and 166.5 [kgmole/h], respectively. In practical, the pressure drops of both shell and tube are constant running at the same value of fouling coefficient, because all parameters is not responsive to fouling coefficient of tube and shell. In the theoretical heat exchanger, the overall heat coefficient increases when the baffle spacing increases, thus the increase in velocity of fluid with more interaction between cold and hot fluid following the growth in overall heat transfer coefficient. However, the overall heat coefficient of the practical heat exchanger still remains constantly in spite of the increase of baffle spacing. When the diameter of baffle spacing increase, the velocity of both cold and hot fluid declines leading to the downward trend in the tube and shell of the theoretical heat exchanger. In the practical heat exchanger, with the rise in of baffle spacing, the pressure drop of the shell decreases while the pressure drop of the tube is unchanged due to the geometrical sizes of tube and shell.

76

6.3.6.

Result Table 25 – Theoretical and experimental results of heat exchanger

Temperature hot fluid [oC] Temperature cold fluid [oC] Pressure hot fluid [kPa] Pressure cold fluid [kPa] Overall U [kJ/h-m2-C]

Specified Pressure Drop [kPa]

Theoretical results Inlet Outlet 450 258.1

Experimental results Inlet Outlet 450 275.1

27.96

222.3

27.96

200

2492

2487

2492

2492

2500

2304

2500

2496

1571

Shell 196.4

Tube 5.011

1999

Shell

Tube 3.841

0.147

77

7. Economic appraisal A comparison on the economic manner is taken into account to evaluate the profitability of the plant between two pathways. The first pathway is the feed stream contain propylene raw material with more than 99% purity and the second one is propylene stream with 5% propane. In term of economic evaluations, the main factors are considered focus on the capital cost, Net Present Value (NPV), payback period and expenditure. However, the summary information provided by Amos (2018) has been used in 2011. Therefore, the chemical engineering plant cost index (CEPCI) will be used to convert this value into the current value for achieving more accurate cost estimation. According to Chemical Chemical Engineering Magazine (update in January 2018), the CEPCI was 585.7 in 2011 and 576.4 is the value of CEPCI in 2018. The equation for converting have been show below: 𝐶𝐸𝑃𝐶𝐼 (2018)

Current cost = Purchase cost * 𝐶𝐸𝑃𝐶𝐼 (2011) . The following table will summarize the various costs associated with plant operation for two pathways. Table 26 – Summary of costs Summary of Costs % of FCI 5% propane impurity feed > 99% propylene 100% $3,568,768 $3,821,281 4% $142,750.73 $152,851.24 7% $249,813.77 $267,489.68 $3,961,333 $4,241,622 Utilities Power $ 8,430.58 $ 9,926.78 Cooling water $ 183,077.78 $ 224,530.32 LP steam $ 127,391.65 $ 159,738.48 DIPB waste treatment $ 4,635,965.31 $ 9,077,142.15 Total $ 4,954,865.32 $ 9,471,337.73 Labour Costs % $510,400 Operating labour cost 100% $102,080 Employee oncost 20% $586,960 Operating supervision 115% $586,960 Laboratory labour 115% Total $1,786,400 Raw Materials and Profits Material cost Benzene $79,123,968 $79,123,968 Propylene with 5% propane impurity $35,287,085 $62,955,367 Catalyst $42,336 $42,336 Total $114,453,389 $142,121,671 Profit Cumene $143,499,233 $143,885,380 Fuel gas $1,412,057 $758,419 Total $144,911,290 $144,643,800 Miscellaneous Costs 15% Contigencies Capex $4,246,834 $4,547,325 Opex $121,444,468 $153,646,899 Savings $23,466,822 -$9,003,099 Total Capital Investment $125,691,302 $158,194,223 Fixed Capital Fixed Capital Investment Land Maintenance and Repairs Total

78

7.1.

Capital cost

Base on the preliminary plant design is simulated on HYSYS, a calculation was undertaken to evalue the costs in both initial pathways. The object of the study is to focus on the subsidiary costs include the capital cost investment, land and contigences. The error related to this calculation have higher value which oscillates between ± 20% - ± 30%. In order to improve the accuracy of economic estimation, it is suggested that the client should specify the definitive cost before the plan is sanctioned. It can increase the accuracy into ± 2% - ± 5% of the actual capital cost (Peter, 2018, University of Adelaide). Table 27 – Cost study of project expenditure

Aspen HYSYS have been used to simulate the cumene production process which achieves 100.000 tons per year with raw material contain benzene and propylene. For developing a fully capital cost, the purchasing cost for all equipment needs to be considered firstly. A detail and sufficient information of equipment cost were given in the Appendix Economic Evaluation (Amos 2018). All estimation will base on the unit sizing, pressure and power. The equation used to calculate the total installed costs was: Total installed cost = Purchased cost *(4 + Pressure factor + Material factor) Table 28 – Material factors associated with different materials (Amos 2018)

Material Carbon steel Stainless steel

Material factor 0 4

79

Table 29 – Pressure factors associated with different pressures (Amos 2018)

Pressure < 10 atm 10-20 atm 20-40 atm 40-50 atm 50-100 atm

Pressure factor 0 0.6 3.0 5.0 10

According to Amos (2018), the cost of valves and pipes has already estimated in the apparatus cost factor, hence there is no costing for valves and pipe included. Moreover, cost for catalyst, trays and vessels have been included the pressure so the pressure factor will be 0 for those unit’s calculation. Also, it is the same for cost of reboiler and condenser which were assumed as a part of cost for purchasing the distillation column. There are two mixers in the PFD, however, the fee for that equipment is very small compares to cost for other units like distillation column or reactor. A full calculation of working capital and installation costs will be presented in Appendix 4

7.2. Chemical engineering plant cost index (CEPCI) The given cost in the project is in the year of 2011, therefore it is important to convert the equipment cost from 2011 into 2018 by appling the CEPCI index. CEPCI is an inflation rate index which aims to perfom a better or more accurate estimation for chemical plant operation. The CEPCI index for 2011 and 2018 are 585.7 and 576.4 respectively (chemengonline.com.au update at January 2018).

7.3. Operating expenses (OPEX) OPEX is an index indicates the annual cost to run a process plant. The estimation was calculated following the gross profit information provided by Amos (2018) and Peter textbook. OPEX estimation is the summary cost of production, fixed charges, general expenses and plant overhead costs

80

7.4. Equipment sizing and costs Base on the design brief and PFD base case for cumene manufacture, the equipment for both pathways are the same. Main equipment required in PFD contain: a furnace, a reactor, a separator, a heat exchanger and two coolers.

7.5.Material of Construction There are two main materials are considered for equipment containing stainless and carbon steel. Base on the equipment properties and operating condition, potential materials were chosen for each unit. A summary of material will be shown in Appendix following applied cost factors to calculate the installation costs.

7.6. Carbon Steel and Stainless Steel Carbon steel is widely used in the chemical industries because of various advantages for the manufacturing process. Beside the availability of difference size, it is considered as a potential material for organic chemical production (Peters and Timmehaus, 1991). In this plant, almost substances are organic as well as carbon steel is cheaper than the stainless steel, therefore carbon steel is used for majority of operating equipment. However, it is suitable for equipment with operating temperature is under 500oC. Stainless steel is other important material which is widely applied for many chemical plants. It is often used to treat the chemical substance cause the corrosion and has a much higher maximum temperature range from 600oC to 1150oC, which depend on the special alloy ingredient of stainless steel (Seider et al. 2009). The most important component in stainless steel is chromium which has the minimum content of 12 %. The more chromium percentage, the higher oxidizing agent resistance for stainless steel. However, this material is much more expensive than carbon steel. *Material of Construction consideration Since carbon steel is cheaper than stainless steel as well as appropriate for organic chemicals, it is widely applied for many pieces of apparatus. However, for main equipment such heat exchanger, reactor or tray sizing which is likely operated with temperature higher than 480 oC

81

7.7.Utilities Utility plays an important role to evaluate the possibility of the project through two pathways. Base on the price and type of utilities given in the design brief. There are three type of energy provide for equipment which include power, cooling water and lower pressure steam (LP). Moreover, DIBP is an unwanted product that needs to be disposed. It means that there is extra cost estimation for waste treatment. It can be clearly seen that the cost for treating the side product occupied the most utilities cost in case of both operation pathways.

Figure 43 – Summary cost for three different type of energy

Figure 44 – Operating cost for DIPB treatment and total utilities cost

82

All calculation and detail summary of the price for equipment utilities in this plant design will be discussed in Appendix 4. In general, it can be briefly indicated that the annual utility costs for the pure feed stream is more double than the annual utility cost spent by the impure feed stream. Therefore, in term of utility manner, the plant will earn more profit by using the feed stream contains 5 % propane.

7.8.Operating Labour Operating labour was determined following the spreadsheet provided by Amos (2018). All equipment need to be considered in HYSYS PFD contain one reactor, one heat exchanger, one vessel, two pumps and 2 distillation column. Total figure of operator required per shift is estimated to be 1.8. In order to fulfill the condition for operating labour cost calculation, various assumptions were presented. Firstly, working period for one operator is 49 weeks with 3 weeks for relaxing or sicking. Also, there are three working shift with 330 working days per year, each shift prolongs 8 hours and then total shift per year will be 245. As a result, the overall operator will be 7.27, which leads to the number of operators needs is around 8. The annual salaries for each worker were assumed to be $63,800 per year, hence the operating labour cost was equal to. In term of operating labour, CEPCI does not associated to operating labour calculation.

7.9. Raw materials and profits The raw materials and profits have a significant impact on the economic feasibility of the cumene plant. Pathway 1 with 5% propane impurity The mass flow rate of benzene and impurity propylene are 8920 kg/h and 5063 kg/h respectively. The operating period is 7920 hours. In order to calculate the annual quantity, the mass flow rate of feed stream is multiplied by operating time. And then the cost for raw profit can be found. Moreover, the fee for catalyst is also considered by void fraction, the volume of the reactor and catalyst density. There are two main products from this pathway. Firstly, the main product is cumene that has mass flow rate of 100,069,200 kg/year, which cost A$143,499,233/ year. The mass flow rate of propylene and propane was separated from separator is sell in price of A$1,4212,057/ year.

83

Table 30 – Various costs of raw materials required for plant operation with 5% propane impurity in fee

The overall profit can be earned from manufacturing process with Pathways 1 is A$144,911,290/ year Pathway 2 with more than 99% propylene purity The mass flow rate of benzene and impurity propylene are 8,920 kg/h and 5,063 kg/h respectively. The operating period is 7920 hours. In order to calculate the annual quantity, the mass flow rate of feed stream is multiplied by operating time. And then the cost for raw profit can be found. Moreover, the fee for catalyst is also considered by void fraction, the volume of the reactor and catalyst density. Table 31 – Various the costs of raw materials required for plant operation with 99% propylene purity in feed. Raw M aterials and Profits Raw M aterials

Quantity (kg/h)

Operating hours Quanlity per year (kg/year)

Price per kg

Annual Cost

Benzene

8,920

7920

70,646,400

1.12

$79,123,968

Propylene

5,063

7920

40,098,960

1.57

$62,955,367

16,800

2.52

$42,336

Catalyst Total Annual Raw M aterial Cost Profit M aterials

Quantity (kg/h)

$142,121,671

Operating hours Quanlity per year (kg/year)

Price per kg

Annual Profit

Cumene

12,669

7920

100,338,480

1.434

$143,885,380

Fuel gas

152

7920

1,203,840

0.63

$758,419

Total Annual Profits

$144,643,800

There are two main products from this pathway. Firstly, the main product is cumene that has mass flow rate of 100,338,480 kg/year, which cost A$142,885,380/ year. The mass flow rate of prolylene and propane was separated from separator is sell in price of A$758,419/ year. The overall profit can be earned from manufacturing process with Pathways 2 is A$144,643,800/year

84

Net Present Value (NPV) Couples of assumptions were classified by Amos (2018), the linear depreciation rate is 10 %, 7.10.

the after-tax internal hurdle rate was 9% p.a while the marginal taxation rate was 35%. In additionally, the construction period of the project is one year and the plant life is operated in 10 year. A cash flow statement was restructured by using above assumptions, which evaluates the return on investment, internal rate of return, especially, the net present value and payback period. It is the most important to estimate the Net Present Value (NPV) for both pathways. The meaning of NPV behavior for future cash flows that the project will earn in a long manufacturing plan. This means that if the NPV is negative, then the project should not be taken into account. In contrast, The NPV index is positive allows the project generally to be operated. Table 32 - Net Present Value

NPV PWPI Index Pre Tax Payback Atfter Tax Payback IRR

5% propane impurity feed stream Purer propylene feed stream $ 94,576,560.34 -$ 41,130,566.23 22.27 -9.05 0.18 -0.51 0.21 -0.58 363% N/A

Based on the NPV calculations, the impure plant should be accepted, and the pure case should be rejected. 7.11.

Payback Period

The payback period is the time the project recovers the original investment from the value of net cash flow is zero. From the calculation on cumulative cash flows in Appendix, the payback period is in the first year. Especially, if the payback period is negative, the plant seems to never gain the profit.

85

8. Conclusion The significant information of Dr. Who Chemicals Ltd that the overall result of the cumene production needs to reach the 105.1251 kgmole/hr with the purity 99.91 wt% in order to produce 100,000 metric ton per year. After the study plans including the physical and chemical properties of reactants and products in the operation process, furthermore, catalyst zeolite is chosen to simulate the cumene product in the PFR reactors due to higher profit probability and reduce energy requirement, thereby more economic than others catalyst. Not only that, benzene is the only concern in chemical hazard during the cumene operation because of its cancer potential disease, but also benzene causes damage genetic impact to both human and animal generation in environmental term. In term of economic assessment, an evaluation was carried out to compare the revenue that may be obtained from the high purity propylene feed contains 5% propane. The net present value, PWPI index as well as payback period and IRR have been calculated. Each measurement of these index proved that using the propylene feed stream contains 5% propane is an economically feasible project. In contrast, the economic indicators has shown that the using the purer feed stream seems to never bring any profit for the company during its life. 

Potential of error

Such assumption of absolute purity of input benzene and propylene as raw materials, thus, the mass and energy in the entire flow sheet are quite simpler than the reality. Moreover, the nonclear observation of safety operation can be seen easily during running the computer software ASPEN HYSYS V10.

86

9. References Amos, S, “HEX Design and Sizing”, CHEM ENG 3030, University of Adelaide, Adelaide, 20th August. Amos, S, “Column”, CHEM ENG 3030, University of Adelaide, Adelaide, 18 th August. Aspen Technology, Inc 2012, Aspen Icarus Reference Guide, 8th edn, Aspentech, viewed 18 October 2018, Chemical Engineering Magazine 2018, ‘CEPCI Junuary 2018 Issue’, view 21 Oct 2018, . Cuno. CW, 1929, “Economic factors in chemical plant location”, Industrial and engineering chemistry, Vol. 21, No.08, p. 739 Gera, V, Kaistha, N, Panahi, M and Skogestad, S (2011) ‘Plantwide Control of a Cumene Manufacture Process’, Indian Institute of Technology Kampur, Chemical Engineering Department [Accessed 10 October 2018] Available at:

Holman, JP 2010, “Heat exchanger”, Heat Transfer, McGraw-Hill, New York, p. 527. Luyben, W,2010, ‘Design and Control of the Cumene Process’, Ind. Eng. Chem. Res., vol 49, no. 2, pp. 719-734 Mangili, PV, Junqueira, PG, Santos, RO, Santos, LS and Prata, DM 2018, “Economic and environmental analysis of the cumene production process using computational simulation”, Chemical Engineering & Processing: Process Intensification, Vol. 130, p. 310. McCabe, WL, Smith JC and Harriott. P 1993, “Distillation”, Unit operation of chemical engineering, 5th edition, McGraw-Hill, pp.521-576.

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Mukherjee, R 1998, “Effectively design shell-and-tube heat exchangers”, Chemical engineering progress, vol. 94, no. 2. Niwa, M, Katada, N and Okumura, K 2010, “Introduction to Zeolite Science and Catalysis”, Characterization and design of zeolite catalysts, Springer, pp. 01 – 02. Norouzi,HR, Hasani, MA, Haddadi-Sisakht, B and Mostoufi, N 2014, “Economic design and optimization of Zeolite-Based Cumene production plant”, Chemical engineering communication, Vol. 201, pp. 1271 – 1291. Peter, M. Timmerhaus, K. & West, R. 2003, Plant Design and Economics for Chemical Engineers, University of Colorado, Mcgraw Hill. Perry, R.H. and Green, D.W. 1997, Perry’s Chemical Engineers’ Handbook, 7th Edition, McGraw-Hill. Treybal, RE 1980, “Distillation”, Mass transfer operation, 3rd Edition, McGraw-Hill, pp.363– 423.

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10. Bibliography Luyben, W 2011, “Design and control of the cumene process”, Principles and case studies of simultaneous design, 1st Edition, John Wiley & Sons. Pathak,A , Agarwal, S, Gera, V and Kaistha,N 2011, “Design and control of a vapourphase conventional process and reactive distillation process for cumene production”, Industrial & Engineering Chemistry Research, Vol.50, pp. 3312 – 3326.

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Appendix Appendix 1 – Calculation procedure of reactor The reaction kinetics is shown as the equation below: The reaction kinetics is shown as the equation below: ri=

ki cb cp mole (gcat)(s)

Where the equilibrium constant k expresses as: −E

Ki=Ai × exp( RTi ) For Rxn-1: k1 =3.5 × 104 × exp(

−24.9 RT

)

Where: L2

L2

3.5 × 104 (mol×g×cat×s) = 3.5 × 104 × mol×g×cat×s ×

1000g kg

1m3

× 1000L = 3.5 ×

m3 L

104 mol×kg×cat×s m3 L

3.5 × 104 mol×kg×cat×s ×

1000mol 1 kmol

×

1600kg.cat 1000L

m3

= 5.6× 107 kmol.s

The void fraction is 0.5: m3

A1= 5.6 × 107 × 0.5 = 2.8 × 107 kmol.s kcal

E1= 24.9 mol =

24.9×4.2×1000kJ kmol

= 104580 kJ/kmol

For Rxn-2: k2 =2.9 × 106 × exp(

−35.08 RT

)

Where: L2

L2

2.9 × 106 (mol×g×cat×s) = 2.9 × 106 × mol×g×cat×s ×

1000g kg

1m3

× 1000L

m3 L

= 2.9 × 106 mol×kg×cat×s m3 L

3.5 × 104 mol×kg×cat×s ×

1000mol 1 kmol

×

1600kg.cat 1000L

The void fraction is 0.5: m3

A1 = 4.64× 109 × 0.5 = 2.32 × 107 kmol.s kcal

E2 =35.08 mol =

35.08×4.2×1000kJ kmol

m3

= 4.64× 109 kmol.s

= 147336 kJ/kmol

90

Appendix 2 – Outline equipment

Figure 45 – PFD for cumene production Table 33 – Outline equipment Equipment Mixers

Pumps

Symbol MIX-100

Image

Description Mixing benzene feed stream and propylene feed stream

MIX-101

Mixing cumene produced from transalkylation reactor with separated liquid

MIX-102

Mixing the outlet stream leaving the PFR-100 with bypass stream

MIX-103

Mixing the benzene stream with DIPB for from trans alkylation reactor

MIX-104

Mixing the recycle benzene with the fresh benzene

P-100

75% efficiency ΔP = 2399 kPa Power = 12.40 kW

91

P-101

75% efficiency ΔP = 1357 kPa Power = 5.028 kW

P-102

75% efficiency ΔP = 1025 kPa Power = 0.1492 kW

P-103

75% efficiency ΔP = 990 kPa Power = 0.8892 kW

E-100

Q= 5,614,399.5 kJ/h

E-102

Q = 8,291,421.01 kJ/h

E-101

Q = 195,548.2 kJ/h

E-103

Q = 31,302.85 kJ/h

E-104

Q = 259,979.35 kJ/h

Furnace

FH-100

75% efficiency ΔP = 2399 kPa Q = 6,419,000 kJ/h

Tees

TEE-100

Split the benzene stream into the stream for transalkylation reactor

TEE-101

Split the initial feed stream before entering the PFR-100

Heat exchangers

92

Reactors

PFR-100

PFR-101

Total reactor volume = 21.02 m3 Length = 10.46 m Diameter = 1.6 m Number of tubes = 1 Wall thickness = 0.005 m Q = 6,746,666.14 kJ /h Total reactor volume = 2.135 m3 Length = 1.388 m Diameter = 1.2 m Number of tubes = 1 Wall thickness = 0.005 m Q = 0 kJ /h

Separator

V-100

ΔP = 0 kPa Vertical direction, flat cylinder

Distillation columns

T-100

Reflux ratio = 1 27 stages; main stage: 15 Weir height = 50.8 mm Weir length = 0.83 m Sieve tray, 24 inch spacing Column diameter = 4.3 ft Active area = 0.81 m2; total area = 0.92 m2 Q condenser = 2,283,152.4 kJ/h ; Q reboiler= 4,868,454.16 kJ/h

T-101

Reflux ratio = 0.66 37 stages; main stage: 7 Weir height = 50.8 mm Weir length = 0.87 m Sieve tray, 24 inch spacing Column diameter = 1.21 m Active area = 0.92 m2; total area = 1.035 m2 Q condenser = 6,011,241.72 kJ/h ; Q reboiler= 6,074,897.015 kJ/h

Valve-100

ΔP = 2389 kPa Percentage of open = 50%

Valve-101

ΔP = 73.68 kPa Percentage of open = 50%

Valve-102

ΔP = 1099 kPa Percentage of open = 50%

Valve

93

Appendix 3 – Calculation procedure for distillation column 1. Splitter Specify in parameter o Top stream

: vapour – vapour fraction

=1

o Bottom steam

: liquid – vapour fraction

=0

o Inlet stream pressure: Column X-100: 101.3 kPa Column X-101: 233.8 kPa Specify splits o Column X-100 In overhead steam 100% recovery Cumene 100% recovery DIPB

In bottoms stream 100% recovery Benzene 100% recovery Propene Figure 46 – Design of splitter X-101 o Column X-101 In overhead steam 100% recovery Cumene

In bottoms stream 100% recovery DIPB

Figure 47 – Design of splitter X-101 Delete the pressure from the stream and defined vapour fraction = 0 (bubble point at 49 oC). Reset PD = 30 psia = 206.8 kPa

94

Heuristic: Column X-100

Column X-101

Heuristic: 1 kPa per stage – assume 27 stages kPA

Heuristic: 1 kPa per stage – assume 37 stages kPA

 PB = 30 psia x (6.895 psia ) + 27

 PB = 30 psia x (6.895 psia ) + 37

 PB = 233.85 kPa

 PB = 243.85 kPa

Create a new stream and defined it as bottom stream Define vapour fraction = 0 and PB = 233.85 kPa for column X-100 and 243.85 kPa for column X-101. There are no occur polymerization and double carbon bond. Table 34 – Determined result of pressure in distillation and bottom in column X-100 and column X-101 Column X-100

Column X-101

Pressure in distillation (kPa)

206.8 kPa

206.8 kPa

Pressure in bottoms (kPa)

233.85 kPa

243.85 kPa

2.

Short-cut column Table 35 – Specifying the parameter for column T-102 and column T-103 Column T-102

Column T-103

Component : Benzene

Component : Cumene

Mole fraction: 1x107

Mole fraction: 1x107

Heavy key in

Component : Cumene

Component : DIPB

distillation

Mole fraction: 0.0001

Mole fraction: 0.0001

Light key in bottoms

Condenser pressure (kPa)

206.8 kPa

206.8 kPa

Reboiler pressure (kPa)

233.8 kPa

243.8 kPa

Minimum reflux ratio (Rmin)

0.466

0.482

External reflux ratio

𝑅 = 2.2𝑅𝑚𝑖𝑛 = 2.2 × 0.466

𝑅 = 1.4𝑅𝑚𝑖𝑛 = 1.4 × 0.479

𝑅=1

𝑅 = 0.675

95

Read the output Table 36 – Summary the result of the output in column T-102 and column T-103 Column T-102

Column T-103

Number of stages N

27

37

Optimal feed stage

16

7

Condenser temperature

59.30oC

183.5oC

Reboiler temperature

190.4oC

252.4oC

3. Rigorous distillation columns 3.1. Column efficiency Table 37 –Temperature of each stage in column T-100 and column T-101 Column T-100

Column T-101

Condenser temperature (oC)

59.12oC

183.2oC

Reboiler temperature (oC)

190.0oC

252.2oC

Average temperature (oC)

124.56oC

217.7oC

The stage which have the closed value with the value of that temperature, is can be considered to be a main stage. o Main stage of column T-101 is stage 14, as the temperature of this stage is close to the average temperature of column ( T = 118.0oC  Taverage = 124.56oC) o Main stage of column T-101 is stage 14, as the temperature of this stage is close to the average temperature of column ( T = 217.1oC  Taverage = 217.7oC) -

Thus, the density and surface tension of main stage are found

Table 38 – Summary the mass density and surface tension in column T-100 and column T101 Column T-100

Column T-101

Mass density ( kg/m3)

 = 806.9 kg/m3

 = 704.9 kg/m3

Surface tension (dyne/cm)

 = 20.70 dyne/cm

 = 12.04 dyne/cm

96

Base on the values of condenser viscosity and reboiler viscosity in HYSYS, the average values are obtained 𝜇𝐿. Table 39 – Summary the viscosity of each stages in column T-100 and column T-101 Column T-100

Column T-101

Viscosity of condenser (cP)

µcondenser = 0.3630 cP

µcondenser = 0.1627 cP

Viscosity of reboiler (cP)

µreboiler = 0.1555 cP

µreboiler = 0.1174 cP

The average value μL

μL = 0.2593

μL = 0.1401

Read out the K-values of Cumene and DIPB for each condenser and reboiler stages respectively. Then calculate the average of K-values for each stage. Table 40 – Determination K-value for each stage in column T-100 and column T-101 Column T-100

Column T-101

Condenser

Condenser



K- values of Cumene in condenser: 0.0243



K- values of Cumene in condenser : 0.9928



K- values of Benzene in condenser: 0.2602



K- values of DIPB in condenser : 0.2906

 𝛼𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟

=

 𝛼𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 =

𝐵𝑒𝑛𝑧𝑒𝑛𝑒 𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 𝐶𝑢𝑚𝑒𝑛𝑒 𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟

0.2602 0.0243

= 10.71

 𝛼𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 =

𝐶𝑢𝑚𝑒𝑛𝑒 𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 𝐷𝐼𝑃𝐷 𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 0.9928

 𝛼𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 = 0.2906 = 3.416

Reboiler

Reboiler



K- values of Cumene in reboiler : 1.017



K- values of Cumene in reboiler : 2.733



K- values of Benzene in reboiler: 4.490



K- values of DIPB in reboiler : 0.9969

 𝛼𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 =

𝐵𝑒𝑛𝑧𝑒𝑛𝑒 𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 𝐶𝑢𝑚𝑒𝑛𝑒 𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 4.490

𝛼𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 = 1.017 = 4.415  𝛼 = (𝛼𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 + 𝛼𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 )/2

𝛼=

1.017+4.490 2

= 8.423

 𝛼𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 =

𝐶𝑢𝑚𝑒𝑛𝑒 𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 𝐷𝐼𝑃𝐷 𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 2.733

 𝛼𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 = 0.9969 = 2.741  𝛼 = (𝛼𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑟 + 𝛼𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟 )/2 𝛼 =

3.416+2.741 2

= 3.079

97

Then calculate the column efficiency by base on O’ Connel correlation: 𝐸𝑜 = 0.492 × (μL × 𝛼)−0.245 ± 10% Table 41- Determination of the column efficiency for column T-100 and column T-101 Column T-100

Column T-101

𝐸𝑜 = 0.492 × (μL × 𝛼)−0.245 ± 10%

𝐸𝑜 = 0.492 × (μL × 𝛼)−0.245 ± 10%

𝐸𝑜 = 0.492 × (0.2593 × 8.423)−0.245 ± 10%

𝐸𝑜 = 0.492 × (0.1401 × 3.079)−0.245 ± 10%

𝐸𝑜 = 0.447

𝐸𝑜 = 0.572

The range of efficiency for aqueous solution is from 40%-90%, thereby the efficiency for both column are accepted.

3.2. The number of actual stages Table 42 – Calculating the number of actual stages for column T-100 and column T-101 Column T-100 𝐸=

Column T-101

𝑁𝑡ℎ𝑒𝑜𝑟𝑦

𝐸=

𝑁𝑎𝑐𝑡𝑢𝑎𝑙

 0.447 =

27

𝑁𝑡ℎ𝑒𝑜𝑟𝑦 𝑁𝑎𝑐𝑡𝑢𝑎𝑙

 0.572 =

𝑁𝑎𝑐𝑡𝑢𝑎𝑙

37 𝑁𝑎𝑐𝑡𝑢𝑎𝑙

 𝑁𝑎𝑐𝑡𝑢𝑎𝑙 = 60

 𝑁𝑎𝑐𝑡𝑢𝑎𝑙 = 65

 Number of stage 60

 Number of stage 65

3.3. Height of column Table 43 – Summary the height of column for column T-100 and column T-101 Column T-100 Plate spacing

24 inches

Column T-101 24 inches

(inches) Disengagement and 5.871 + 5.871 = 11.742 ft

5.871 + 5.871 = 11.742 ft

column sump at the top and bottom (ft) Height of column (ft)

𝐻 = 22 × 𝑁𝐴𝑐𝑡𝑢𝑎𝑙 ×

𝐻 = 22 × 60 × 𝐻 = 122 𝑓𝑡

25 3

25 3

× 10−2 + 11.742

× 10−2 + 11.742

𝐻 = 22 × 𝑁𝐴𝑐𝑡𝑢𝑎𝑙 ×

𝐻 = 22 × 65 × 𝐻 = 131 𝑓𝑡

25 3

25 3

× 10−2 + 11.742

× 10−2 + 11.742

98

3.4.Column diameter Read out the HYSYS parameter Table 44 – Parameter specification in column diameter calculation Parameter

Symbol Units

HYSYS flow Variable

Liquid flow rate

L

lb/hr

Reflux – Mass flow

Vapour flow rate

G

lb/hr

To condenser – Mass flow

Liquid surface tension



dyne/cm

Reflux – Surface tension

Density of liquid

L

lb/ft3

Reflux – Mass density

Density of vapour

G

lb/ft3

To condenser – Mass density

Molar weight of liquid

ML

Kg/mole

Reflux – Molecular weight

Molar weight of vapour

MG

Kg/mole

Boil up – Molecular weight

lb/hr

Boil up – Mass flow

Flow rate of boiling up vapour Gboil up

Flow parameter could be determined by the equation: 𝐹𝐿𝐺 =

𝐿 𝜌𝐺 ×√ 𝐺 𝜌𝐿

1.1.1. Column T-100 𝐹𝐿𝐺 =

29220𝑙𝑏/ℎ𝑟 50.96 𝑙𝑏/𝑓𝑡 3 ×√ 7665 𝑙𝑏/ℎ𝑟 44.01 𝑙𝑏/𝑓𝑡 3 𝐹𝐿𝐺 = 4.10

Tray spacing = 0.6096 m = 24 inches Csb could be found by using Figure X: Flooding velocity determination plot and the value of flow parameter above.

99

Figure 48 – Flooding velocity determination plots for column T-100 Base on the figure X, Csb  0.01 ft/s = 0.00305 m/s Non foaming system  Ff = 1 for most distillation system 𝐴ℎ

The ratio of hole area and active area 𝐴𝑎 = 0.1  Fha = 1 The velocity of flooding U f : 𝑈𝑓 = 𝐶𝑠𝑏 × (

𝜎 0.2 𝜌𝐿 − 𝜌𝐺 0.5 ) × 𝐹𝑓 × 𝐹ℎ𝑎 × ( ) 20 𝜌𝐺

20.70 0.2 816.3 − 704.9 0.5 𝑈𝑓 = 0.00305 × ( ) ×1×1×( ) 20 704.9 𝑚 𝑈𝑓 = 0.0012 = 14.173 𝑓𝑡/ℎ 𝑠 Assume that flooding is 80%, therefore the velocity of vapour : 𝑈𝐺 = 0.8 × 𝑈𝑓 𝑈𝐺 = 0.8 × 0.0012 𝑚 𝑈𝐺 = 0.00096 = 11.338 𝑓𝑡/ℎ𝑟 𝑠 As FLG ≤ 0.1  The ratio of downcomer area and total area

𝐴𝑑 𝐴𝑡

= 0.1

100

Determine the column diameter: 0.5

4×𝐺 ) 𝐷𝑡 = ( 𝐴𝑑 𝑈𝐺 × 𝜋 × (1 − 𝐴𝑡 ) × 𝜌𝐺 0.5 4 × 7665 𝐷𝑡 = ( ) 11.338 × 𝜋 × (1 − 0.1) × 50.96

𝐷𝑡 = 4.33 𝑓𝑡 𝐻

122𝑓𝑡

Check the ratio of height – diameter 𝐷 = 4.33𝑓𝑡 = 28.17 ∈ [20,30]  Acceptable height for 𝑡

column design 1.1.2. Column T-101 𝐹𝐿𝐺 =

1303 𝑙𝑏/ℎ𝑟 44.41 𝑙𝑏/𝑓𝑡 3 ×√ 27920 𝑙𝑏/ℎ𝑟 41.28 𝑙𝑏/𝑓𝑡 3 𝐹𝐿𝐺 = 0.048

Tray spacing = 0.6096 m = 24 inches Csb could be found by using Figure X: Flooding velocity determination plot and the value of flow parameter above.

Figure 49 – Flooding velocity determination plots for column T-101

101

Base on the figure X, Csb  0.38 ft/s = 0.1158 m/s Non-foaming system  Ff = 1 for most distillation system 𝐴ℎ

The ratio of hole area and active area 𝐴𝑎 = 0.1  Fha = 1 The velocity of flooding U f : 𝑈𝑓 = 𝐶𝑠𝑏 × (

𝜎 0.2 𝜌𝐿 − 𝜌𝐺 0.5 ) × 𝐹𝑓 × 𝐹ℎ𝑎 × ( ) 20 𝜌𝐺

12.04 0.2 711.4 − 661.3 0.5 𝑈𝑓 = 0.1128 × ( ) ×1×1×( ) 20 661.3 𝑚 𝑈𝑓 = 0.028 = 330.71 𝑓𝑡/ℎ 𝑠 Assume that flooding is 80%, therefore the velocity of vapour: 𝑈𝐺 = 0.8 × 𝑈𝑓 𝑈𝐺 = 0.8 × 0.028 𝑚 𝑈𝐺 = 0.0224 = 264.57 𝑓𝑡/ℎ𝑟 𝑠 As FLG ≤ 0.1  The ratio of downcomer area and total area

𝐴𝑑 𝐴𝑡

= 0.1

Determine the column diameter: 0.5

4×𝐺 ) 𝐷𝑡 = ( 𝐴𝑑 𝑈𝐺 × 𝜋 × (1 − 𝐴𝑡 ) × 𝜌𝐺 𝐷𝑡 = (

0.5 4 × 27920 ) 264.57 × 𝜋 × (1 − 0.1) × 44.41

𝐷𝑡 = 1.833 𝑓𝑡 𝐻

131𝑓𝑡

Check the ratio of height – diameter 𝐷 = 1.833𝑓𝑡 = 71.45 𝑡

102

3.5. Multi-pass Trays Table 45 – Summary the liquid flow rate for column T-100 and column T-101

Liquid flow rate

Column T-100

Column T-101

13250 kg/h = 29220

591.2 kg/h = 1303 lb/hr

704.9 kg/m3 = 50.96 lb/ft3

661.3 kg/m3-= 41.28 lb/ft3

(lb/hr) Density of Liquid (lb/ft3) Unit conversation from lb/hr to gallon/min (gpm) The liquid flow rate QL (gal/min)

𝑄𝐿 =

𝐿×

𝑄𝐿 =

1 ℎ𝑟 ( ) 60 min

𝜌𝐿

𝑔𝑎𝑙

× 6.2288 (𝑓𝑡 3 )

𝑙𝑏 1 ℎ𝑟 29220 ( )× ( ) ℎ𝑟 60 min 𝑙𝑏 50.96 ( 3 ) 𝑓𝑡

𝑔𝑎𝑙

×

𝑄𝐿 =

𝐿×

𝑄𝐿 =

1 ℎ𝑟 ( ) 60 min

𝜌𝐿

𝑔𝑎𝑙

× 6.2288 (𝑓𝑡 3 )

𝑙𝑏 1 ℎ𝑟 )× ( ) ℎ𝑟 60 min 𝑙𝑏 41.28 ( 3 ) 𝑓𝑡

1303 (

×

𝑔𝑎𝑙

6.2288 (𝑓𝑡 3 )

6.2288 (𝑓𝑡 3 )

𝑄𝐿 = 59.53 𝑔𝑝𝑚

𝑄𝐿 = 2.65 𝑔𝑝𝑚

The number of tray pass can be calculated by using a plot of liquid flow rate Q L vs diameter DT

Figure 50 – A plot of liquid flow rate (gal/min) and column diameter (ft) Column T-100 had the liquid flow rate of 2.65 gpm and 4.33ft diameter, so the column is single pass Column T-101 had the liquid flow rate of 59.53 gpm and 1.833 ft diameter, so the column is single pass.

103

3.6. Pressure drop First, the velocity of vapour can be found by based on total cross-sectional area. 𝑢𝑣ap = 𝐺 ×

1 4 × 𝜌𝐺 𝜋 × 𝐷𝑇2

Table 46 – Vapour velocity in 2 columns Column T-100

Column T-100

1

4

1

𝑢𝑣ap = 𝐺 × 𝜌 × 𝜋×𝐷2 𝐺

𝐺

𝑇

𝑙𝑏

1

ℎ𝑟

1

𝑓𝑡 3

𝑢𝑣ap = 7665 (ℎ𝑟) × 3600 ( 𝑠 ) × 50.96 ( 𝑙𝑏 ) × 4

𝑙𝑏

4

(𝑓𝑡 2 )

𝑢𝑣ap = 0.0028

ℎ𝑟

(𝑓𝑡 2 )

𝑢𝑣ap = 0.0165

𝑠

1

1

𝜋×1.8332 𝑓𝑡

𝑇

𝑓𝑡 𝑠

Hole area is 10% of total active area Table 47 – Summary the velocity of hole for column T-100 and column T-101 Column T-100 𝑢𝑜 =  𝑢𝑜 =

Column T-100

𝑢𝑣ap

𝑢𝑜 =

0.1 0.0028 0.1

= 0.028

𝑓𝑡

 𝑢𝑜 =

𝑠

𝑢𝑣ap 0.1 0.0165 0.1

= 0.165

𝑓𝑡 𝑠

Then, calculating the dry tray pressure drop (Hd): 𝑈0 2 𝜌𝐺 ℎ𝑑 = 0.186 ( ) × ( ) 𝐶0 𝜌𝐿 Table 48 – Summary the dry tray pressure drop for column T-100 and column T-101 Column T-100

Column T-100 0.028 2

50.96

ℎ𝑑 = 0.186 ( 0.66 ) × (44.01)  ℎ𝑑 = 0.3876 𝑖𝑛𝑐ℎ𝑒𝑠

0.165 2

 ℎ𝑑 = 0.0125 𝑖𝑛𝑐ℎ𝑒𝑠

Table 49 – Summary the weir height for column T-100 and column T-101

(inches)

44.41

ℎ𝑑 = 0.186 ( 0.66 ) × (41.28)

The height of weir (hw) can be obtained from HYSYS

Height of weir

1

𝑓𝑡 3

𝑢𝑣ap = 27920 (ℎ𝑟) × 3600 ( 𝑠 ) × 44.41 ( 𝑙𝑏 ) ×

1

𝜋×4.332

4

𝑢𝑣ap = 𝐺 × 𝜌 × 𝜋×𝐷2

Column T-100

Column T-100

50.8mm = 2 inches

50.8mm = 2 inches

104

Next is determining the equivalent head on tray (hl): 2

𝑞𝑙 3 ℎ𝑙 = 𝜑𝑒 [ℎ𝑤 + 𝐶 ( )] 𝐿𝑤 𝜑𝑒 Table 50 – Summary the equivalent head on tray for column T-100 and column T-101 Column T-100 𝐴𝑑 𝐴𝑇

Superficial

Column T-100 𝐴

𝐴𝑑

= 0.1 𝐴𝑎 = 0.9

𝑢𝑎 =

𝐴𝑇

𝑇

𝑢𝑣ap 0.9

=

0.0028

= 0.003 𝑓𝑡/𝑠

0.9

𝐴

= 0.1 𝐴𝑎 = 0.9

𝑢𝑎 =

𝑇

𝑢𝑣ap 0.9

=

0.0165

= 0.018 𝑓𝑡/𝑠

0.9

vapour velocity (ft/s) Capacity parameter (ft/s)

𝑘𝑆 = 𝑢𝑎 × (𝜌

𝜌𝐺 𝐿 −𝜌𝐺

)

0.5

𝑘𝑆 = 𝑢𝑎 × (𝜌

50.96

 𝑘𝑆 = 0.003 × (50.96−44.01)

0.5

relative froth density

Weir length

𝐿 −𝜌𝐺

)

0.5

44.41

 𝑘𝑆 = 0.018 × (41.28−44.41)

 𝑘𝑆 = 0.0081 𝑓𝑡/𝑠 Effective

𝜌𝐺

0.5

 𝑘𝑆 = 0.0678 𝑓𝑡/𝑠

0.91 )

0.91 )

𝜑𝑒 = 𝑒 (−4257𝑘𝑠

𝜑𝑒 = 𝑒 (−4.257𝑘𝑠 0.91

 𝜑𝑒 = 𝑒(−4.257𝑥 0.0081

0.91

 𝜑𝑒 = 𝑒(−4.257𝑥 0.0678

)

 𝜑𝑒 = 0.95

)

 𝜑𝑒 = 0.69

𝐶 = 0.362 + 0.317 × 𝑒 −3.5ℎ𝑤

𝐶 = 0.362 + 0.317 × 𝑒 −3.5ℎ𝑤

 𝐶 = 0.362 + 0.317 × 𝑒 −3.5×2

 𝐶 = 0.362 + 0.317 × 𝑒 −3.5×2

 𝐶 = 0.362

 𝐶 = 0.362

𝐿𝑊 = 073𝐷𝑇 = 0.73 × 4.33

𝐿𝑊 = 073𝐷𝑇 = 0.73 × 1.833

 𝐿𝑊 = 3.16 𝑓𝑡 = 37.92 inches

 𝐿𝑊 = 1.338 𝑓𝑡 = 16.06 inches

(inches) Weir height (inches)

ℎ𝑙 = 𝜑𝑒 [ℎ𝑤 + 𝐶 (𝐿

𝑞𝑙

𝑤 𝜑𝑒

2 3

ℎ𝑙 = 𝜑𝑒 [ℎ𝑤 + 𝐶 (𝐿

)]

𝑄𝐿

𝑤 𝜑𝑒

2 3

2/3

)]

2.65

 ℎ𝑙 = 0.95 [2 + 0.362 (37.92×0.95)]

 ℎ𝑙 = 0.69 [2 + 0.362 (16.06×0.69)]

 ℎ𝑙 = 1.795 𝑖𝑛𝑐ℎ𝑒𝑠

 ℎ𝑙 = 1.127 𝑖𝑛𝑐ℎ𝑒𝑠

59.53

2 3

105

Assume that the maximum bubble diameter is equal to the diameter of dry tray pressure drop (DB(max) = DH) Table 51 – Summary the pressure drop due to surface tension for column T-100 and column T-101 Column T-100

Column T-100

Bubble diameter

DB(max) = DH = 6.35 mm-=6.35 x 103

DB(max) = DH = 6.35 mm-=6.35 x 103 m

(m)

m

Pressure drop due

ℎ𝜎 = 𝑔𝜌

to surface tension ℎ𝜎 =

(inches)

6𝜎

ℎ𝜎 = 𝑔𝜌

𝐿 𝐷𝐵(max)

6×20.70×102 (𝑑𝑦𝑛𝑒/𝑚) 9.81

𝑚 𝑙𝑏 ( 2 )×44.01 ( 3 )×6.35×103 𝑠 𝑓𝑡

ℎ𝜎 =

6𝜎 𝐿 𝐷𝐵(max)

6×12.04×102 (𝑑𝑦𝑛𝑒/𝑚) 𝑚

𝑙𝑏

9.81 ( 2 )×41.28 ( 3 )×6.35×103 𝑠 𝑓𝑡

ℎ𝜎 = 4.53 × 10−3 𝑚

ℎ𝜎 = 2.81 × 10−3 𝑚

ℎ𝜎 = 0.17835 inches

ℎ𝜎 = 0.11063 inches

The equation of total heat loss is: ℎ𝑡 = ℎ𝑑 + ℎ𝐼 + ℎ𝜎 Table 52 – Summary the total heat loss for column T-100 and column T-101 Column T-100

Column T-100

Total heat loss

ℎ𝑡 = ℎ𝑑 + ℎ𝐼 + ℎ𝜎

ℎ𝑡 = ℎ𝑑 + ℎ𝐼 + ℎ𝜎

(inches)

ℎ𝑡 = 0.3876 + 1.795 + 0.17835

ℎ𝑡 = 0.0125 + 1.127 + 0.11063

ℎ𝑡 = 2.361 𝑖𝑛𝑐ℎ𝑒𝑠

ℎ𝑡 = 1.250 𝑖𝑛𝑐ℎ𝑒𝑠

Tray pressure drop: Table 53 – Summary the tray pressure drop for column T-100 and column T-101 Column T-100 Tray pressure drop (psi)

Column T-100 𝑘𝑔

ℎ𝑡 × 𝜌𝐿 = 2.361 (𝑖𝑛𝑐ℎ𝑒𝑠) × 704.9 (𝑚3)

𝑘𝑔

ℎ𝑡 × 𝜌𝐿 = 1.250 (𝑖𝑛𝑐ℎ𝑒𝑠) × 661.3 (𝑚3)

= 1664.27 (𝑖𝑛. 𝑘𝑔/𝑚 3 )

= 826.63 (𝑖𝑛. 𝑘𝑔/𝑚 3 )

= 0.963 𝑝𝑠𝑖

= 0.478 𝑝𝑠𝑖

106

3.7 Other dimensions Table 54 – Summary the result of other dimensions for column T-100 and column T-101 Column T-100

Column T-101

12.7 mm

12.7 mm

3.404 mm

3.404 mm

Weir height (mm)

50.8 mm

50.8 mm

Wire length (ft)

2.7 ft

2.885 ft

Downcomer width

0.736 ft

0.312 ft

Flow path length

𝐹𝑝𝑙 = 𝐷𝑇 − 2 × 𝑤𝑑𝑙

𝐹𝑝𝑙 = 𝐷𝑇 − 2 × 𝑤𝑑𝑙

(ft)

𝐹𝑝𝑙 = 4.33 − 2 × 0.736

𝐹𝑝𝑙 = 1.833 − 2 × 0.312

𝐹𝑝𝑙 = 2.86 𝑓𝑡

𝐹𝑝𝑙 = 1.21 𝑓𝑡

38.1 mm

38.1 mm

0.1

0.1

Sieve tray hole (mm) Deck thickness (mm)

(ft)

Downcomer clearance (mm) Downcomer clearance/ tower area

107

Appendix 4 – Calculation procedure for HEX

Figure 51 – Overall Heat-Transfer Coefficients in Tubular Heat Exchangers

108

Fig LMTD correction factors for heat exchangers with one shell pass, two or more tube passes.

109

110

111

Figure 52 – Reynold number-Friction factor plot

112

Figure 53 – Reynold number-Friction factor plot

113

Appendix 5 – Economic Evaluation Fixed Capital Investment The fixed capital investment (FCI) is a total cost for all the equipment needed. All calculation will be carried out in MS Excel 2016 and the price for all equipment needs to be convert into current cost (2018). Moreover, necessary condition for calculation such as specified size for reactor, separator and distillation column, duty and area of pump, heat exchanger and cooler were determined from HYSYS. The trays and vessels which were required for operating also defined by using HYSYS simulation

Table 55 – Fixed Capital Cost for design process using the propylene feed stream contains 5% propane Equipment Type

Equipmen Material type t Name

Pressure (atm)

Fixed capital investment Pressu Materia User Input, re l Factor Unit factor

Utilization

Equipment Purchased cost cost

Current Cost

Total Installation Cost

P-100

Carbon Steel

13.4

0.6

0

power, kW

5.028

$630

$1,202

$1,182

$5,439

P-101

Carbon Steel

23.68

3

0

power, kW

12.71

$630

$1,742

$1,714

$11,995

Stainless Steel

0

0

4

area, m2

21.87

$1,030

$6,558

$6,451

$51,610

E-101

Carbon Steel

0

0

0

area, m2

17.84

$1,030

$5,803

$5,709

$22,837

E-102

Carbon Steel

0

0

0

area, m2

4

$1,030

$2,366

$2,328

$9,312

FH-100

Carbon Steel

0

0

0

duty, kW

1797.32

$635

$254,961

$250,826

$1,003,303

-

$42,997

$42,300

$338,397

-

$21,036

$20,695

$165,559

Pump Heat exchanger E-100 Heater/cooler Furnace

Stainless Steel Reactor

-

PFR-100

-

4

pressure, bar

0.071

height, m

10.45

diameter, m Stainless Steel Vessel

pressure, bar -

V-100

-

4

Stainless Steel

Tray Tower

-

4

Stainless Steel T-102

-

-

4

Total FCI

1.013

height, m

5.95

diameter, m

1.09

pressure, bar -

T-100

1.6

1

height, m

41.5

vessel diameter, m

1.3

pressure, bar

2.3

height, m

44.5

vessel diameter, m

0.55

-

$138,774.64 $136,524

$1,092,190

-

$110,305.03 $108,516

$868,127

$3,568,768

95

Table 56 – Fixed Capital Cost for design process using the propylene feed stream (99%)

96

Labor Costs The labour costs for both pathways is similar, which is show in the table below: Table 57 – Operating labour Operating Labor Equipment Auxiliary Facilities Air Plants Boilers Chimneys and Stacks Cooling Towers Water Demineralizers Electric Generation Plants Portable Generation Plants Electric Substations Incinerators Mechanical Refrigeration Units Waste Water Treatment Plants Water Treatment Plants Evaporators Vaporizers Furnaces Fans Blowers and Compressors Heat Exchangers Towers Vessels Pumps Reactors

Number of equipment Operators per Shift 0 0 0 0 0 0 0 0 0 0 0 0 Process Equipment 0 0 0 0 0 6 2 1 2 1

A single operator works on average 8-hour shifts per week: Total shift per year Process plant normally Plant operates Total operators required for the operation in a given shift Total Number of Operators Needed Expected Annual Salary Annual Cost of Operating Labor

1 1 0 1 0.5 0.5 3 0 2 0.5 2 2

0 0 0 0 0 0 0 0 0 0 0 0

0.3 0.05 0.5 0.05 0.15 0.1 0.35 0 0 0.5

0 0 0 0 0 0.6 0.7 0 0 0.5 (3 weeks’ time off for vacation and sick leave)

49 weeks per year 5 245 3 shifts per day 330 days 4.04 7.27

97

Total operators per shift

8 $63,800 $510,400.00

Moreover, it is important to evaluate other costs associates to Labour operating costs such as: employee oncost, Operating supervision, Laboratory labour. Each type of these factors is calculated by multiplying corresponding percentage with the operating labour cost. All specified information is provided in the design brief. The detail calculation on annual cost is presented by the Table 59: Table 58 – Detail calculation on annual cost Labour Costs Operating labour cost Employee oncost Operating supervision Laboratory labour Total

% 100% 20% 115% 115%

98

$510,400 $102,080 $586,960 $586,960 $1,786,400

Utilities The information of mass and energy flows is summarized from the HYSYS simulation and the cost of each utility type is give in the design brief. All utilities information are shown in the Tables below: Table 59 – Summary of utilities required in the process with propylene feed stream contain 5% propane Equipment Name P-100 P-101 E-102 E-103 PFR-100 T-100 Condenser T-101 Condenser T-100 Reboiler T-101 Reboiler DIPB disposal

Energy stream Q1 Q2 Q100 Q106 Q3 Q101 Q8 Q6 Q9 Bottom B2

Utility Type Electricity Electricity Cooling Water Cooling Water Cooling Water Cooling Water Cooling Water LP Steam LP Steam Waste Treatment

Utility per unit Operating hours (h) Quanlity per year Price per unit 5.03E+00 kW 7920 3.98E+04 $0.06 kWh 1.27E+01 kW 7920 1.01E+05 $0.06 kWh 3.96E+05 Kg/h 7920 3.14E+06 $0.02 m3 9350 Kg/h 7920 7.41E+04 $0.02 m3 3.23E+05 Kg/h 7920 2.55E+06 $0.02 m3 1.51E+05 Kg/h 7920 1.20E+06 $0.02 m3 2.95E+05 Kg/h 7920 2.34E+06 $0.02 m3 2691 Kg/h 7920 2.13E+07 $0.00 kg 2759 Kg/h 7920 2.19E+07 $0.00 kg 5.95E+02 Kg/h 7920 4.71E+06 $1 kg Total Utility Cost per Year

Annual Utility Cost Current cost $2,429 $ 2,389.73 $6,140 $ 6,040.86 $62,790 $ 61,771.32 $1,481 $ 1,457.02 $51,100 $ 50,271.01 $23,950 $ 23,561.61 $46,776 $ 46,016.83 $63,938 $ 62,901.09 $65,554 $ 64,490.56 $4,712,400 $ 4,635,965.31 $5,036,558 $4,954,865

Table 60 – Summary of utilities required in the process with propylene feed stream (99%) Equipment Name P-100 P-101 E-102 E-103 PFR-100 T-100 Condenser T-101 Condenser T-100 Reboiler T-101 Reboiler DIPB disposal

Energy stream Q1 Q2 Q100 Q106 Q3 Q101 Q8 Q6 Q9 Bottom B2

Utility Type Electricity Electricity Cooling Water Cooling Water Cooling Water Cooling Water Cooling Water LP Steam LP Steam Waste Treatment

Utility per unit Operating hours (h) Quanlity per year Price per unit 4.986 KJ/h 7920 3.95E+04 $0.06 kWh 15.9 KJ/h 7920 1.26E+05 $0.06 kWh 5.24E+05 Kg/h 7920 4.15E+06 $0.02 m3 29560 Kg/h 7920 2.34E+05 $0.02 m3 3.20E+05 Kg/h 7920 2.54E+06 $0.02 m3 2.66E+05 Kg/h 7920 2.11E+06 $0.02 m3 3.01E+05 Kg/h 7920 2.39E+06 $0.02 m3 3883 Kg/h 7920 3.08E+07 $0.00 kg 2840 Kg/h 7920 2.25E+07 $0.00 kg 1.17E+03 Kg/h 7920 9.23E+06 $1 kg Total Utility Cost per Year

99

Annual Utility Cost Current cost $2,409 $ 2,369.77 $7,682 $ 7,557.01 $82,954 $ 81,608.57 $4,682 $ 4,606.36 $50,736 $ 49,912.59 $42,134 $ 41,450.98 $47,726 $ 46,951.81 $92,260 $ 90,763.63 $67,478 $ 66,383.91 $9,226,800 $ 9,077,142.15 $9,624,861 $9,468,747

Raw materials and profits The cost for purchasing raw material and the profit from selling the product are calculated. Moreover, DIPB is unexpected product, which needs to be disposed. This means company have to considers cost of waste treatment and the cost of each utility type is give in the design brief. All raw material cost and profit information are shown in the Table below: Table 61 – Estimation of raw material cost and Profits (Propylene feed stream contain 5% propane impurity)

Raw Materials Benzene Propylene with 5% propane Catalyst Profit Materials Cumene Fuel gas

Raw Materials and Profits Quantity (kg/h) Operating hours Quanlity per year (kg/year) Price per kg Annual Cost 8,920 7920 70,646,400 1.12 $79,123,968 5,063 7920 40,098,960 0.88 $35,287,085 16,800 2.52 $42,336 Total Annual Raw Material Cost $114,453,389 Quantity (kg/h) Operating hours Quanlity per year (kg/year) Price per kg Annual Profit 12,635 7920 100,069,200 1.434 $143,499,233 283 7920 2,241,360 0.63 $1,412,057 Total Annual Profits $144,911,290

Table 62 – Estimation of raw material cost and Profits (Propylene feed stream (99%) Raw Materials Benzene Propylene Catalyst Profit Materials Cumene Fuel gas

Raw Materials and Profits Quantity (kg/h) Operating hours Quanlity per year (kg/year) Price per kg 8,920 7920 70,646,400 1.12 5,063 7920 40,098,960 1.57 16,800 2.52 Total Annual Raw Material Cost Quantity (kg/h) Operating hours Quanlity per year (kg/year) Price per kg 12,669 7920 100,338,480 1.434 152 7920 1,203,840 0.63 Total Annual Profits

99

Annual Cost $79,123,968 $62,955,367 $42,336 $142,121,671 Annual Profit $143,885,380 $758,419 $144,643,800

Summary of all costs Table 63 – Summary of all costs

% of FCI

5% propane impurity feed > 99% propylene Fixed Capital Investment 1 $ 3,568,768 $ 3,821,281 Land 0.04 $ 142,751 $ 152,851 Maintenance and Repairs 0.07 $ 249,814 $ 267,490 Total $ 3,961,333 $ 4,241,622 Utilities $ 4,954,865 $ 9,471,338 $ 1,786,400 Labour Costs Material cost $ 114,453,389 $ 142,121,671 Profit $ 144,911,290 $ 144,643,800 Contigencies 15% Capex $ 4,246,834 $ 4,547,325 Opex $ 121,444,468 $ 153,646,899 Savings $ 23,466,822 -$ 9,003,099 Total Capital Investment $ 125,691,302 $ 158,194,223 Capex, Opex and savings cost are calculated following these equation: 

Capex cost = Contingencies* ( Fixed Capital Investment)+ Land hiring cost



Opex cost = Maintenance and Repair cost+ Total (Utilities cost+ Labour cost + Raw material cost)



Savings = Profit from selling product – Opex cost

100

Net Present Value calculation tables

101

Table 64 – Summary of NPV and other economic indicator for pathways 1

Year

Capital 0 1 2 3 4 5 6 7 8 9 10

Total NPV PWPI Index Pre Tax Payback Atfter Tax Payback IRR

Savings $4,246,834 0 0 0 0 0 0 0 0 0 0 $4,246,834 $94,576,560.34 22.27 0.18 0.21 363%

$23,466,822 $23,466,822 $23,466,822 $23,466,822 $23,466,822 $23,466,822 $23,466,822 $23,466,822 $23,466,822 $23,466,822 $234,668,217

NPV Calculations Pre-tax cashflow Depreciation Taxable savings Tax paid After-tax cashflow -$4,246,834 $0 $0 $0 -$4,246,834 $23,466,822 $424,683 $23,042,138 $8,064,748 $15,402,073 $23,466,822 $424,683 $23,042,138 $8,064,748 $15,402,073 $23,466,822 $424,683 $23,042,138 $8,064,748 $15,402,073 $23,466,822 $424,683 $23,042,138 $8,064,748 $15,402,073 $23,466,822 $424,683 $23,042,138 $8,064,748 $15,402,073 $23,466,822 $400,000 $23,066,822 $8,073,388 $15,393,434 $23,466,822 $400,000 $23,066,822 $8,073,388 $15,393,434 $23,466,822 $400,000 $23,066,822 $8,073,388 $15,393,434 $23,466,822 $400,000 $23,066,822 $8,073,388 $15,393,434 $23,466,822 $400,000 $23,066,822 $8,073,388 $15,393,434 $230,421,383 $4,123,417 $230,544,800 $80,690,680 $149,730,703 Taxation rate 35% Discount rate 9%

102

Table 65 – Summary of NPV and other economic indicator for pathways

Year

Capital 0 1 2 3 4 5 6 7 8 9 10

Total NPV PWPI Index Pre Tax Payback After Tax Paypack IRR

Savings $4,547,325 0 0 0 0 0 0 0 0 0 0

($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099) ($9,003,099)

$4,547,325 ($90,030,991) -$41,130,566.23 -9.05 -0.51 0.58

NPV Calculations Pre-tax Taxable cashflow Depreciation savings Tax paid After-tax cashflow -$4,547,325 $0 $0 $0 -$4,547,325 -$9,003,099 $454,732 -$9,457,832 -$3,310,241 -$5,692,858 -$9,003,099 $454,732 -$9,457,832 -$3,310,241 -$5,692,858 -$9,003,099 $454,732 -$9,457,832 -$3,310,241 -$5,692,858 -$9,003,099 $454,732 -$9,457,832 -$3,310,241 -$5,692,858 -$9,003,099 $454,732 -$9,457,832 -$3,310,241 -$5,692,858 -$9,003,099 $400,000 -$9,403,099 -$3,291,085 -$5,712,014 -$9,003,099 $400,000 -$9,403,099 -$3,291,085 -$5,712,014 -$9,003,099 $400,000 -$9,403,099 -$3,291,085 -$5,712,014 -$9,003,099 $400,000 -$9,403,099 -$3,291,085 -$5,712,014 -$9,003,099 $400,000 -$9,403,099 -$3,291,085 -$5,712,014 -$94,578,315 $4,273,662 $94,304,653 -$33,006,629 -$61,571,687 Taxation rate 35% Discount rate 9%

#NUM!

103

Appendix 6 – HAZOP Line no

Description

Guide Word

Deviation

Feed flow None

Pressure

Possible Causes

Consequenc es / Concern

Action Required

No feed input

No reaction

Add feed

Not feasible

Temperature Not feasible malfunctioned

Feed 26

Benzen feed Feed flow Too much feed input More of

Pressure

malfunctioned

Temperature

Increased feed temperature

104

Increase reaction rate, increase pressure, increase temperature, product deterioration , risk of catastrophic thermo runaway reaction Pipeline leak/burst Product deterioration ,

Install flowmeter and flow alarm

Monitor feed input

Perform regular valve inspection and maintenance Reduce feed temperateure Monitor feed temperature

catastrophic thermal runaway reaction Feed flow

malfunctioned feed line blocked

Less of

Pressure

malfunctioned

Low efficiency Low efficiency

Check, perform regular valve/line inspection and maintenance Install flowmeter and flow alarm Install flowmeter and flow alarm, perform regular valve inspection and maintenance

Temperature Not feasible Impurities in input feed

Control system malfunction

Decrease reaction efficiency, product contaminati on, risk of leak or explosion Decrease reaction efficiency, side reactions Reactor inactive Reactor inactive

No feed input

No reaction

Impurities Feed line corroded More than

Phase

No power Other than

25

Benzen recovery feed

None

Start-up failed Feed flow

Feed temperature changed

Power outage

105

Ensure feed quality, filter feed properly

Perform regular valve/line inspection and maintenance

Monitor feed temperature

Shut off system and wait for restart Check system, perform regular system inspection and maintenace check recovery line

Pressure

Not feasible

Temperature Not feasible malfunctioned

Feed flow Too much recovery

More of

Pressure

Temperature

Feed flow

Malfunctioned

Increased temperature

Malfunctioned Feed line blocked

Less of Pressure

Malfunctioned

106

Increase reaction rate, increase pressure, increase temperature, product deterioration , risk of catastrophic thermo runaway reaction Pipeline leak/burst Product deterioration , catastrophic thermal runaway reaction Low efficiency Low efficiency

Install flowmeter and flow alarm

Monitor feed input

Perform regular valve inspection and maintenance Reduce feed temperateure

Monitor feed temperature

Check, perform regular valve/line inspection and maintenance Install flowmeter and flow alarm Install flowmeter and flow alarm, perform regular valve inspection and maintenance

Temperature Not feasible Impurities in input feed Impurities Feed line corroded More than

Phase

Other than

No power

Product flow

Feed temperature changed

Power outage

Ensure feed quality, filter feed properly

Perform regular valve/line inspection and maintenance

Monitor feed temperature

Shut off system and wait for restart

Power outage

Reactor inactive

Shut-off all valves and wait for restart

No feed output

Reactor inactive

check feed line, performen regular valve/line inspectrion and maintenance

Increase reactor

Install flowmeter and flow alarm

None 4

Decrease reaction efficiency, product contaminati on, risk of leak or explosion Decrease reaction efficiency, side reactions Reactor inactive

Mixed line Pressure

Not feasible

Temperature Not feasible More of

Product flow

High flow rate

107

Pressure

High output pressure

Cooling coil malfunctioned Temperature

Insufficient cooling water supply

108

pressure, increase reactor temperature, product deterioration and risk of catastrophic thermal runaway reaction Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion Damage the structural integrity of the reactor vessel and its components Increase reactor pressure, increase reactor

Install mixer pressure monitor, pressure alarm, pressure high trip and automatic relief valve, install flowmeter and flow alarm

Install vessel temperature monitor, temperature alarm and temperature high trip

Stirrer malfunctioned Thermocouple malfunction

Product flow

Mixed line blocked

Pressure

Mixer leaked

Less than

109

temperature, product deterioration and risk of catastrophic thermal runaway reaction Local Perform regular vessel inspection and maintenance overheating Perform regular thermocouple inspection and maintenance Increase reactor pressure, increase reactor temperature, damage the Perform regular inspection and maintenance structural integrity of the reactor vessel and its components Potential hazardous for the Check and maintain vessel routinely environment Install vessel pressure monitor, pressure alarm and employees

Risk of explosion Line 4 partualy blocked

Temperature Feed flow too high

Reactor components corroded

More than

Impurities Mixer leaked

Impure feed

110

Low efficiency Unknown side reactions may occur Decrease reaction efficiency, product contaminati on, risk of leak or explosion Side reactions, product contaminati on Side reactions, product contaminati on

Install flowmeter and flow alarm

Install temperature alarm

Perform regular pipe line and reactor vessel inspection and maintenance

Perform regular pipe line and reactor vessel inspection and maintenance

Monitor feed quality

Line no

Description

Guide Word none

Deviation flow rate flow rate

2

propene and propane feed line

more of

less of

none 3

after pumb line more of

pressure

Possible Causes No feed input Too much flow rate High flow rate Carbon deposit

Consequences / Concern

Action Required

No feed line in

Add feed

increase pressure, increase temperature pipe leak/brust

Perform regular inspection and maintenance Perform regular inspection and maintenance Perform regular inspection and maintenance

Product deterioration, catastrophic thermal runaway reaction increase Product deterioration, temperature feed catastrophic thermal runaway temperature reaction feed line low efficiency flow rate blocked low low efficiency pressure feed line pressure leak in feed line No feed No feed line in flow rate output None pump Unchange pressure, unchange pressure duty temperature None pump Unchange pressure, unchange temperature duty temperature Too much increase pressure, increase flow rate flow rate temperature High flow pipe leak/brust pressure rate

111

Monitor feed temperature

install flowmeter and flow alarm install flowmeter and flow alarm, preform regular inspection and maintenance install presuremeater Check feed line check pump duty check pump duty Perform regular inspection and maintenance Perform regular inspection and maintenance

Carbon deposit

less of

Line no

Product deterioration, catastrophic thermal runaway reaction high output Product deterioration, temperature temperature catastrophic thermal runaway reaction output line low efficiency flow rate blocked leak in pipe low efficiency pressure

Description Guide Word Deviation none

flow rate

Possible Causes no feed input Too much feed input

flow rate

mix

cool stream in

More of pressure

less than

High flow rate Carbon deposit

Perform regular inspection and maintenance Monitor output temperature

install flowmetar and flow alarm install presuremeater

Consequences / Concern

Action Required

No feed in line

check feed line

Increase pressure, increase temperature, product deterioration, risk of catastrophic thermo runaway reaction pipe leak/brust

monitor feed input

Product deterioration, catastrophic thermal runaway reaction increase Product deterioration, temperature feed catastrophic thermal temperature runaway reaction feed line low efficiency flow rate blocked

112

Perform regular inspection and maintenance require pipe lines maintenance

Perform regular inspection and maintenance install flowmeter and flow alarm

none

more of 5

cool stream out

low pressure feed line pressure leak in feed line no feed flow rate input pipe blocked Flow control Flow rate failure High output Pressure flow rate Machine performent temperature failure output line blocked leak in feed pressure line Machine performent temperature failure flow rate

less than

nore 14

hot steam out

flow rate flow rate

More of Pressure

low efficiency

install flowmeter and flow alarm, preform regular inspection and maintenance install presuremeter

No feed in line

Check feed line

increase pressure on pipes

Perform regular inspection and maintenance install FAH, FAHH

pipe leak/brust

install FAH, FAHH on pipe

Product deterioration, catastrophic thermal runaway reaction, change product purity low efficiency

Monitor output temperature, install TAH, TAHH

low efficiency

Install PAH,PAHH to mornitor pressure on pipe install TAH, TAHH to monitor output temperature

Product deterioration, catastrophic thermal runaway reaction, change product purity no output stream

pipe blocked Flow control increase pressure, reduce failure product purity High flow increase pressure on pipes rate

113

install FAH, FAHH on pipe

Perform regular inspection and maintenance install FAH, FAHH install FAH, FAHH on pipe

less than

wrong temperature machine funtrion output line flow rate blocked leak in feed pressure line

Product deterioration, catastrophic thermal runaway reaction low efficiency

Perform regular inspection and maintenance

low efficiency

install presuremeter

114

install flowmeter and flow alarm

115

Line Description Guide Word no

none

Deviation

flow rate

flow rate

more of 6

pressure

after heat line

other than

Phase change

equipment failure output line flow rate blocked equipment failure temperature air control failure operator incorrect errors, feed speed design errors temperature

less than

Possible Causes flow control failure No input feed high input flowrate air control failure

Consequences / Concern

Action Required

No reaction

install FAL

check feed line decrease product quality, changes product composition

install FAH,FAHH on the product line install air flow controller, perform regular inspection and maintenance

Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion

Install reactor vessel pressure monitor, pressure alarm, pressure high trip and automatic relief valve

reaction speed rises, side reaction can occurs low efficiency

install TAH, TAHH on the vessel

Product deterioration, catastrophic thermal runaway reaction

Perform regular inspection and maintenance, install TAL

incoract reaction, Change product purities

stant operation, test plant design and simluation

116

install flowmeter and flow alarm

Lin e no

Descriptio n

Guid e Word

none

Deviation

flow rate

flow rate 7

reactor product line pressure more of

temperatue

Possible Causes

Consequences / Concern

Action Required

flow control failure reactor operator failure pipe blocked

No product gain, damage structural integrity of separator vessel

install FAL,

power outage

reactor inactive

flow control failure

reduce product yield, reduce reactor efficiency

high flow rate

Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion

high reaction rate increase product voidage increase feed stream temperature high reaction rate side reaction

Perform regular inspection and maintenance,

Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion reduce product yield, reduce reactor efficiency, risk of leak or explosion, damage the structural integrity of the reactor vessel and its components reduce poduct purities

117

shutdown process and wait for restar install FAH, FAHH, perform regular inspection and maintenance install FAH, FAHH

shutdown process and wait for restar, perform maintenance install TAHH, TAH, perform regular inspection and maintenance

less of

flow rate

temperatue

More than

Lin e no

Descriptio n

Guid e Word

Impurities

Deviation

flow control failure pipe partial blocked reduce reactor duty Reactor components corroded

low product output, pipes burst,

Impure feed Possible Causes

Side reactions, product contamination Consequences / Concern

pump failure

reduce product purity, reduce reaction rate Decrease reaction efficiency, product contamination, risk of leak or explosion

16

pressure temperatur e

feed line separator

install FICA to feed line, perform regular inspection and maintenance

Increase reaction rate, increase pressure, increase temperature, product deterioration, risk of catastrophic thermo runaway reaction fall in reaction rate, rate of separation fall/reduce turbulence flow occurs,

Mornitor feed input

level valve failure No feasible No feasible too much feed input

more of

flow rate

Perform regular pipe line and reactor vessel inspection and maintenance Monitor feed quality Action Required

No separation process, no production, pump damage

flow rate none

intall FAL, Perform regular pipe line and reactor vessel inspection and maintenance Check reactor duty

level control failure partial failure pump

118

install independent level transmitter and FAH install automatic pump shutdown

pressure temperatur e

flowrate less than

pressure temperatur e

feed line blocked increase feed temperature pipe leaked

blocked in pipe partial failure pump control inlet valve failure leaked in steam No feasible

18

none

pressure temperatur e more of

flow rate

install flowmeter and flow alarm Mornitor feed input temperature Perform regular inspection and maintenance, install FAL

Low efficiency

install flowmeter and flow alarm

Power outage

Separator inactive

LIC failure

No product gain, damage structural integrity of separator vessel

Shut-off all valves and wait for restart install LIC control, FAL, perform regular inspection and maintenance

flow rate underline separator stream

Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion Product deterioration, catastrophic thermal runaway reaction rate of production reduces, rate of separation reduces, level fall.

valve blocked feed line blocked No feasible No feasible level control failure

reduce product purity, low separation efficiency

119

Install FAH, FAHH, install automatic pump shutdown

partial failure pump value fully open Pipe blocked pressure

temperatur e

value fully open Stirrer malfunctione d stream blocked

flow rate less than pressure temperatur e

Line no

20

Description

column under stream

Guide Deviation Word flow rate none pressure temperature

valve partial blocked leaked in steam high water flowrate

Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion pipes broken, leaked, reduce product yield local overheat

Damage the structural integrity of the reactor vessel and its components, risk of leak or explosion

Product deterioration, catastrophic thermal runaway reaction Decrease reaction rate

Possible Consequences / Concern Causes Product feed No feed in line pump failed isolating valve jammed no feasible no feasible

120

Perform regular inspection and maintenance, install PAHH,PAH install PAH, PAHH Perform regular inspection and maintenance, install TAH, TAH install flowmeter and flow alarm

install pressuremeter install flowmeter and flow alarm

Action Required check feed line,Install flowmeter and flow alarm

flow rate

pressure more of temperature

high feed flow Increase reaction rate, increase rate pressure, increase temperature, product deterioration, risk of catastrophic thermo runaway reaction high flow rate pipe leak/burst

Mornitor feed input

high reboiler temperature

decrease product purity

high column temperature expantion of hot product

loss of feed

Install vessel temperature monitor, temperature alarm and temperature (high and low) trip

brust pide

preform regular inspection and maintenance

hot product load on column Product feed pump failed

high temperature in furnace

install pyrometer in furance

temperature rise in column, drop in liquid level in column, overheating

Install flowmeter and flow alarm, perform regular valve inspection and maintenance

flow rate less than pressure

more than

isolating valve jammed leaked Potential hazardous for the column environment and employees Risk of explosion

water form atmosphere contamination through vent

water turns steam and explodes

121

Perform regular valve inspection and maintenance

Check and maintain reactor vessel routinely Install vessel pressure monitor, pressure alarm locate tank to be in hot system, stream vent at high point in pipe system

LIC failed none

flow rate

flow rate more of 19

column upper stream more than

less than

none

18

feed stream

pump damage

valve failed over flow feed more feed in condenser, lower product yield failure in increase pressure on condenser valve condenser

Install flowmeter and flow alarm install pressure indicator on column, install high pressure alarm

failed flow stream cooling of condenser and column after shutdown LIC failed

install thermocouple on valve

pressure

reverse flow

excress pressure suck back air into column on cooling

pump damage

Install flowmeter and flow alarm, perform regular valve inspection and maintenance

feed pump failed valve jammed flow controller fault

temperature rise in column, drop in liquid level in column, overheating

install TIC to the vessel to control temperature

level risen in column,hence temperature fall, reboiler capacity reached, column stop operating

install independent high flow alarm

feed temperature increase

Product deterioration, catastrophic thermal runaway reaction

Mornitor feed input, install TAH and TIC to control feed temperature

flow rate valve jammed flow rate

flow rate more of temperature

Install flowmeter and flow alarm, perform regular valve inspection and maintenance

122

pressure

flow rate less than pressure

valve pipe line leak/burst malfunctioned

preform regular inspection and maintenance

feed pump blocked valve blocked

install TIC to the vessel to control temperature, FIC to control feed flow

temperature rise in column, drop in liquid level in column, overheating

valve Low efficiency malfunctioned

123

install FICA, perform regular inspection and maintenance

Appendix 7 – Date of meeting

Date of meeting

15/09/2018

Time & location

12 pm – 1 pm at Hub central

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

No

Facilitator

Nhut Nguyen

Recorder

Thanh Le 20/09/2018

Next meeting scheduled

Points to be discussed

Decisions made

Responsibilities assignedWhat/Who/When/Where

Background information

Research information about 5 chemical

Everyone – upload files into

components

group chat on Facebook – due date 20/09/2018

Assessment tasks

Review assessment tasks to understand all

Everyone – at this meeting

requirements and the process of project

– in Hub central

124

Points discussed - Pending decisions

Literature reviews

Look for literature articles relating to

Everyone – upload files into

cumene production

group chat on Facebook – due date 20/09/2018

Record meeting minutes

Thanh Le – record every

One of member will record all minutes of meetings for project

meeting

Problems encountered

Solutions

Communication

New group chat is created on Facebook to contact with each members

125

Date of meeting

20/09/2018

Time & location

12pm – 7pm – Cat suite B.23

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

No Hieu Nguyen

Facilitator Recorder

Thanh Le

Next meeting scheduled

28/09/2018

Points to be discussed

Decisions made

Responsibilities assignedWhat/Who/When/Where

Means-End Analysis

-

The analysis will begin to work after

Everyone – at this meeting

researching for background information.

– in Cat suite B.23

-

Points discussed - Pending decisions

Calculation overall mass balance, the flow rate reactants and products required, and Eliminate differences in composition ( Step 1, 2 and 3 )

Process flow diagram (PFD)

Everyone – at this meeting

Discussed about draft PFD

-

– in Cat suite B.23

Create a draft PFD and compare with others from published literature articles.

Problems encountered

Make a decision for a final PFD.

Solutions

126

Find a suitable parameter of flow rate, temperature and pressure for

-

Base on the Mean – End analysis to decide on these parameter

PFD

-

Read articles and lectures

127

Date of meeting

28/09/2018

Time & location

12pm – 1pm, Cat Suite B.24

Present

Hieu Nguyen, Nhut Nguyen, , Thanh Le,

Absent

No Phuc Truong

Facilitator Recorder

Trang Tran

Next meeting scheduled

03/10/2018

Points to be discussed

Means-End Analysis

Decisions made

-

Responsibilities assigned-

Points discussed - Pending

What/Who/When/Where

decisions

Hieu Nguyen – due date

Complete remain steps of this section.

03/10/2018 Build a base case

-

Everyone – due date 03/10/2018

Began to create a base case on HYSYS

Decide to choose only one base case for group

Problems encountered

Solutions

128

Date of meeting

03/10/2018

Time & location

2pm – 3pm, Cat Suite 2.35

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

No Hieu Nguyen

Facilitator Recorder

Thanh Le

Next meeting scheduled

12/10/2018

Points to be discussed

Decisions made

Build a base case

-

Responsibilities assignedWhat/Who/When/Where

Chose Nhut Nguyen’s Base case to work

-

on

Points discussed Pending decisions

Everyone – at this meeting – Cat Suit 2.35

-

Nhut Nguyen & Hieu Nguyen – Reactor – Due date 14/10/2018

Process design

-

Divided into small group.

Each group has responsible for each part

Phuc Truong & Thanh Le –

Need to repair on data for

Separation Columns – Due date

Distillation Columns to

14/10/2018

match with the

of project.

requirement of project

-

Hieu Nguyen & Trang Tran – Heat exchanger – Due date 14/10/2018

129

Problems encountered

Solutions

130

Date of meeting

12/10/2018

Time & location

3pm – 9pm Cat Suite G.20

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

No Hieu Nguyen

Facilitator Recorder

Trang Tran

Next meeting scheduled

19/10/2018

Points to be discussed

Decisions made

Responsibilities assignedWhat/Who/When/Where

The desired purity of

-

-

cumene does not achieve

Re-run the column with another parameters of temperature and pressure.

-

Points discussed Pending decisions

Phuc Truong & Thanh Le – at that meeting

Accept the difference between calculated and displayed results in HYSYS.

Problems encountered

Solutions

131

Date of meeting

19/10/2018

Time & location

2 pm – 7 pm, Cat Suite 2.35

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

No Trang Tran

Facilitator Recorder

Thanh Le

Next meeting scheduled

22/10/2018

Points to be discussed

Decisions made

Responsibilities assignedWhat/Who/When/Where

Contents of report

Made a draft table of contents

Phuc Truong

Introduction

Trang Tran – due date 24/10/2018 – upload to

Literature review

group chat on Facebook

Mean - Ends Analysis

Nhut Nguyen – due date 24/10/2018 – upload

Base case design

to group chat on Facebook

Process design: Reactor

Hieu Nguyen – due date 24/10/2018 – upload

Process optimization: Reactor

to group chat on Facebook

Group report

Calculation procedure for Reactor Process design: Separation column

Thanh Le – due date 24/10/2018 – upload to

Process optimization: Separation column

group chat on Facebook

Calculation procedure for column

132

Points discussed Pending decisions

Process design: heat exchanger

Trang Tran – due date 24/10/2018 – upload to

Process optimization: Heat exchanger

group chat on Facebook

Calculation procedure for heat exchanger Phuc Truong – due to 24/10/2018 – upload to

Hazard

group chat on Facebook Problems encountered

Solutions

133

Date of meeting

22/10/2018

Time & location

3pm – 5-pm, Cat Suite 2.35

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

None Thanh Le

Facilitator Recorder

Phuc Truong

Next meeting scheduled

Points to be discussed

Hazop meeting

24/10/2018

Decisions made

Responsibilities assigned-

Points discussed -

What/Who/When/Where

Pending decisions

Everyone – in this meeting – Cat suite 2.35

Estimated the potential risks related to operating condition of cumene production plant

Hazop record

Project report

All ideas are recorded to complete the

Phuc Truong – in this meeting – Cat suite

HAZOP assessment

2.35

All contents were almost done

Everyone

Problems encountered

Solutions

134

Date of meeting

24/10/2018

Time & location

5pm – 9pm, Cat Suite 2.35

Present

Hieu Nguyen, Nhut Nguyen, Phuc Truong, Thanh Le, Trang Tran

Absent

None Hieu Nguyen

Facilitator Recorder

Trang Tran

Next meeting scheduled

Points to be discussed

Project report

Decisions made

Responsibilities assigned-

Points discussed -

What/Who/When/Where

Pending decisions

Everyone – in this meeting – Cat suite 2.35

Combine all the parts of project report Problems encountered

Solutions

135