CPD-3284 - TU Delft

CPD-3284 Conceptual Process Design Process Systems Engineering DelftChemTech Faculty of Applied Sciences Delft University of Technology

Subject Novel catalytic method for alcohol oxidation

Authors Telephone Lin Luo (1113259) +(31)- 641246881 Weimin Wang (1118668) +(31)- 619922615 Shuang Zhao (1118447) +(31)- 641389123 Zhengjie Zhu (1113658) +(31)- 624518226

Keywords 3,3-dimethylbutyraldehyde, 3,3-dimethylbutanol, TEMPO, RuCl2(PPh3)3, Catalytic oxidation

Assignment issued : Oct 10th, 2002 Report issued : Jan 15th, 2003 Appraisal : Feb 11th, 2003

ST4931 Conceptual Process Design Design of a plant utilizing novel catalytic method for alcohol oxidation CPD-3284

Lin Luo (1113259) Weimin Wang (1118668) Shuang Zhao (1118447) Zhengjie Zhu (1113658)

79

List of Symbols

Symbol Description SI Units A area m2

c ratio between Hs and D’ - CA concentration of reactant kmol/m3 Ci Liq. mole composition - CG capacity factor - CP heat capacity kJ/kmol Kd diameter m D diameter m D’ diameter m FC flow rate of CW kg/h FLG flow parameter - Ft temperature correction

factor -

g gravitational acceleration m2/s G flow rate kg/s h height m H height m H specific enthalpy kJ/kmol H heat kJ Hf, enthalpy of formation kJ/kmol Hr enthalpy of reaction kJ/kmol Hv enthalpy of evaporation kJ/kmol k reaction constant h-1

L liquid - Ls flow path length m m mass flow rate kmol/h ML the mass flow rate of

liquid kg/s

MG the mass flow rate of gas kg/s Nmix mixing number - Np power number - P pressure Pa Pd discharge pressure bar Ps suction pressure bar Ps power input per stirrer w

P pressure difference across the pump

N/m2

Q amount of heat J QP flow rate m3/s Q heat transferred per unit

time kW

R gas constant J/mol K



80

Symbol Description SI Units S entropy J/mol K S solid - t temperature oC (K) T temperature oC (K) Tm the main temperature

difference, the temperature driving force

oC (K)

tm mixing time s

u superficial velocity ms U overall heat transfer

coefficient kw/m2 K

v velocity m/s V volume m3

w work J (kJ) Greek Description SI Unit

j activity coefficient of component I

-

jv fugacity coefficient of

component i -

total specific power input w/kg homogeneity factor - primary eddy size - average viscosity index - viscosity -2Nm s v flow rate m3/s surface tension -

L density for the liquid kg/m3

g density for the gas kg/m3

L liquid holdup -

0 orifice coefficient -

p pump efficiency -

reaction degree -



81

Reference 1. A. Dijksman, Catalytic oxidation of alcohols. TU Delft, Uitgave: Ponsen & Looijen BV, Wageningen, 2001, ISBN 90-6464848-4; Ozs/Ins NIOK; 2. A. Dijksman, A. Marino-González, A. Mairata i Payeras, I.W.C.E. Arends, R.A. Sheldon, Efficient and Selective Aerobic Oxidation of Alcohols into Aldehydes and Ketones using Ruthenium/TEMPO as Catalytic System, In: Journal of the American Chemical Society, Jaargang: 123, 2001, p. 6826-6833, ISSN 00027863; Ozs/Ins NIOK; BTA wet6; 3. NIST Chemistry Web Book, http://webbook.nist.gov/chemistry, NIST Standard Reference Database Number 69-July 2001 Release; 4. James M. Douglas, Univ. of Massachusetts (USA), Conceptual Design of Chemical Processes, Mc-Graw-Hill, 1988; 5. M. hudlicky, Oxidation in Organic Chemistry, ACS, Washington DC, 1990; 6. Instuction Manual, Conceptual Design, J. Grievink, C.P.Luteijn, P.L.J. Swinkels, July 2002; 7. Chemical Market Report, http://www.chemicalmarketreporter.com/home/frameset.htm 8. Coulson & Richardson, R. K. Sinnott, University College, Swansea (U. K.), Chemical Engineering, Volume 6, Design, second edition, Pergamon Press, 1993; 9. Smith, Van Ness & Abbott, Introduction to Chemical Engineering Thermodynamics, Fifth edition, Mcgraw-Hill International Editions, 1996; 10. CRC Handbook of Chemistry and Physics; 11. http://www.sigmaaldrich.com; 12. J. D. Seader & Ernest J. Henley, Separation Process Principles, John Wiley & Sons, Inc., 1998; 13. H. Scott Foggler, Elements of Chemical Reaction Engineering, third edition, Prentice Hall, 1999; 14. J. Harmsen, G. Korevaar, Sustainable Technology, Course Book, September 2002; 15. Kirk-Othmer, Encyclopaedia of Chemical Technology; 16. James G. Speight, Chemical and Process Design Handbook, Mc-Graw-Hill, 2002; 17. G. Stephanopoulos, 1984, Chemical Process Control: An Introduction to Theory and Practice, Prentice Hall Inc., 1984;



82

18. J. M. Pransitz, Computer Calculations for Multi-component Vapor-liquid and Liquid- liquid equilibra, Prentice-hall inc., 1990; 19. L. Venkatesh, Choice of Thermodynamic Models for use in Simulation Programs, Chem.Eng. World, 1997; 20. W. D. Seider, Process design principles, Synthesis, Analysis, and Evaluation, John Wiley & Sons, Inc., 1998; 21. R. C. Reid, The Properties of Gases & Liquids, 1987; 22. Method for preparing 3, 3-dimethybutyraldehyde, US5973209, Chapeau Marie- Christine D (US), Prakash Indra (US), 1999-10-26; 23. Preparation of 3, 3-dimethybutyraldehyde by oxidation of 3, 3-dimethybutanol, US5856584, Prakash Indra (US), Ager David. J (US), Katritzky Alan R (US), 1999-01-05; 24. H. J. Pasman, S. M. Lemkowitz, Chemical Risk Management, Course Book, 2001; 25. Dow’s Fire and Explosion Index Hazard Classification Guide, 7th edition, AIChE Technical Manual; 26. http://www.wesellchemicals.com/Methyl%20Isobutyl%20Carbinol.doc; 27. Highlights from the Literature, Some Items of Interest to Process R&D Chemists and Engineers as Selected by Trevor Laird and Stephen A. Hermitage, Organic Process Research & Development 2001, 5, 460-466; 28. http://www.dct.tudelft.nl/~meeuse/workshop_korevaar.pdf; 29. J.J. Heijnen, 2001, Bioprocess Technology, St 322 course syllabus; 30. Coulson J.M & Richardson J.F, 1996, Coulson & Richarson’s Chemical Engineering, volume 1, Fluid flow, Heat transfer and Mass transfer, 5th edition, Oxford, Butterworth – Heinemann; 31. A.T.Jackson, process engineering in biotechnology, 1990; 32. Walas, 1959, Reaction Kinetics for Chemical Engineers, McGraw-Hill; 33. P.A.Schweitzer, handbook of separation techniques for chemical engineers (third edition), 1996; 34. Don W. Green, James O. Maloney and Robert H. Perry (edited), 1998, Perry’s Chemical Engineers’ handbook, 7th edition, McGraw – Hill; 35. Dr. Z. Olujic, Laboratory for Process Equipment, Delft University of Technology; 36. J.Davidson & O.v. Bertele, process fan and compressor selection, 1996;



83

37. S.M.Walas, hemimcal process equipment selection and design, 1988.



i

Summary

This design report is the final report of Conceptual Process Design of the plant for the production of 3,3-dimethylbutyraldehyde from the oxidation of 3,3-dimethylbutanol. 3,3-dimethylbutyraldehyde can be wildly used as fragrances in the perfume industry and intermediates in organic synthesis. Several companies are found from the literature to produce 3,3-dimethylbutyraldehyde by using different ways. Therefore the product is currently available in the market. However, it is expected to have a bigger market for this aldehyde. The product will be sold to aldeheyde manufacturers such as NutraSweet Company and its competitors. This novel catalytic method for alcohol oxidation has been developed by Biocatalysis and Organic Chemistry section of Delft University of Technology. The project principal is Dr. I. W. C. E. Arends. It applies new catalyst and different starting raw materials and it is very promising in term of economical potential and yield. It is aimed to design the plant with the annul production scale of 1000 tons 3,3-dimethylbutyraldehyde with the proposed price of 143.5 Euro/kg. This product is formed during the oxidation reaction from 3,3-dimethylbutanol. This new chemical route is carried out using a ruthenium catalyst in combination with TEMPO as co-catalyst to get 3,3-dimethylbutyraldehyd. The feedstock includes 3,3-dimethylbutanol; air, ruthenium catalyst and TEMPO, and all of them are commercially available. The total investment for this plant is determined to be 3.98 million euros per year. The economical life of the plant will be 11 years in which 1 year for design and construction and 10 years of working period. The annual income is expected to be 292.67 million euros and the total production cost is determined to be 290.25 million euros per year. This report is divided into 12 chapters. The introduction chapter provides the conceptual design aspects. The process options and the option chosen are included in chapter 2. The kernel of the design so called Basis of Design forms chapter 3 in which background information such as feedstock, products, wastes, utilities, plant location, costs etc. are covered. The thermodynamic properties such as the equations and values for parameters, reaction equilibrium is written in chapter 4. The “back-bone” of the process such as the Process Flow Schemes (PFS), Batch Cycle Diagram, Process Stream Summary, Mass and Heat balances are covered in chapters 5 and 7. Chapter 6 provides the design of the control system of the plant. Process and equipment design and the resulting equipment data sheets are produced in chapter 8. Chapter 9 deals with the wastes produced by the plant while chapter 10 looks into the environmental and safety aspects. Economic evaluation based on income, investment and operating costs will result in chapter 11. Creativity and group process methods are shown in chapter 12. The conclusions and recommendations are given in the last chapter of this report, chapter 13.



ii

Table of Contents Chapter Title Page

Summary i

Table of Content ii

Chapter 1 Introduction 1

Chapter 2

2.1

2.2

2.3

2.4

Process Options and Selection

Process option and option chosen

Continuous, batch or combination

Stoichiometry and catalysts

Reaction kinetics

3

3

6

6

7

Chapter 3

3.1

3.2

3.2.1

3.2.2

3.2.3

3.2.4

3.3

3.4

Basis of Design

Description of the design

Process definition

Process concept chosen

Block scheme

Thermodynamic properties & Reaction kinetics

List of Pure Component Properties

Basic assumption

Economic margin

9

9

11

11

16

18

19

21

26

Chapter 4

4.1

4.2

4.2.1

4.2.2

4.3

Thermodynamic Properties

Model for vapor/liquid equilibrium

Thermodynamic data

Reaction enthalpy data

Specific heat data

Validation of method

29

29

31

31

31

31

Chapter 5

5.1

5.2

5.3

5.4

5.5

Process Structure& Description

Criteria and selection

Process flow scheme (PFS)

Process stream summary

Utilities

Process yield

33

33

34

35

35

37



iii

Table of Contents (cont’d) Chapter Title Page

Chapter 6

6.1

6.1.1

6.1.2

6.2

Process Control

Reaction selection

Control of dissolving tank (T05)

Control of reactor (R01)

Separation section

38

38

38

38

39

Chapter 7

7.1

7.2

7.3

Mass and Heat Balances

Practical aspects

Balance for total stream

Balance for stream components

42

42

43

45

Chapter 8

8.1

8.2

8.3

8.4

8.5

8.6

8.7

8.8

8.9

8.10

8.11

8.12

8.13

8.14

8.15

8.16

8.17

Equipment Design

Storage tank for raw material

Storage tank for TEMPO

Storage tank for product

Storage tank for deactivated TEMPO

Dissolving tank

Reactor

Gas–liquid separator

Decanter

Flash drum

Evaporator system

Compressor

Heat exchangers

Distillation column (S03)

Distillation column (S04)

Mixing equipment

Pumps and pipes

Equipment data sheets

47

47

47

47

47

48

49

50

50

51

51

52

52

54

55

55

55

57

Chapter 9

9.1

9.2

9.3

Wastes

Introduction of wastes

Classification of direct wastes

Methods for wastes treatment

58

58

58

58



iv

Table of Contents (cont’d) Chapter Title Page

Chapter 10

10.1

10.2

10.3

Health, Safety and Environment

Introduction of safety

Dow Fire and Explosion Index (FEI)

Hazard and Operability study (HAZOP)

60

60

60

63

Chapter 11

11.1

11.2

11.3

11.3.1

11.3.2

11.3.3

11.4

11.5

11.6

11.7

Economy

Capital investment cost

Annual income

Operation costs

Raw materials costs

Utilities costs

Summary of production costs

Gross income, net cash flow and economic criteria

Cost review

Sensitivities

Discussion

66

66

66

67

67

67

68

69

69

70

70

Chapter 12

12.1

12.2

12.3

12.4

12.5

12.6

12.7

12.8

Creativity & Group Process Methods

What is creativity?

Creativity and process design

Why should creativity be utilized in the design?

What kind of creativity is utilized in the design?

Rules and utilize creativity in the design

How to improve creativity during the design

How creativity is utilized in the design?

Group rules

71

71

71

72

72

73

73

74

76

Chapter 13 Conclusions and Recommendations 77

List of Symbols 79

References 81

Appendices A



v

Appendixes

Table of Contents

Number Title Page

Table of Content i

Appendix 1 Block Scheme 1-1

Appendix 2 ASPEN Simulation 2-1

Appendix 3 Process Flow Scheme (PFS) 3-1

Appendix 4

Appendix 4.1

Appendix 4.2

Stream Summary and Utility Summary

Process Stream Summary

Utility Summary

4-1

4-1

Appendix 5

Appendix 5.1

Appendix 5.2

Appendix 5.3

Appendix 5.4

Appendix 5.5

Appendix 5.6

Appendix 5.7

Appendix 5.8

Appendix 5.9

Appendix 5.10

Appendix 5.11

Equipment Design

Design of storage tanks

Design of dissolving tank T05

Design of reactor R01

Gas-liquid separator & Flash drum

Design of decanter

Evaporator design

Design of distillation column

Design of heat exchangers

Design of pipes and pumps

Design of raw material transfer pump P01


5-1

5-1

5-3

5-7

5-11

5-13

5-14

5-18

5-24

5-32

5-42

5-43

Appendix 6

Appendix 6.1

Appendix 6.2

Appendix 6.3

Ecnonomy

Summary of equipment purchase costs

Capital investment cost calculation

Discounted Cash Flow Rate of Return (DCFROR)

6-1

6-1

6-3

6-4

ST4931 Conceptual Process Design Appendix Design of a plant utilizing novel catalytic method for alcohol oxidation CPD-3284

Lin Luo (1113259) Appendix - - Weimin Wang (1118668) Shuang Zhao (1118447) Zhengjie Zhu (1113658)

i

Appendixes

Table of Contents

Number Title Page

Table of Content i

Appendix 1 Block Scheme 1-1

Appendix 2 ASPEN Simulation 2-1

Appendix 3 Process Flow Scheme (PFS) 3-1

Appendix 4

Appendix 4.1

Appendix 4.2

Stream Summary and Utility Summary

Process Stream Summary

Utility Summary

4-1

4-1

Appendix 5

Appendix 5.1

Appendix 5.2

Appendix 5.3

Appendix 5.4

Appendix 5.5

Appendix 5.6

Appendix 5.7

Appendix 5.8

Appendix 5.9

Appendix 5.10

Appendix 5.11

Equipment Design

Design of storage tanks

Design of dissolving tank T05

Design of reactor R01

Gas-liquid separator & Flash drum

Design of decanter

Evaporator design

Design of distillation column

Design of heat exchangers

Design of pipes and pumps

Design of raw material transfer pump P01


5-1

5-1

5-3

5-7

5-11

5-13

5-14

5-18

5-24

5-32

5-42

5-43

Appendix 6

Appendix 6.1

Appendix 6.2

Appendix 6.3

Ecnonomy


Capital investment cost calculation


6-1

6-1

6-3

6-4

ST4931 Conceptual Process Design Chapter 1 Design of a plant utilizing novel catalytic method for alcohol oxidation CPD-3284


1

Chapter 1 Introduction The catalytic oxidation of primary and secondary alcohols affords aldehydes and ketones, respectively. The aldehydes can be wildly used as fragrances in the perfume industry, intermediates in organic synthesis. While, on the other hand, ketones can be served as intermediates in organic synthesis. Some ketones, especially aliphatic and cyclic ones with ten carbon atoms or less, are also used as flavors and fragrances. Several possible industrial routes can obtain these interesting chemicals. As indicated in the literature, primary and secondary alcohols are often used as starting materials for the production of aldehydes and ketones, respectively. The goal is the development of a widely applicable, environmentally friendly, chemoselective catalytic method for the oxidation of alcohols. The combination of RuCl2(PPh3)3 and TEMPO affords an efficient catalytic system for the aerobic oxidation of a variety of primary and secondary alcohols, giving the corresponding aldehydes and ketones in > 95% selectivity in all cases. The scope of this promising RuCl2(PPh3)3/TEMPO system is the aerobic oxidation of a broad range of alcohols, e.g. aliphatic, benzylic, allylic and cyclic. 3,3-dimethylbutyraldehyde can be produced from oxidation reaction of 3,3-dimethylbutanol using this catalytic system, and wildly used as fragrances in the perfume industry and intermediates in organic synthesis. A novel catalytic method for alcohol oxidation has been developed by Biocatalysis and Organic Chemistry section of Delft University of Technology. The project principal is Dr. I. W. C. E. Arends. The oxidation reaction is carried out using a ruthenium catalyst in combination with TEMPO as co-catalyst to get corresponding aldehyde or ketone from alcohol. The substrate (alcohol) serves as solvent, and 8% oxygen is used as the oxidizing agent to stay below explosion limits at all the time. Hence 10 bars for pressure is chosen in order to reach a reasonable stream of oxygen. This is a new chemical route, which applies new catalyst and different starting raw materials and it is very promising in term of economical potential and yield compared to the existed route mentioned in this chapter (refer to part 2.1.1 for the comparison of these routes). The objective of the project is to design a plant for a production of 3,3-dimethylbutyraldehyde from the oxidation of 3,3-dimethylbutanol. The production scale of the plant is to produce 1000 tons of this aldehyde per year. The supplied form of the product is liquid in 100L containers and the product will be sold to aldeheyde manufacturers such as NutraSweet Company and its competitors. It is expected to have a market price of 143.5 euro/kg of this product [11]. From the literature research, several companies are found to produce 3,3-dimethylbutyraldehyde by using different ways from what we used. Therefore the product is currently available in the market. However, it is expected to have a bigger market for this aldehyde. Since no toxic substance is used in process of alcohol oxidation, the process of producing aldehydes from its corresponding alcohols is obviously a green process. Therefore, from both environmental and economic point of view, the process is sustainable. The manufacturers will be interested in these aldehydes if they can be purchased at an acceptable price. Therefore, they can not only still make a lot of profits but also reduce the harm in making these products. The feed of the plant includes 3,3-dimethylbutanol; air, ruthenium catalyst and TEMPO, and all of them are commercially



2

available. After searching a lot of information of the design project, we did not find any similar plant from the literature. 3,3-dimethylbutyraldehyde is chosen to be product in liquid form. The normal boiling point of this compound is 142oC (refer to chapter 4 of thermodynamic properties and table 3.4 for list of pure component physical and chemical properties). The feed of the plant includes 3,3-dimethylbutanol; air, ruthenium catalyst and TEMPO, and all of them are commercially available. After searching a lot of information of the design project, we did not find any similar plant from the literature. Many difficulties are encountered during the design process. For instance, the information in literature on thermodynamic data and pure component data is lacked; the price of the raw materials is currently unavailable. Therefore, many assumptions, educated guesses are made during the design based on existing data and information. Some data of physical properties is assumed based on the results from the ASPEN plus simulation. The wastes produced by the process include the off-gas, wastewater and deactivated TEMPO. The methods of waste treatment in the design as mentioned in chapter 9 are chosen to ensure the process environmentally safe. Dow Fire and Explosion Index (FEI) and Hazard and Operability study (HAZOP) are used to assess safety aspects from a process design point of view. The most dangerous unit, which is reactor and the most dangerous component (3,3-dimethylbutanol) are taken for the determination of the degree of hazard in the plant. It is expected that this product, 3,3-dimethylbutyraldehyde, will be able to compete in the market of fragrances in the perfume industry and intermediates in organic synthesis. It is believed that the plant in the design is very safe for investment and very promising.



3

Chapter 2 Process Options and Selection In this chapter, a comparison of different chemical routes is firstly made so that the advantage of the method used can be seen easily. Secondly, the best process option is chosen after consideration of several options. Then the reason for choosing continuous process, stoichiometry and catalyst, and reaction kinetics are present. 2.1 Process option and option chosen 2.1.1 Chemical routes a) Chemical route 1 This invention relates to a method for preparing 3,3-dimethylbutyraldehyde comprising the step of isomerizing vaporized 3,3-dimethyl-1, 2-epoxybutane in the presence of silica gel. The invention also relates to the above-described method further comprising the step of oxidizing 3,3-dimethylbutene to form 3,3-dimethyl-1, 2-epoxybutane prior to the step of isomerization. Yet another embodiment of this invention relates to the method of preparing 3,3-dimethyl -1,2-epoxybutane by treating 3,3-dimethylbutene with dimethyldioxirane. The method of this invention allows for the preparation of 3,3-dimethylbutyraldehyde in a reproducible and highly economical manner so that use of the aldehyde in the preparation of a sweetener derived from aspartame is commercially practicable.

(1)

b) Chemical route 2 This invention is directed to a method for preparing 3,3-dimethylbutyraldehyde from 3,3-dimethylbutanol using an oxidizing component described in more particularity below. In one embodiment, 3,3-dimethylbutanol is oxidized to 3,3-dimethylbutyraldehyde in the vapor phase by contacting it with an oxidizing metal oxide compound. In another embodiment, the oxidation of 3,3-dimethylbutanol is carried out by treating it with 2,2,6,6-tetramethyl-1-piperidinyloxy, free radical and an oxidizing agent in a solvent to produce 3,3-dimethylbutyraldehyde. The method of this invention provides a commercially practicable means of preparing 3,3-dimethylbutyraldehyde.

(2)

(3)



4

c) Chemical route 3 The RuCl2(PPh3)3/TEMPO catalysed aerobic oxidation is avalialbe for a broad range of primary and secondary alcohols[2], e.g. aliphatic, ebnzylic, allylic and cyclic, canbe smoothly oxidized to their corresponding aldehydes and ketones. In our process, 3,3-dimethylbutyraldehyde is manufactured in this way. We use Ru-100, combined with TEMPO as catalyst. The oxygen, instead of metal oxidite, is chosen as oxidizing agent. This is the reaction.

(4)

d) Chemical route 4 To explore economical synthetic routes to 3,3-dimethylbutyraldehyde the group at the NutraSweet Company has developed a process that involves oxidation of 1-chloro-3, 3-dimethylbutane with DMSO in the presence of a base and substoichiometric amounts of MX (M=Na, K; X=Br, I). In addition they report the purification of the aldehyde via a bisulphate adduct [27].

(5)

e) Criteria for choosing the chemical routes Table 2.1: Criteria for choosing the chemical routes

Chemical routes 1 2 3

Raw material availability

3,3-dimethyl-1, 2-epoxybutane is not

easy to obtain. It is generated by

treating 3,3-dimethylbutene

with dimethyldioxirane.

Easy

Easy

Operating conditions

High temperature (200 – 400˚C) High pressure

(60 bar)

High temperature

(350 ˚C)

Low temperature (80 – 100 ˚C) Low pressure

(10 bar)

Equipments Parr reactor is

placed in the dry ice bath

A metal column packed with the oxidizing metal

oxide

No special equipment

Over-Oxidation problem

Use the temperature

control to preclude over-oxidation

30% of acid was formed along with the 3,3-

dimethylbutyraldehyde.

No



5

Toxic Material None KBr None Yield (%) 70 90 96 Decision Rejected Rejected Chosen

According to Table 2.1, we can see route 3 overwhelmed the other two routes for many aspects of advantage, whereby it is chosen as our production process. 2.1.2 Options chosen a) Option 1 (O2/N2 feeding system) Feed oxygen is an oxidant. However, it couldn’t enter the reactor directly because of the explosion limit. According to the dissertation [4], the volume ratio between oxygen and nitrogen is 8:92. When the gases out of the reactor, two ways can be considered to deal with it. Two measures are expected to guarantee the inlet oxygen composition satisfy the above requirement. First, put a nitrogen production device ahead of the reactor (See Appendix Block Scheme). This device separate oxygen from atmosphere air and provide 99% pure of nitrogen. Meanwhile, another stream of air bypass this device. By controlling the flow rate of these two streams, the outlet stream with 8%(v/v) oxygen composition can be made. Alternatively, the ‘used’ air from the reactor can be recycled to ‘dilute’ the oxygen in the feed air. The reactor itself is a nitrogen produce machine. This proposal shouldn’t be adopt unless it’s verified by serious mathematics. Fortunately, after simulating by ASPEN PLUS, this method is tested and succeeded. The follow table illustrates the advantages of this method Table 2.2: Evaluation of option 1

Option Advantages Disadvantages DecisionWith gas recycle

No need for nitrogen product device

Simple Economical

Difficult to control the concentration of oxygen at unsteady state.

Chosen

Without gas recycle

Be able to control the oxygen concentration at both steady and unsteady state.

A nitrogen product device needed

Expensive operating cost

Rejected

b) Option 2 (Further separation of TEMPO from reacted mixture) The mixture from the decanter (D01) contains some amount of TEMPO (about 3% m/m). Generally, separating TEMPO from this mixture by distillation is not energy favorable. Also it’s not necessary because most of the TEMPO (90%) is recycled. No issue happens if the recycle contains alcohol, aldehyde. So a flash drum is added to provide a crude separation. Here, two ways can be chosen to the outlet of the flash drum. First, separate Ru-100 catalyst directly and purge the liquid phase out. Secondly, it can be treated with a distillation column again (See the block scheme). The distillate is alcohol rich and the bottom product is TEMPO rich.



6

Table 2.3: Evaluation of option 2 Option Advantages Disadvantages DecisionFurther

separation Save the raw material Increased the

production Less waste purged

Increased capital investment

Complex

Chosen

No further separation

Simple

Considerable amount of alcohol and aldehyde in the waste

Waste raw material

Rejected

Since about 40 tons annually of raw material (alcohol) is wasted if no further distillation added. That will be a huge waste from the cost point of view; especially the price of the raw material is high. Therefore, a distillation to separate TEMPO from the alcohol is needed. 2.2 Continuous, batch or combination The factors that favor batch operation are [4]:

1. Production rate a. Sometimes batch if less than 10×106 lb/yr; b. Usually batch if less than 1×106 lb/yr; c. Multiproduct plant;

2. Market forces a. Seasonal production; b. Short product lifetime;

3. Scale-up problems a. Very long reaction times; b. Handling slurries at low flow rates; c. Rapidly fouling materials.

The reasons for continuous process chosen are as follow: 1. The production scale of this 3,3-dimethylbutyraldehyde plant is 1000 tons/yr

(2.2×106 lb/yr) > 1×106 lb/yr; 2. It is not multiproduct plant; 3. It is not seasonal production and product lifetime is not short; 4. No scale-up problems exist.

Therefore batch process is not necessarily chosen. According to the production scale, continuous process is chosen for plant operation. 2.3 Stoichiometry and catalysts a) Stoichiometry The stoichiometric ratio between alcohols and their corresponding aldehydes or ketones is 1:1. In this process, ratios between 3,3-dimethylbutanol and 3,3-dimethylbutyraldehyde are 1:1.



7

(6)

(7)

From the experiment, if the reaction take place under an inert atmosphere, the aldehyde and TEMPH as products in a 3:2 ratio (reaction 3). On the other hand, in the presence of oxygen TEMPOH is rapidly oxidized to TEMPO. Whereby, TEMPO acts as hydrogen transfer mediator, the reaction of TEMPO might not be considered during our process model.

b) Catalyst Name Dichlorotris (tirphenylphosphine) ruthenium (II) C.A.S No 15529-49-4 Catalog ID Ru-100 Formula RuCl2(PPh3)3 Colour black crystal Precious Metal Content 10.5% Comments Dissociates in solution Name TEMPO (2,2,6,6-tetramethylpiperidinyloxy) C.A.S No 2564-83-2 Formula C9H18NO* 2.4 Reaction kinetics The dependence of the initial rate of catalytic oxidation on the temperature can be employed to determine the activation energy of the reaction (Arrhenius plot). The data for the Ru/TEMPO-catalyzed oxidation of octan-2-ol plot in Figure 2.1 can be readily fitted

to the familiar expression exp( )aEk A

RT

, to give the activation energy ( aE ) of 47.8

kJ/mol.



8

Figure 2.1: The correlation of initial rate and the temperature (40-120 ˚C) for the Ru/TEMPO-catalyzed aerobic oxidation of octan-2-ol

The linear function in the figure is 5757.1

ln( ) ln( ) 8.52aEk A

RT T [2]

Due to the lack data of reaction of oxidation of 3,3-dimethylbutyraldehyde, we use the data of octan-2-ol combined with several specifications:

1st order reaction; Reaction condition is 10 bar, 100 ˚C; The conversion is 99% after 7 hours reaction time; the 1% byproduct is acid

due to the over-oxidation. The reaction used in the process is 30 minutes, and the conversion is 23%.



9

Chapter 3 Basis of Design 3.1. Description of the design A novel catalytic method for alcohol oxidation has been developed by Biocatalysis and Organic Chemistry section of Delft University of Technology. The project principal is Dr. I. W. C. E. Arends. The oxidation reaction is carried out using a ruthenium catalyst in combination with TEMPO as co-catalyst to get corresponding aldehyde or ketone from alcohol. The substrate (alcohol) serves as solvent, and 8% oxygen is used as the oxidizing agent to stay below explosion limits at all the time. Hence 10 bars for pressure is chosen in order to reach a reasonable stream of oxygen. This is a new chemical route, which applies new catalyst and different starting raw materials and it is very promising in term of economical potential and yield compared to the existed route mentioned in this chapter. The catalytic oxidation of primary and secondary alcohols affords aldehydes and ketones, respectively. The aldehydes can be wildly used as fragrances in the perfume industry, intermediates in organic synthesis. While, on the other hand, ketones can be served as intermediates in organic synthesis. Some ketones, especially aliphatic and cyclic ones with ten carbon atoms or less, are also used as flavors and fragrances. Several possible industrial routes can obtain these interesting chemicals. As indicated in the literature, primary and secondary alcohols are often used as starting materials for the production of aldehydes and ketones, respectively. The oxidation of alcohols into their corresponding aldenhydes and ketones is of significant importance in synthetic organic chemistry. Many reagents for alcohol oxidations are known, e.g. hypochlorite, chromium (VI) oxide, dichromate, manganese (IV) oxide, permanganate and ruthenium (VIII) oxide. Unfortunately, one or more equivalents of these – often hazardous or toxic – oxidizing agents are required. According to the literature, one of the options to produce aldehydes and ketones is simple dehydrogenation over a solid metal catalyst producing the desired aldehyde/ketone and one equivalent of hydrogen gas. However, rather high temperatures (>250°C) are required for this process. On the other hand, lower temperatures can be employed in oxidation processes. Already in 1984, Semmelhak et al. report a CuCl/TEMPO catalytic system, which was capable of oxidizing alcohols to the corresponding aldehydes and ketones using molecular oxygen as terminal oxidant. The major disadvantage of this system was that only activated benzylic and allylic alcohols could be oxidized. Moreover, DMF and large amounts of catalyst were required to obtain good activity. From both an economic and environmental point of view, the quest for effective catalytic oxidation processes that use clean, inexpensive primary oxidants, such as molecular oxygen or hydrogen peroxide, e.g. a ‘green method’ for converting alcohols to carbonyl compounds on an industrial scale, remains an important challenge. The goal is the development of a widely applicable, environmentally friendly, chemoselective catalytic method for the oxidation of alcohols. The combination of



10

RuCl2(PPh3)3 and TEMPO affords an efficient catalytic system for the aerobic oxidation of a variety of primary and secondary alcohols, giving the corresponding aldehydes and ketones in > 95% selectivity in all cases. The scope of this promising RuCl2(PPh3)3/TEMPO system is the aerobic oxidation of a broad range of alcohols, e.g. aliphatic, benzylic, allylic and cyclic. The objective of the project is to design a plant that uses the novel catalytic method mentioned above and select the oxidation route for which the novel catalytic method will lead to a substantial improvement in ‘greenness’ and economy. The basis of the design is to select the capacity of the plant and the product specifications based on the product selected. As a low calorie, extremely potent sweetening sweetener, Neotame, N-[3,3-dimethylbutyl-L-α-aspartyl]-L-phenylalanine-1-methyl ester, would be extensively used in the food and beverage industries instead of Aspartame or other sweetener. 3,3-dimethylbutyraldehyde is one of two raw materials in the process of producing neotame. H2, Cd/C 3,3-dimethylbutyraldehyde + aspartame neotame ( reductive amination) Therefore after choosing the process and product from fine chemical industry, 3,3-dimethylbutyraldehyde is decided to be the product of the design process. Since 3,3-dimethylbutyraldehyde can be wildly used as fragrances in the perfume industry and intermediates in organic synthesis, the market demand is expected to be very high. The production scale of the plant is to produce 1000 tons of this aldehyde per year. The supplied form of the product is liquid in 100L containers and the product will be sold to aldeheyde manufacturers such as NutraSweet Company and its competitors. It is expected to have a market price of 143.5 euro/kg of this product [11]. Since no toxic substance is used in the process, it is obviously a green process. Therefore, from both environmental and economic point of view, the process is sustainable. And from safety point of view, 8% oxygen was chosen as oxygen-source to stay below explosion limits at all the time. From the literature research, several companies are found to produce 3,3-dimethylbutyraldehyde by using different ways from what we used. For example, the NutraSweet Company have developed a process that involves oxidation of 1-chloro-3,3-dimethylbutane with DMSO in the presence of a base and substoichiometric amounts of MX (M=Na, X=Br, I) [27]. Therefore the product is currently available in the market. However, it is expected to have a bigger market for this aldehyde. As mentioned above that no toxic substance is used in the process of alcohol oxidation, the process of producing aldehydes from its corresponding alcohols is a green process from environmental point of view. The manufacturers will be interested in these aldehydes if they can be purchased at an acceptable price. Therefore, they can not only still make a lot of profits but also reduce the harm in making these products. The feed of the plant includes 3,3-dimethylbutanol; air, ruthenium catalyst and TEMPO, and all of them are commercially available. After searching a lot of information of the design project, we did not find any similar plant from the literature.



11

Many difficulties are encountered during the design process. For instance, the information in literature on thermodynamic data and pure component data is lacked; the price of the raw materials is currently unavailable. Therefore, many assumptions, educated guesses are made during the design based on existing data and information. Some data of physical properties is assumed based on the results from the ASPEN plus simulation. This report presents the process options and option chosen. The block schemes, thermodynamic properties, reaction kinetics, pure component properties, process stream summary and basic assumptions are also included. Margin as the total value of products less the total value of feedstock is calculated and presented as well. 3.2. Process definition In this section, a comparison of different chemical routes is firstly made so that the advantage of the method used can be seen easily. Secondly, the best process option is chosen after consideration of several options. Then the reason for choosing continuous process, stoichiometry and catalyst, reaction kinetics, block scheme, thermodynamic properties, list of pure component properties, process stream summary, and mass balance are present. 3.2.1 Process concept chosen 1) Chemical routes a) Chemical route 1 This invention relates to a method for preparing 3,3-dimethylbutyraldehyde comprising the step of isomerizing vaporized 3,3-dimethyl-1, 2-epoxybutane in the presence of silica gel. The invention also relates to the above-described method further comprising the step of oxidizing 3,3-dimethylbutene to form 3,3-dimethyl-1, 2-epoxybutane prior to the step of isomerization. Yet another embodiment of this invention relates to the method of preparing 3,3-dimethyl -1,2-epoxybutane by treating 3,3-dimethylbutene with dimethyldioxirane. The method of this invention allows for the preparation of 3,3-dimethylbutyraldehyde in a reproducible and highly economical manner so that use of the aldehyde in the preparation of a sweetener derived from aspartame is commercially practicable.

(1)

b) Chemical route 2 This invention is directed to a method for preparing 3,3-dimethylbutyraldehyde from 3,3-dimethylbutanol using an oxidizing component described in more particularity below. In one embodiment, 3,3-dimethylbutanol is oxidized to 3,3-dimethylbutyraldehyde in the vapor phase by contacting it with an oxidizing metal oxide compound. In another



12

embodiment, the oxidation of 3,3-dimethylbutanol is carried out by treating it with 2,2,6,6-tetramethyl-1-piperidinyloxy, free radical and an oxidizing agent in a solvent to produce 3,3-dimethylbutyraldehyde. The method of this invention provides a commercially practicable means of preparing 3,3-dimethylbutyraldehyde.

(2)

(3)

c) Chemical route 3 The RuCl2(PPh3)3/TEMPO catalysed aerobic oxidation is avalialbe for a broad range of primary and secondary alcohols[2], e.g. aliphatic, ebnzylic, allylic and cyclic, canbe smoothly oxidized to their corresponding aldehydes and ketones. In our process, 3,3-dimethylbutyraldehyde is manufactured in this way. We use Ru-100, combined with TEMPO as catalyst. The oxygen, instead of metal oxidite, is chosen as oxidizing agent. This is the reaction.

(4)

d) Chemical route 4 To explore economical synthetic routes to 3,3-dimethylbutyraldehyde the group at the NutraSweet Company has developed a process that involves oxidation of 1-chloro-3,3-dimethylbutane with DMSO in the presence of a base and substoichiometric amounts of MX (M=Na, K; X=Br, I). In addition they report the purification of the aldehyde via a bisulphate adduct [27].

(5)



13

e) Criteria for choosing the chemical routes Table 3.1: Criteria for choosing the chemical routes

Chemical routes 1 2 3

Raw material availability

3,3-dimethyl-1, 2-epoxybutane is not

easy to obtain. It is generated by

treating 3,3-dimethylbutene

with dimethyldioxirane.

Easy

Easy

Operating conditions

High temperature (200 – 400˚C) High pressure

(60 bar)

High temperature

(350 ˚C)

Low temperature (80 – 100 ˚C) Low pressure

(10 bar)

Equipments Parr reactor is

placed in the dry ice bath

A metal column packed with the oxidizing metal

oxide

No special equipment

Over-Oxidation problem

Use the temperature

control to preclude over-oxidation

30% of acid was formed along with the 3,3-

dimethylbutyraldehyde.

No

Toxic Material None KBr None Yield (%) 70 90 96 Decision Rejected Rejected Chosen

According to Table 2.1, we can see route 3 overwhelmed the other two routes for many aspects of advantage, whereby it is chosen as our production process. 2) Options chosen a) Option 1 (O2/N2 feeding system) Feed oxygen is an oxidant. However, it couldn’t enter the reactor directly because of the explosion limit. According to the dissertation [4], the volume ratio between oxygen and nitrogen is 8:92. When the gases out of the reactor, two ways can be considered to deal with it. Two measures are expected to guarantee the inlet oxygen composition satisfy the above requirement. First, put a nitrogen production device ahead of the reactor (See Appendix Block Scheme). This device separate oxygen from atmosphere air and provide 99% pure of nitrogen. Meanwhile, another stream of air bypass this device. By controlling the flow rate of these two streams, the outlet stream with 8%(v/v) oxygen composition can be made. Alternatively, the ‘used’ air from the reactor can be recycled to ‘dilute’ the oxygen in the feed air. The reactor itself is a nitrogen produce machine. This proposal shouldn’t be adopt unless it’s verified by serious mathematics. Fortunately, after simulating by ASPEN PLUS, this method is tested and succeeded. The follow table illustrates the advantages of this method



14

Table 3.2: Evaluation of option 1 Option Advantages Disadvantages Decision

With gas recycle

No need for nitrogen product device

Simple Economical

Difficult to control the concentration of oxygen at unsteady state.

Chosen

Without gas recycle

Be able to control the oxygen concentration at both steady and unsteady state.

A nitrogen product device needed

Expensive operating cost

Rejected

b) Option 2 (Further separation of TEMPO from reacted mixture) The mixture from the decanter (D01) contains some amount of TEMPO (about 3% m/m). Generally, separating TEMPO from this mixture by distillation is not energy favorable. Also it’s not necessary because most of the TEMPO (90%) is recycled. No issue happens if the recycle contains alcohol, aldehyde. So a flash drum is added to provide a crude separation. Here, two ways can be chosen to the outlet of the flash drum. First, separate Ru-100 catalyst directly and purge the liquid phase out. Secondly, it can be treated with a distillation column again (See the block scheme). The distillate is alcohol rich and the bottom product is TEMPO rich. Table 3.3: Evaluation of option 2

Option Advantages Disadvantages DecisionFurther

separation Save the raw material Increased the

production Less waste purged

Increased capital investment

Complex

Chosen

No further separation

Simple

Considerable amount of alcohol and aldehyde in the waste

Waste raw material

Rejected

Since about 40 tons annually of raw material (alcohol) is wasted if no further distillation added. That will be a huge waste from the cost point of view; especially the price of the raw material is high. Therefore, a distillation to separate TEMPO from the alcohol is needed. 3) Continuous, Batch or Combination The factors that favor batch operation are [4]:

1. Production rate a. Sometimes batch if less than 10×106 lb/yr; b. Usually batch if less than 1×106 lb/yr; c. Multiproduct plant;

2. Market forces



15

a. Seasonal production; b. Short product lifetime;

3. Scale-up problems a. Very long reaction times; b. Handling slurries at low flow rates; c. Rapidly fouling materials.

The reasons for continuous process chosen are as follow: 1. The production scale of this 3,3-dimethylbutyraldehyde plant is 1000 tons/yr

(2.2×106 lb/yr) > 1×106 lb/yr; 2. It is not multiproduct plant; 3. It is not seasonal production and product lifetime is not short; 4. No scale-up problems exist.

Therefore batch process is not necessarily chosen. According to the production scale, continuous process is chosen for plant operation. 4) Stoichiometry and Catalysts a) Stoichiometry The stoichiometric ratio between alcohols and their corresponding aldehydes or ketones is 1:1. In this process, ratios between 3,3-dimethylbutanol and 3,3-dimethylbutyraldehyde are 1:1.

(6)

(7)

From the experiment, if the reaction take place under an inert atmosphere, the aldehyde and TEMPH as products in a 3:2 ratio (reaction 3). On the other hand, in the presence of oxygen TEMPOH is rapidly oxidized to TEMPO. Whereby, TEMPO acts as hydrogen transfer mediator, the reaction of TEMPO might not be considered during our process model.

b) Catalyst Name Dichlorotris (tirphenylphosphine) ruthenium (II) C.A.S No 15529-49-4 Catalog ID Ru-100 Formula RuCl2(PPh3)3



16

Colour black crystal Precious Metal Content 10.5% Comments Dissociates in solution Name TEMPO (2,2,6,6-tetramethylpiperidinyloxy) C.A.S No 2564-83-2 Formula C9H18NO* 5) Reaction kinetics The dependence of the initial rate of catalytic oxidation on the temperature can be employed to determine the activation energy of the reaction (Arrhenius plot). The data for the Ru/TEMPO-catalyzed oxidation of octan-2-ol plot in Figure 2.1 can be readily fitted

to the familiar expression exp( )aEk A

RT

, to give the activation energy ( aE ) of 47.8

kJ/mol.

Figure 3.1: The correlation of initial rate and the temperature (40-120 ˚C) for the Ru/TEMPO-catalyzed aerobic oxidation of octan-2-ol

The linear function in the figure is 5757.1

ln( ) ln( ) 8.52aEk A

RT T [2]

Due to the lack data of reaction of oxidation of 3,3-dimethylbutyraldehyde, we use the data of octan-2-ol combined with several specifications:

1st order reaction; Reaction condition is 10 bar, 100 ˚C; The conversion is 99% after 7 hours reaction time; the 1% byproduct is acid

due to the over-oxidation. The reaction used in the process is 30 minutes, and the conversion is 23%.

3.2.2 Block scheme 1) Block Scheme (See Appendix A1) 2) Description of block scheme The process of manufacturing 3.3-dimethylbutyraldehyde by oxidation of 3,3-Dimethyl-1-butanol works as described according to the following steps:



17

All the raw materials (except air) are in the storage area. Feedstock (98.5% 3,3-Dimethyl-1-butanol) and TEMPO are in the different storage tanks. Ru-100 is in the catalyst cans. The air is blasted into the system.

Since Ru-100 is solid at normal temperature, it should be dissolved in tempo before entering the reactor. TEMPO is also heated above its melting point (37 ˚C, 1atm). The molar ratio of Ru-100/TEMPO is 1:3.

Before each batch the reactors, the storage tanks for raw materials are cleaned according to Good Manufacturing Practices. Feed oxygen is an oxidant. According to the dissertation, the volume ratio between oxygen and nitrogen is 8:92. Therefore, the air needs to be diluted with Nitrogen. To do this, a vapor phase recycle is added to ‘dilute’ the oxygen composition in the feed. (See the block scheme). In this case, the reactor can be regarded as a nitrogen produce device. By controlling the flow rate of stream, the reactor can be fed with the gas mixture contain 8%(v/v) of oxygen which meet the specification.

In the reactor (block R01), the oxidation reaction takes place at 100 ˚C, 10 bar. Several specifications for this reaction are: the molar ratio of substrates and catalyst is 67:1. After reaction time of 1 hour, the conversion is 21% and the selectivity exceeds 99%. Since oxygen takes part in the reaction, this reaction is exothermic, so the reactor is cooled down using cooling water to maintain the temperature at 100 ˚C.

After the reactor, the reacted mixture release from the top of the reactor, mixture mainly includes nitrogen, oxygen, alcohol, aldehyde and TEMPO. This mixture is vapor-liquid phase. A flash drum is used to separate these two phases. The top product is mainly oxygen and nitrogen which is fed back to the reactor. The bottom liquid mixture then goes into a decanter. Because organic components doesn’t dissolve in the produced water (from oxidizing reaction) , the water can be easily discharged by the decanter. The mixture from the decanter (D01) contains some amount of TEMPO (about 3% m/m). Generally, separating TEMPO from this mixture by distillation is not energy favorable. Also it’s not necessary because most of the TEMPO (90%) is recycled. No issue happens if the recycle contains alcohol, aldehyde. Moreover, the boiling point of Ru-100/TEMPO (more than 500K) is much higher than alcohol and aldehyde. So a flash drum is added to provide a crude separation. After the flash drum, the outlet can be treated with a distillation column (S04) again. The distillate is alcohol rich and the bottom product is TEMPO rich. Ru-100/TEMPO passes through block S05 for a further separation. In this unit, the deactivated Ru-100 is separated. This deactivated Ru catalyst is disposed. Fortunately, according to Johnson Matthery Inc, we can exchange the deactivated catalyst for flesh one with them, a half-price charged.



18

The liquid from S02 combined with the distillate from distillation column S04 goes to aldehyde column S03.where the 3,3-Dimethyl-1-butanol is separated as recycle, the top product is 3.3-dimethylbutyraldehyde, the purity is above 99%.

3.2.3 Thermodynamic properties & reaction kinetics (Please also refer to chapter 4 for more information on thermodynamic properties) 1) Model for vapor/liquid equilibrium Phase equilibrium data are needed for the design of all separation processes. Experimental data have been published for several thousands binary and many multi-component systems. But no universal equation is available for computing the thermodynamic properties of various kinds of mixture. Instead, there are three types of models to calculate phase equilibrium: the equation-of-state method, the activity coefficient method and a special application method. These are based on constitutive equations because they depend on the constitution or nature of the components in the mixture. First of all, the equation-of-state method is used to describe both liquid and vapor phase behavior. This type is recommended to apply for weak non-ideal solution, such as most hydrocarbon and light gas mixture systems at high and moderate pressure (at least not below atmospheric pressure). This method is applicable for systems where the interaction of the components in the liquid phase is assumed to minimize. Because the design system is a polar, non-ideal solution, comprising of alcohol, aldehyde and water, under the condition of low pressure and moderate temperature, the equation-of-state method is not applicable. Secondly, the activity coefficient method is a combination of two different equations: one to calculate vapor phase properties (j

v, EOS/IG), methods Nothnagel and Hayden-O’Connell with Henry are effective. The other one is to describe liquid phase behavior (j in the unsymmetrical form). This method is widely used in the highly non-ideal mixtures, which are comprised of acetone, alcohol, ester, ether, etc. Especially, high accuracy would be obtained for (chemical) systems with dissimilarity in size, shape and intermolecular forces. Since the chemicals involved in the system are raw material, 3,3-dimethylbutanol with impurity, n-hexane and 1-hexene, and products, 3,3-dimethylbutyraldehyde and water, it is a non-ideal mixture with intermolecular forces, the activity coefficient method should be applicable to model vapor/liquid equilibrium. Thirdly, the special application method is available for systems that for some reasons cannot easily fall into one of the other two methods. In the activity coefficient method, typical equations model liquid phase behaviors are:

Wilson NRTL Uniquac Van Laar UNIFAC

Wilson and Van Laar are only suitable for binary vapor-liquid systems which interaction is small, while NRTL and Uniquac can also be used in cases that there are two liquid



19

phases present. NRTL and Uniquac have better accurate in systems containing water than Wilson, however contrary result is showed for vapor-liquid system. None of these models should be used under critical conditions. A special, widely used activity coefficient model is UNIFAC. This method can predict phase equilibrium for systems for which no experimental data are available. It doesn’t depend on binary interaction parameters, instead bases upon the so-called group contribution method. And a modified NRTL method, NRTL-HOC is used Hayden-O’Connell method to calculate fugacity coefficient. Since the thermodynamic data for raw material-3, 3-dimethylbutanol and 3,3-dimethybutyraldehyde are hardly available in database, even in ASPEN Simulation engine; therefore UNIFAC and NRTL-HOC are chosen to predict vapor/liquid equilibrium for this system. This design aimed to separate the product (3,3-dimethylbutyraldehyde) from the others (mainly water and raw material). Because the heaviest component is the alcohol and the volatile key is water since it has the lowest boiling point among the others. T/x @p constant diagrams are produced for the key components and can be viewed in chapter 4 (figure 4.1 and 4.2). The graphs are obtained from ASPEN PLUS simulation engine. From these graphs, it can be seen that the separation of alcohol from mixture in the bottom or separation of product from water can be done by distillation. 2) Thermodynamic data a) Reaction enthalpy data The reaction enthalpy is calculated from the heat of formation of each component at reaction temperature 100C (Note: Data of heat of formation at 100C are taken from ASPEN simulation engine and Henry’s chemical engineer handbook). Table 4.1 (chapter 4) presents heat of formation of the components at 100oC. From this table, it is determined that the reaction is exothermic (please refer to chapter 4 for explanation) b) Specific heat data Typical constant pressure specific heat of each component can be obtained from ASPEN PLUS simulation engine. Specific heat of the components at 1atm, 25oC (obtained from ASPEN) is given in table 4.2 (chapter 4). 3) Validation of method For separations Carlson (1996) suggests to check boiling points of the pure components against values from literature. As it can be seen from the table 4.3 (chapter 4), the boiling point of each component from UNIFAC model is higher than literature and average error is + 3.1 %. 3.2.4 List of pure component properties



20

Table 3.4: List of pure component physical and chemical properties

PURE COMPONENT PROPERTIES Component Name Technological Data Metdical Data Notes

Design systematic Formula CAS RN Mol. Weight g/mol

Boiling point K

Melting point 0C

Density of liquid kg/m3

Viscosity of liquid CP/ 0C

Solubilityin water % (wt)

MAC value Mg/m3

LD50 Oral g

Hexane n-hexane

C6H14 110-54 -

3 86.18 341.7 -95 0.6548

hexane

1-hexene C6H12 592-41-6

84.16 336.6 -139.8 0.673

Neo-hexanol

3,3-dimethyl-butanol

C6H14O 624-95-3

102.18 415.15 -60 0.817 4.5782.271 0.781

0.60 0.51 0.44

250C 500C 1000C

Neo-hexanal

3,3-dimethylbutyraldehyde

C6H12O 2987-16-8

100.16 379.15 0.798 0.583 0.429 0.262

1.75 1.40 1.12

250C 500C 1000C

TEMPO 2,2,6,6-(Tetramethylpiperidino) --oxy

C9H18NO 2564-83-2

156.25 535.15 36-39 n.a

Ru (II) catalyst

Ru (II) catalyst RuCl2

(PPh3)3 15529-49-4

958.85 - 159 n.a



21

3.3. Basic assumptions 3.3.1 Plant capacity The production scale is 1000 tons/year of 3,3-dimethylbutyraldehyde. The reaction part and the separation part will be operated continuously the whole year. 1) Feedstock The feed of the plant are 3,3-dimethylbutanol, Air, Ru-100 and TEMPO. a) Feed condition Table 3.5: Feed condition Pressure (bar) 10 Temperature (°C) 100

b) Feed composition Table 3.6: Composition for stream <1> Stream name: Feedstock <1> Components Specification (%) Notes Additional Information

(Also ref. Note numbers) Available Design 3,3-dimethyl-

butanol Hexane Hexene

Heavy ends Light ends

Metal

90-99

<1.0 <1.0

<0.004

99.5

0.30.2

(2) (2) (2) (1)

(1) The price of the feedstock is 24.4 euro/kg. The supply phase is liquid. The compositions taken in consultation with Principal. (2) Impurities not harmful for the process. Compounds not included in mass balance.

Total 100.0

Table 3.7: Composition for stream <3> Stream name: Air <3>

Components Specification (%) Notes Additional Information (Also ref. Note numbers) Available Design

Nitrogen Oxygen Argon

Carbon Dioxide

75.7 23.3 0.93 0.03

79.021.0

(2) (2) (1)

(1) This air stream is used to provide oxygen and nitrogen from air. The supply phase is gas. (2) Components not included in mass balance. Total 100.0



22

Table 3.8: Composition for stream <35> Stream name: TEMPO <35>

Components Specification (%) Notes Additional Information (Also ref. Note numbers) Available Design

2,2,6,6-tetramethyl-

piperidinyl-1-oxy Impurities

99.0

1.0

100.0

(2) (1)

(1) The price of TEMPO is 30.8 euro/kg. The supply phase is solid. (2) Components not included in mass balance.

Total 100.0 Table 3.9: Composition for stream <36> Stream name: Ru-100 <36> Components Specification (%) Notes Additional Information

(Also ref. Note numbers) Available Design RuCl2(PPh3)3

Impurities 99.01.0

100.0 (2)

(1)

(1) The price of Ru-100 is 989.0 euro/kg. The supply phase is solid. (2) Components not included in mass balance.

Total

100.0

2) Product The product of the plant is 3,3-dimethylbutyraldehyde, which is in stream 29 (top of the purification column S03). The purity of the product is 99.25% (as a design parameter). The production scale of the plant is 1000 tons/year. It will be stored in 100L containers and will be sold to aldeheyde manufacturers such as NutraSweet and its competitors. The product could be kept in the storage for the distribution depending on the demands. Table 3.10: Production composition Stream name: Product <29>

Components Weight percentage (%)

Notes Additional Information (Also ref. Note numbers)

3,3-dimethyl-butyraldehyde 3,3-dimethyl-

butanol Hexane Hexene Hexanal Water

Nitrogen Oxygen

99.25

0.08

0.290.160.020.010.170.02

(1) (2)

(1) The product is sold 143.5 euro/kg. The supply phase is liquid. (2) The product from the purification column is collected to the storage tank and then 100L containers of the product will be filled and stored.

Total 100.0



23

3) Waste After about 30 minutes of working time, the catalyst RuCl2(PPh3)3/TEMPO will be deactivated. Part of TEMPO will be disposed as waste, the rest will be recycled. However, RuCl2(PPh3)3 could be recycled, and the price of the deactivated RuCl2(PPh3)3 can be estimated as 692 euro/kg (70% of the original price). After some treatments of the deactivated RuCl2(PPh3)3, it can be sold and used again. Table 3.11: Composition for stream <16> Stream name: Wastewater <16>

Components Weight percentage (%) Notes Additional Information (Also ref. Note numbers)

Water 100.0 (1)

(1) The decanter D01 is used to separate the water and other components.

Total 100.0

Table 3.12: Composition for stream <33> Stream name: Waste TEMPO<33>



2,2,6,6-tetramethyl-piperidinyl-1-oxy

3,3-dimethylbutanol

3,3-dimethyl-butyraldehyde

47.1

52.3

0.6

(1)

(1) This waste stream will be at the top of the evaporator S05.

Total 100.0 Table 3.13: Composition for stream <13> Stream name: Off-gas <13>



Nitrogen Oxygen

3,3-dimethyl-butyraldehyde 3,3-dimethyl-


90.746.611.10

1.19

0.030.020.010.30

(1)

(1) Part of the gases from flash drum S01 is recycled, the rest is considered as waste.

Total 100.0



24

3.3.2 Location The plant should be located close to the feedstock supply and convenient for product distribution. The product is 3,3-dimethyl-butyraldehyde, which can be sold to NutraSweet Company in large amount. Since NutraSweet Company is located in the United States, and its competitors can be found in America as well. On the other hand, the main feedstock, 3,3-dimethylbutanol, hopefully can also be purchased in America. Hence the plant will be built is industrial area in the United States. 3.3.3 Battery limit 1) Units inside the plant Table 3.14: Units inside the plant

Unit Number Additional Information Storage Tank Storage Tank Storage Tank Storage Tank

Dissolving Tank Reactor

Flash Dram Decanter

Flash Dram Evaporator Compressor

Heat exchanger Heat exchanger Heat exchanger Heat exchanger Heat exchanger Heat exchanger Heat exchanger Heat exchanger

Distillation column Distillation column

Mixer/Splitter Pump

Vacuum Pump

T01 T02 T03 T04 T05 R01 S01 D01 S02 S05 C01 H01 H02 H03 H04 H11 H12 H13 H14 S03 S04

M01-07 P01-07

P08

Outside the plant, for feedstock (3,3-dimethylbutanol) For TEMPO storage For product storage For deactivated TEMPO storage For dissolving catalyst RuCl2(PPh3)3 in TEMPO Oxidation reaction taking place For gas – liquid separation To separate water and other substances For vapor – liquid separation To separate TEMPO from deactivated Ru-100 To supply air at 10 bar to the reactor As a cooler for feed to flash dram S01 As a heater for feed to flash dram S02 As a heater in distillation column S03 As a cooler for feed air As a reboiler in distillation column S03 As a condenser in distillation column S03 As a reboiler in distillation column S04 As a condenser in distillation column S04 To separate the product from 3,3-dimethylbutanol To separate the catalysts from other subtances For the continuous mixing of low viscosity fluids For transportations of the substances To create vacuum for evaporator S05

2) Units outside the plant Air is needed for preparation of oxidant, but there is no cost for air. Steam is needed for heating the reactor to 100oC and for the distillation column reboiler. Cooling water for the cooling jacket and for the condenser of the distillation column. Power is needed for the pumps and electricity inside the plant.



25

Table 3.15: Units outside the plant Utility Unit Price (euro/unit) Quantity (unit/year) Cost (euro/year) Steam

Cooling water Electricity

ton ton

kWh

13.640.0450.10

2329211706542333

317639623

54233Notes: (1) Refer to table 5.2 of utilities cost. (2) The prices of utilities are taken from CPD Instruction manual (refer to “Cost Data, WEBCI/DACE”, 18th Edition November 1995). (3) Refer to table 5.3b of utilities summary for quantity of the utilities.

3.3.4 Define in-and outgoing streams 1) Ingoing streams Table 3.16: Summary of ingoing streams

Component Steam Price (euro/kg)

Quantity (tons/year)

Total Cost (×106 euro/year)

Notes

Feedstock 3,3-dimethyl-

butanol Hexane Hexene

Heavy ends Light ends

Metal

<1>

24.4 1123.3 27.4 (1)

Oxidant Air

<3> 0.0 1123.3 0

Catalyst TEMPO Ru-100

<35> <36>

30.8989.0

98.5201.5

3.0

199.3

(2)

Notes: (1) The price of feedstock is estimated from the price by Sigma-Aldrich. (2) The prices of catalysts are estimated from the prices by CPChem Company. And price of recycle RuCl2(PPh3)3 is estimated to 70% of the price of original RuCl2(PPh3)3.



26

2) Outgoing streams Table 3.17: Summary of outgoing streams

Component Steam Price (euro/kg)

Quantity (tons/year)

Total Cost (×106euro/year)

Notes

Product 3,3-dimethyl-butyraldehyde 3,3-dimethyl-


Nitrogen Oxygen

<29>

143.5 1067.7 153.2 (1)

Notes: (1) The price of feedstock is estimated from the price by Sigma-Aldrich. (2) Assume the selectivity of the reaction is 99%.

3.4 Economic margin 3.4.1 Estimation of margin Margin is the total value of product and waste minus that of feedstock and process chemicals. Utilities are not considered in this estimation. The estimation is tabulated in Table 3.18. Table 3.18: Estimation of margin

In Quantity

(tons/year) Price

(euro/kg) Total Cost

(×106 euro/year)

Feedstock (1) 1100.0 24.4 26.8

RuCl2(PPh3)3 (2) 109.4 989.0 108.2

TEMPO (3) 53.7 30.8 1.7

Total raw materials cost (4) = (1) + (2) + (3) 136.7

Out

3,3-dimethylbutyraldehyde (5) 1043.3 143.5 149.7

Recycle RuCl2(PPh3)3 (6) 109.4 692.0 54.0

Total products & waste value (7) = (5) + (6) 203.7

Margin = Out - In 67.0



27

Remarks: 1. The price of catalyst is estimated from the price of CPChem Company and

modified by the large quantity factor. And price of recycle catalyst (deactivated RuCl2(PPh3)3) is estimated to 70% of price of active RuCl2(pph3)3

2. The price of raw material is estimated on the basis of the price from Sigma-Aldrich co. and modified by the large quantity factor.

3. Only cost of raw materials, product and waste values are considered here. Assume the recycle rate of catalyst is 90%, and selectivity of the reaction is 99%. From the Table 4.1, note that the cost for chemical catalyst (RuCl2(PPh3)3 ) counts for the large proportion of total cost (79%). Loss of activity of the catalyst should be extremely avoided. From the estimated margin, the process is highly feasible in practice. 67.0 million euro/year or 670.0 millions for project life of 10 working years. But the margin calculated in this chapter did not include the costs of utilities, labor, shipping and packaging, insurance etc. and also the interest was not considered. 3.4.2 Determination of maximum allowed investment The following assumptions are made to determine maximum allowed investment:

The economic plant life of project is 11 years The design/building period is 1 year and the working period is 10 years The productivity of the last year of this project is still 1000 ton/year. The interest rate (r) is 8% per year The margin is 67.0 millions Euro as Table 4.1

The following formula are used to calculate maximum allowed investment:

a) Margin

Present margin =(1 )nr

b) n=11

n=2

Margin Cumulative of margin in working years =

(1 )nr

Where n is year, r is the year interest rate

Cumulative net cash flow at the end of projectDCFROR =

Life of project × Original investment

Cumulative of margin in working years - Original investmentDCFROR =


As DCFROR is 10% or 0.1:



28

Cumulative of margin in working years - Original investment0.1


Cumulative of margin in working years

Original investment = 0.1× Life of project + 1

Original investment = 416.3/ 2.2 = 189.2 million Euro Maximum acceptable investment is determined to be 189.2 millions Euro at DCFROR (Discount Cash Flow rate of Return) of 10%.



29

Chapter 4 Thermodynamic Properties 4.1 Model for vapor/liquid equilibrium Phase equilibrium data are needed for the design of all separation processes. Experimental data have been published for several thousands binary and many multi-component systems. But no universal equation is available for computing the thermodynamic properties of various kinds of mixture. Instead, there are three types of models to calculate phase equilibrium: the equation-of-state method, the activity coefficient method and a special application method. These are based on constitutive equations because they depend on the constitution or nature of the components in the mixture. First of all, the equation-of-state method is used to describe both liquid and vapor phase behavior. This type is recommended to apply for weak non-ideal solution, such as most hydrocarbon and light gas mixture systems at high and moderate pressure (at least not below atmospheric pressure). This method is applicable for systems where the interaction of the components in the liquid phase is assumed to minimize. Because the design system is a polar, non-ideal solution, comprising of alcohol, aldehyde and water, under the condition of low pressure and moderate temperature, the equation-of-state method is not applicable. Secondly, the activity coefficient method is a combination of two different equations: one to calculate vapor phase properties (j

v, EOS/IG), methods Nothnagel and Hayden-O’Connell with Henry are effective. The other one is to describe liquid phase behavior (j in the unsymmetrical form). This method is widely used in the highly non-ideal mixtures, which are comprised of acetone, alcohol, ester, ether, etc. Especially, high accuracy would be obtained for (chemical) systems with dissimilarity in size, shape and intermolecular forces. Since the chemicals involved in the system are raw material, 3,3-dimethylbutanol with impurity, n-hexane and 1-hexene, and products, 3,3-dimethylbutyraldehyde and water, it is a non-ideal mixture with intermolecular forces, the activity coefficient method should be applicable to model vapor/liquid equilibrium. Thirdly, the special application method is available for systems that for some reasons cannot easily fall into one of the other two methods. In the activity coefficient method, typical equations model liquid phase behaviors are:

Wilson NRTL Uniquac Van Laar UNIFAC

Wilson and Van Laar are only suitable for binary vapor-liquid systems which interaction is small, while NRTL and Uniquac can also be used in cases that there are two liquid phases present. NRTL and Uniquac have better accurate in systems containing water than Wilson, however contrary result is showed for vapor-liquid system. None of these models should be used under critical conditions. A special, widely used activity coefficient model is UNIFAC. This method can predict phase equilibrium for systems for which no experimental data are available. It doesn’t depend on binary interaction parameters, instead bases upon the so-called group contribution method. And a modified NRTL method, NRTL-HOC is used Hayden-O’Connell method to calculate fugacity coefficient. Since the thermodynamic data for



30

raw material-3, 3-dimethylbutanol and 3,3-dimethybutyraldehyde are hardly available in database, even in ASPEN Simulation engine; therefore UNIFAC and NRTL-HOC are chosen to predict vapor/liquid equilibrium for this system. This design aimed to separate the product (3,3-dimethylbutyraldehyde) from the others (mainly water and raw material). Because the heaviest component is the alcohol and the volatile key is water since it has the lowest boiling point among the others. To determine what kind of separation method to be used, the T/x @p constant diagram is plotted for key components by ASPEN PLUS simulation engine.

T-xy for C6H14O/1-C6H12O

Liquid/Vapor Molefrac C6H14O

Tem

pera

ture

F

0 0.2 0.4 0.6 0.8 1

280

300

320

T-x 14.696 psi

T-y 14.696 psi

Figure 4.1: T/x diagram at p = 1 atm for the key component, raw material

T-xy for WATER/C6H12O

Liquid/Vapor Molefrac C6H12O

Tem

pera

ture

F

0 0.2 0.4 0.6 0.8 1

210

220

230

240

250

260

270

280

T-x 14.696 psi

T-y 14.696 psi

Figure 4.2: T/x diagram at p = 1 atm for the key component, product From the figures, it can be seen that separation of alcohol from mixture in the bottom or separation of product from water can be done by distillation. For 0.9 mole ratio of 3,3-



31

dimethylbutanol (recycle to reaction), relative volatility is 1.8. For 0.95 mole ratio of product, relative volatility is 2.5. 4.2 Thermodynamic data 4.2.1 Reaction enthalpy data The reaction enthalpy is calculated from the heat of formation of each component at reaction temperature 100C (Note: Data of heat of formation at 100C are taken from ASPEN simulation engine and Henry’s chemical engineer handbook). Table 4.1: Heat of formation of the components at 100oC (from ASPEN)

Component Hf, 100C (kJ/kmol)3,3-dimethybutanol -328,1503,3-dimethybutyraldehyde -241,942Water -285,703

The reaction is: C6H14O + ½ O2 C6H12O + H2O Hence the heat of the reaction can be calculated as:

o p ro d u c t re a c ta n tre a c t io n a t 1 0 0 CH H H

,100 ,100 ,100 ,100 ,100

,100 ( 285,703) ( 241942) ( 328150) 0 199495( / )

reaction C water C product C rawmaterial C oxygen C

reaction C

H H H H H

H kj kmol

The reaction is exothermic. 4.2.2 Specific heat data Constant pressure specific heat of each component in the Table 2.5 is obtained from ASPEN PLUS simulation engine. As an example of specific heat data of pure components is given below:

Table 4.2: Example of specific heat of the component at 1atm, 25oC Component Specific heat (kJ/kmol-K) at 1 atm3,3-dimethybutanol 156.753,3-dimethybutyraldehyde 148.05Water 73.75TEMPO 213.48

4.3 Validation of method For separations Carlson (1996) suggests to check boiling points of the pure components against values from literature.



32

Table 4.3:Comparison the boiling points of the components obtained by using Joback method and literature values

Component Normal boiling point (C)

Literature Joback & Unifac Error 3,3-dimethybutanol 142.0 152.0 +7.0 %3,3-dimethybutyraldehyde 106.0 109.0 +2.8 %Water 100.0 100.0 0Tempo 247.0 272.0 + 6.1 %RuCl2(PPh3)3 - - -3,3-dimethybutane 47.5 49.7 + 2.2 %3,3-dimethybutene 41.0 41.2 + 0.5 %

As it can be seen from the above table, the boiling point of each component from UNIFAC model is higher than literature and average error is + 3.1 %.



33

Chapter 5 Process Structure & Description 5.1 Criteria and selection This chapter will explain how all information from the foregoing chapters is used to assess the various design criteria. Some of the criteria, which determine the choice of design, are, for example, productivity, product purity, safety, and cost, etc. Moreover, it is explained how these criteria have determined the choice of the following design elements: - Combinations of various tasks into one unit operation - Unit operations, i.e., sequence, type, specific equipment, means of phase contact, type of reactor, type of S/L separator, etc - Final process conditions - Process chemicals, e.g., hydroxides, acids, process water, extractants, etc. - Etc. Please refer to the Process Flow Scheme (part 5.2, chapter 5) in order to follow the explanation below: 5.1.1 Storage tanks Feed alcohol, TEMPO, the product aldehyde and waste TEMPO are stored in storage tanks. The size of the tanks for feed alcohol and product aldehyde is designed to same according to the amount of raw material. It is designed that the liquid only occupies 80% of the volume of the tank for safety reasons. Since it is still considered to be a danger one, the outdoor storage can reduce the explosion potential of the unit and provide good ventilation in case of explosion occurs (refer to Appendix A5.1.1). Sizes of storage tank are designed according to requirement of By-product, because amount of by-product is slight large than that of TEMPO. The risk of explosion is very little, so the tanks can be placed inside the building for the convenience of transportation (refer to Appendix A5.1.2). Hence, safety and productivity criteria determine these choices. 5.1.2 Solid storage Solids (Ru-100 and cold TEMPO) are required to be conveyed to the process at steady rate. Overhead bunkers, also called bins and hoppers, are used for short-term storage of them. They are arranged so that the material can be withdrawn at a steady rate from the base of bunker on to a suitable conveyor. 5.1.3 Dissolving process The catalyst Ru-100 dissolves in TEMPO. It is desirable for the reaction to dissolve it first before it reacts. This dissolving process is done in a separated dissolving tank before entering the reactor. This tank is designed to amount required to 1 day. Hence, productivity criteria determine the choice (refer to Appendix A5.2). 5.1.4 Compressor selection The type of equipment best suited for the pumping of gases in pipelines depends on the flow-rate, the differential pressure required, and the operating pressure. So the choice of compressor is determined by productivity criteria (refer to Appendix A5.10).



34

5.1.5 Separation process Separation unit design is very important in this design, because it determines the quality and the quantity of the product, as well as its economic and environmental effects. A variety of separation processes are used in the design. Water is separated form organic phase in a decanter because of the sufficient difference in density between them (refer to Appendix A5.3 for decanter design and Chap 9 for waste water treatment). Some organic vapor goes out accompany with diluted air from the vent of the reactor can be separated by cooling system and a liquid/vapour separator (refer to Appendix A5.4 and A 5.8 for design details). Flash drum can be used to separate the catalyst and the product stream because of the sufficient difference in boiling between them (refer to Appendix A5.4). Vacuum evaporator is used to separate deactivated Ru-100 from waste (mainly denaturalized TEMPO) (refer to Appendix A5.6). So from the productivity and the economic criteria and also the property of the substances involved, various separation processes are chosen. 5.1.6 Mixing method in the tanks chosen (T05, R01) Internal impellers are chosen as the mixing device in the tanks because the costs of other devices or methods (bubble mixing and external pump) are much more expensive (for details please refer to part 8.5 and 8.6 chapter 8). Therefore, economic criteria determined the choice of mixing method for the tanks. 5.1.7 Cooling method chosen for reactor (R01) Since internal heat transfer surface and external heat exchanger will cost more than jacket heat exchanger, so jacket reactor is used (for details please refer to part 8.6 chapter 8). Hence, economic criteria determined the choice of cooling method for the reactor. 5.1.8 Pumps chosen (P01-P08) Pumps are used in the plant either to transfer liquid through the whole process (P01-07) or to create vacuum (P08). Because the amount of liquid only depends on the requirement of productivity, so the choice of pump for transportation is determined by productivity criteria while the choice of pump for evaporator is determined by vacuum level (which affects the level of separation (refer to Appendix A5.9 for details). 5.1.9 Heat exchangers chosen (H01-H04, H11-H14) The same reason as explained in pump chosen part, productivity criteria also determined the choice of heat exchangers (refer to Appendix A5.8 for more details). 5.1.10 Utility chosen After considering the cost of different utilities and required amount of heat transfer in the plant (calculation is based on plant productivity criteria), the cheapest utilities are chosen from all possible choices (refer to Appendix A5.8 and part 5.4, chapter 5 for detail calculation and selection). Therefore, the utilities selection here is determined by productivity criteria and cost criteria. 5.2 Process flow scheme (see Appendix 3) 5.3 Process stream summary (see Appendix 4.1)



35

5.4 Utilities 5.4.1 Utilities requirement

Table 5.1: The requirements of utilities (1 year = 360 day · 24hrs = 8640hrs)

UTILITIES REQUIREMENTS Requirements

Name

Function

Utility type

Load (kW)

CW (kg/h)

Steam (kg/h)

Electricity (kWh/a)

CW (kg/a)

Steam (kg/a)

T05 Stirrer of dissolving tank Electricity 0.30 2592.0 R01 Stirrer of reactor Electricity 1.44 12441.6 R01 Reactor CW 2447 21142080 C01 compressor Electricity 56.30 486432.0 S05 evaporator steam 12.0 103680H01 Heat exchanger CW 2340 20217600 H02 Heat exchanger Steam 64.0 552960H03 Heat exchanger CW 1152 9953280 H04 Heat exchanger CW 852 7361280 H11 Condenser CW 17496 1.51×108

H12 Reboiler Steam 192.0 0 1658880H13 Condenser CW 216 1866240 H14 Reboiler Steam 1.6 13824P01 Pump Electricity 0.32 2764.8 P02 Pump Electricity 0.80 6912.0 P03 Pump Electricity 0.45 3888.0 P04 Pump Electricity 0.76 6566.4 P05 Pump (into dissolving tank) Electricity 0.10 864.0 P06 Pump(out of dissolving tank) Electricity 0.10 864.0 P07 Pump (in the evaporator) Electricity 1.10 9504.0 P08 Vaccum Pump Electricity 1.00 8640.0 Total 62.67 24503 269.6 541468.8 2.12×108 2329344

Remarks: - Please also refer to appendix A4.2 for the summary of utilities - Power needed for a stirrer of dissolving tank is calculated in appendix A5.2 - Power needed for a stirrer of reactor is calculated in appendix A5.3 - Steam and cooling water for heating and cooling the reactor are calculated in appendix A5.8 - Column feed heater, reboiler & condenser are calculated in appendix A5.8 - Power for the pumps are calculated in appendix A5.9 of pump design

5.4.2 Options and selection There are several options for choosing utilities. Cold utilities, which are CW and CA (cooling air), can be chosen. Hot utilities, which are LP, MP, HP and hot oil, are available. Cooling water and low pressure steam are chosen as cold utility, hot utility because:

- Cold utility: CW is chosen as cold utility, because the mixture in the reactor needs to be maintained at 1000C during the reaction. It is also because the temperature of liquid should not be too high when it goes through the pipes. It means a large amount of heat should be removed from the reactor (refer to Appendix A5.8 for calculations). For cooling water the heat capacity (4.18 kJ/kg



36

K) and heat film coefficient (0.20) of water are higher than that of air (1.00 kJ/kg K and 0.10 respectively). It means that less amount of cooling water needed to achieve the same effect as air and the area needed is much less.

- Hot utility: MP steam is chosen as hot utility. All hot utilities available including hot oil (280-2600C), HP steam (25bar 2250C), HP steam (18bar, 2050C), MP steam (4bar 1400C) and LP steam (1.5 bar 1100C) can be used as hot utility for both the flash drum (S02) and distillation columns (S03-S04). The reason to choose MP steam is because of economic reason and heating efficiency. The price of MP steam is cheaper than HP steam and the cost of equipment for MP steam needed is also cheaper than that for HP steam. On the other hand, the heating efficiency of MP steam is higher than that of LP steam. Based on the comparison of the prices and efficiencies between LP steam MP and HP steam, it is decided to use MP steam as hot utility in the process (please refer to Appendix A5.8 for heat calculations).

5.4.3 Major users and possibilities for future reduction

Table 5.2: The cost of the utilities

Utility Unit Cost per unit Total cost (Dfl) (Euro/a) Electricity kWh 0.22 54,233Steam ton 30.00 31,763Cooling water ton 0.10 9,623

Total 95,619 From the table 5.11 it is found that electricity power and steam counts for the major cost of utilities. The reason for high cost for electricity is its high amount demanded, while the high cost for steam is due to its high price. Electricity is needed for the compressor, the stirrers and all the pumps power (the power needed for the pumps are relatively low). Among all, the compressor is the main electricity-consumer. Some alternative process is considered without using the compressor or to reduce the power needed by the compressor. The first idea is to remove the compressor. In order to realize high pressure (10bar) in the reactor, liquid nitrogen can be injected into the reactor. It seems that cost for liquid nitrogen is also not low. The second idea is to reduce the power needed by the compressor by using a relatively low pressure in the reactor, i.e., 8bar. As a result, the transportation of oxygen from vapor phase to liquid phase has been reduced. This problem can be solved either by inject more air into the reactor or by enhancing mixing in the reactor. Both of them seem to be of low cost. To reduce the steam used in the process, it is recommended to introduce heat integration through the whole process by combining the cooling system with the heating system. On the other hand, equipment used for heat exchange will add to the investment as well.



37

5.5 Process yields Table 5.3a: The process yield (excluding the utilities)

Table 5.3b: The process yield (the utilities)

Utilities Name

Ref. Stream

t/a

kWh/a

t/t product

kWh/

t product Steam - 2,329 2.173Cooling water - 211,706 197.586Electricity - 542,333 506.163

Comments: - Yield of the product is relatively high because of high selectivity of the reaction

under the determined reaction condition (100C, 10bar). - The amount of steam and cooling water required is relative high because of

multiple-stage of separation process. The two distillation column and the two flash drums need much energy.

- Electricity is need for compressor, mixing-stirrers and pumps.

Stream Name

Ref. Stream

t/a t/t product IN OUT IN OUT

HEXANE 3.37 3.37 0.003 0.003HEXENE 2.25 1.91 0.002 0.002HEXNAL 0.00 0.40 0.000 0.000WATER 0.00 192.72 0.000 0.180TEMPO 98.50 98.50 0.092 0.092E 1117.58 24.51 1.043 0.023A=Product 0.00 1071.46 0.000 1.000O2 235.87 64.65 0.220 0.060N2 887.33 887.33 0.828 0.828Ru-100 208.57 208.57 0.195 0.195Total 2553.46 2553.42 2.383 2.383



38

Chapter 6 Process Control The process can be divided to two parts: reaction section and separation section. 6.1 Reaction section 6.1.1 Control of dissolving tank (T05) The liquid level in the dissolving tank is controlled by two inlet valves. And the outlet goes accompany with recycle stream <20> into the reactor. The control of this stream is depicted in part 6.1.2. (Figures between <> is the stream number in PFS. Please also refer to 5.2 Chapter 5)

Figure 6.1 the control configuration for the dissolving tank

6.1.2 Control of reactor (R01) The control of the reactor includes four stages:

1. Inlet flow control: There are three inlet flows to the reactor. The flow rate of all these streams is controlled.

2. Injected air temperature control: A large amount of air is injected to the reactor, thus its temperature affects the reactor temperature a lot. It is decided to control the temperature of this stream by adjust the flow rate of cooling water.

3. Reactor temperature control: The temperature in the reactor is control by adjust the flow rate of cooling water in the cooling jacket.

4. Liquid level in the reactor control: The liquid level is controlled by the flow rate of outlet stream <10>.

LC



39

Figure 6.2 Control configurations for reactor

6.2 Separation section 6.2.1 Flash drums S01-S02: 1. Pressure control: The pressure in the flash drum is controlled by the valve on the

outlet stream for vapor phase <12>. 2. Temperature control: The temperature in the flash drum is controlled by the flow rate

of the heat exchange mediate (CW for S02, or steam for S01). 3. Liquid level control: The liquid level in the flash drum is controlled by the flow rate

of the liquid outlet.

Figure 6.3 Control configurations for flash drum

FC

TC

CW

LC

TC

CW

FC

FC

LC

TC

PC



40

6.2.2 Distillation columns S03-S04: The top stream will go through the condenser and the reflux drum. In the condenser, it is important to control the flow rate of cooling water to make sure that heat can be removed effectively and vapor can condense to liquid. If the follow rate of cooling water varies, the liquid level in the reflux accumulator will also vary. So the flow rate of cooling water should be controlled. It is obvious that a simple flow controller can fulfill this task. Alternatively, this can also be done by control the pressure on the top of the column by adjusting the flow rate of the cooling water. The second choice seems to be simpler. The liquid level in the reflux accumulator should be taken care. If the level is too high or too low, the flow rate of the leaving stream and reflux stream will vary, and the balance of column system will be disturbed. The liquid level is controlled by the flow rate of leaving stream. If the level is too high, the leaving stream’s flow rate will increase; if the level is too low, the leaving stream’s flow rate will decrease. The liquid level can’t be controlled by another stream, reflux stream. Otherwise, in the column the vapor-liquid equilibrium will be disturbed frequently. Secondly, the flow rate of the leaving stream (especially for the product) should be kept stable. So the cascade controller is used here. If the steam flow rate changes a bit, the flow controller will first adjust it back to the required value. Then, the stable product flow goes out. If the liquid level varies, the level controller will control the flow rate more easily and effectively. The reflux stream is critical for the operation of column. If the reflux flow changes, the whole column system will be disturbed. The flow rate of down-flow liquid will change and disturb the equilibrium of liquid and vapor on every tray. The whole operation and product quality will be affected. So the reflux stream flow rate is important. It is obvious that a flow controller can fulfill this task as shown in the figure below.

Figure 6.4 Control configurations for the tope stream form the distillation column

Inside the bottom of the column, there is also liquid, which should be kept at certain level. If there is no liquid, the column cannot be operated; if the liquid is too much, more heat is needed and efficiency decreases. To control the liquid level in the bottom of column, only the leaving stream can be controlled to keep the level. A liquid controller is used here. The stream, which enters the reboiler, cannot be controlled because it will affect the amount of vapor and further the system balance. The temperature in the bottom of the column is important for the quality of the product. Because the source of heat is from the steam entering the reboiler, it is obvious that the

PC

FC

FC

LC

H11



41

temperature can be controlled by the steam flow rate. A controller is installed to control the temperature by adjusting the steam.

Figure 6.5 Control configurations for the bottom stream from distillation column

6.2.3 Evaporator S05: The control of evaporator is the same as flash drum. (Refer to 6.2.1 Chapter 6 and 5.2 Chapter 5).

TC

H12

LC



42

Chapter 7 Mass and Heat Balances 7.1 Practical aspects

In the following section, the mass and heat balances for total stream and for stream components are made. The mass and heat balance in reaction part is given in table 7.1, while in table 7.2, the mass and heat balance in separation part is shown. As it can be seen in the tables, no imbalances do occur. All mass and heat coming to the plant are the same as those going out from the plant. In table 7.3, the overall component mass and stream heat balance is shown. It provides the difference between enthalpies of IN and OUT going streams. This difference should equal to the difference between heats “IN” and “OUT” as introduced or removed by heat exchanger equipment (steam, cooling water, air cooling, etc). As it can be seen from table 7.3, imbalances do occur in the mass and heat balance. These imbalances are due to round-up factor. Since the calculation is done with both Microsoft Excel and Aspen Simulation, and the results are rounded-up to 3 decimal places, this round up might be carried on until the last calculation and causes imbalances at the end. However, the errors are just very small (0.00135% for the mass balance and 0.02 % for the heat balance) and they will not give any influence on the final result. In table 7.4, the overall component mass and stream heat balance is shown. It provides the difference between enthalpies of IN and OUT going streams. This difference should equal to the difference between heats “IN” and “OUT” as introduced or removed by heat exchanger equipment (steam, cooling water, air cooling, etc).



43

7.2 Balance for total stream

Table 7.1: Heat and mass balance for streams total of reaction part

HEAT & MASS BALANCE FOR STREAMS TOTAL

REACTION PART

IN OUT

Plant EQUIPMENT EQUIPM. EQUIPMENT Plant

Mass Heat Mass Enthalpy Stream IDENTIF. Stream Mass Enthalpy Mass Heat

kg/hr kW kg/hr kW Nr. Nr. kg/hr kW kg/hr kW

130.000 -135.964 130.000 -135.964 <1> P01 <2> 130.000 -135.855

-0.320 -0.320

130.000 -135.855 <2> M01 <6> 554.882 -500.609

424.882 -364.754 <31>

130.000 -136.284 684.882 -636.893 Total 684.882 -636.464

130.000 -0.005 130.000 -0.005 <3> M02 <4> 722.984 -18.229

592.984 -18.224 <14>

130.000 -0.005 722.980 -18.229 Total 722.980 -18.229

722.984 -18.229 <4> C01 <7> 722.984 71.951

-56.300 -56.300

722.984 71.951 <7> <32> 722.984 -20.555

H04 -92.506 -92.510

-56.300 1445.968 -2.578 Total 1445.968 -41.110 0 -92.510

35.54.0 -1.228 <5> M04 <8> 443.913 -92.291

408.373 -91.062

443.913 -92.290 Total 443.913 -92.291

443.913 -92.291 <8> P02 <9> 443.913 -92.103

-0.800 -0.800

-0.800 443.913 -94.691 Total 443.913 -92.103

554.882 -500.609 <6> R01 <10> 1721.773 -698.584

722.984 -20.555 <32>

443.913 -92.103 <9>

-1.440 -1.440 -85.320 -85.320

-1.440 1721.8.00 -617.590 Total 1721.800 -783.900 -85.320

11.400 -0.520 11.400 -0.520 <35> T05 <5> 35.540 -1.228

24.140 -0.708 24.140 -0.708 <36>

0.100 0.100 P05

0.100 0.100 P06

35.540 -1.028 35.540 -1.028 Total 35.540 -1.228

295.540 -195.217 5498.996 -1462.660 Total 5498.996 -1665.330 0 -177.830

OUT-IN: -295.540 17.387



44

Table 7.2: Heat and mass balance for streams total of separation part HEAT & MASS BALANCE FOR STREAMS TOTAL

SEPARATION PART

IN

EQUIPM.

OUT

Plant EQUIPMENT EQUIPMENT Plant

Mass Heat Mass Heat Stream IDENTIF. Stream Mass Heat Mass Heat


1721.773 -698.584 <10> <11> 1721.773 -750.376

H01 -51.792 -51.792

1721.773 -750.376 <11> S01 <12> 705.933 -21.695

<15> 1015.840 -728.681

3443.550 -1448.960 Total 3443.550 -1552.540 0.000 -51.790

705.933 -21.695 <12> M03 <13> 112.949 -3.471 112.950 -3.471

<14> 592.984 -18.224

705.930 -21.695 Total 705.930 -21.695 112.950 -3.471

1015.840 -728.681 <15> D01 <16> 21.955 -96.718 21.955 -96.718

<17> 993.885 -631.963

1015.800 -728.680 Total 1015.800 -728.680 21.955 -96.718

993.885 -631.963 <17> H02 <18> 993.885 -592.821

-39.142 -39.142

993.885 -592.821 <18> S02 <19> 453.747 -101.180

<25> 540.138 -419.280

-39.140 1987.770 -1263.930 Total 1987.770 -1113.280

453.747 -101.180 <19> M06 <20> 408.373 -91.062

<21> 45.375 -10.118

453.75 -101.180 Total 453.750 -101.180

45.375 -10.118 <21> S04 <22> 37.050 -2.692

<23> 8.325 -7.966

-0.540 -0.540 -1.040 -1.040

-0.540 45.375 -11.198 Total 45.375 -9.826 -1.040

8.325 -7.966 <23> M07 <24> 8.325 -7.966

8.325 -7.966 <24> M05 <26> 548.463 -427.246

540.138 -419.280 <25>

548.460 -427.250 Total 548.460 -427.250

548.463 -427.246 <26> H03 <27> 548.463 -520.248

-93.002 -93.002

-0.294 548.463 -520.248 <27> P04 <28> 548.463 -519.954

-0.760 -0.760

-1.050 1096.930 -948.250 Total 1096.930 -1133.200 0.000 -93.000



45

Table 7.2: Heat and mass balance for streams total of separation part (cont’d) HEAT & MASS BALANCE FOR STREAMS TOTAL

SEPARATION PART

IN

EQUIPM.

OUT

Plant EQUIPMENT EQUIPMENT Plant

Mass Heat Mass Heat Stream IDENTIF. Stream Mass Heat Mass Heat


548.463 -519.954 <28> S03 <29> 123.581 -109.261 123.581 -109.261

<30> 424.882 -364.957

-94.000 -94.000 -81.300 -81.300

-94.000 548.460 -425.950 Total 548.463 -555.518 123.581 -190.561

424.882 -364.957 <30> P03 <31> 424.882 -364.754

-0.450 -0.450

-0.450 424.882 -364.957 Total 424.882 -364.754

37.050 -2.692 <22> S05 <33> 21.770 -1.012 10.260 -2.423

-7.300 -7.300 <37> 25.280 -1.686 26.790 -0.269

-1.100 -1.100 P07

-1.000 -1.000 P08

-9.400 37.050 7.708 Total 37.050 -2.692 37.050 -2.692

-244.586 10307.960 -5942.150 TOTAL 10307.950 -6010.620 295.536 -259.834

OUT-IN: 295.536 -15.248

Table 7.3: Mass and heat balances

Mass Heat IN OUT IN OUT

Kg/h Kg/h KW KWReaction part 295.540 0.000 -195.217 -177.830Separation part 0.000 295.536 -244.586 -261.934Total OUT-IN 0.00400 0.039ERROR (%) 0.00135 0.020 From table 7.3, it can be seen that the net heat is almost the same as the net enthalpy; the small error is just due to round-up factor. 7.3 Balance for stream components Mass balance per component inside battery limit is shown in the following table:



46

Table 7.4: Overall component mass and stream heat balance Overall Component Mass and Stream Heat Balance

Stream Nr. <1>+<3>+<35> IN <13>+<16>+ OUT OUT-IN +<36> <29>+<33>+<37> Name Total Plant Total Plant Total Plant Component ton/year kmol/year ton/year kmol/year ton/year kmol/year HEXANE 3.37 39.10 3.37 39.10 0.00 0.00 HEXENE 2.25 26.69 1.91 22.66 -0.34 -4.03 HEXNAL 0.00 0.00 0.40 4.01 0.40 4.01 WATER 0.00 0.00 192.72 10697.45 192.72 10697.45 TEMPO 98.50 630.38 98.50 630.37 0.00 -0.01 E 1117.58 10937.10 24.51 239.91 -1093.07 -10697.19 A 0.00 0.00 1071.46 10697.34 1071.46 10697.34 O2 235.87 7371.03 64.65 2020.52 -171.22 -5350.51 N2 887.33 31674.76 887.33 31674.81 0.00 0.05

Ru-100 208.57 217.47 208.57 217.47 0.00 0.00

Total 2553.46 50896.53 2553.42 56243.64 -0.05 5347.11 Enthalpy (MJ) -1185.39 -1832.91 -647.52



47

Chapter 8 Equipment Design 8.1 Storage tank for feedstock (T01) T01 is designed as storage tank, which can contain the amount of 3,3-dimethylbutanol needed for 10 days production in the plant. The storage amount of the main material is chosen on the basis of balancing conveniently arrangement of production and transportation against economical and safety factors. Approach: - Approximated amount of liquid is used to calculate design volume - Chosen diameter (for good tank size) is used calculate the height Table 8.1: Design parameters of storage tank T01 Volume (m3) 47.7 Remark: refer to appendix A5.1.1 for

detailed calculation. Diameter (m) 3.8Height (m) 4.2

8.2 Storage tank for TEMPO (T02) T02 is designed as storage tank, which can contain the amount of 3,3-dimethylbutanol needed for 30 days’ production in the plant. The storage amount of the main material is chosen on the basis of balancing conveniently arrangement of production and transportation against economical and safety factors. Approach: - Approximated amount of liquid is used to calculate design volume

- Chosen diameter (for good tank size) is used calculate the height Table 8.2: Design parameters of storage tank T01 Volume (m3) 12.9 Remark: refer to appendix A5.1.2 for

detailed calculation. Diameter (m) 2.4Height (m) 2.9

8.3 Storage tank for product (T03) T03 is designed as product storage tank before product is filled into containers. Approach: - the same approach as for raw material storage tank T01

Table 8.3: Design parameters of product storage tank T03

Volume (m3) 47.7 Remark: refer to appendix A5.1.1 for detailed calculation. Diameter (m) 3.8

Height (m) 4.2 8.4 Storage tank for deactivated TEMPO (T04) T04 is designed as storage tank, which can contain deactivated TEMPO, in order to further deal with it as a waste or to be sold as a by-product Approach: - the same approach as for TEMPO storage tank T02Table 8.4: Design parameters of deactivated TEMPO storage tank T04 Volume (m3) 12.9 Remark: refer to appendix A5.2 detailed

calculation. Diameter (m) 2.4Height (m) 2.9



48

8.5 Dissolving tank (T05) T05 is designed as a tank for dissolving catalyst RuCl2(PPh3)3 in TEMPO. Mixing method chosen: There are three normal ways to stir the contents in the tank:

- Using internal impellers; - Gas bubbles; - Pump around.

The ‘pump around’ method is always used in big scale conditions and the external pump is required. Because the concentration of catalyst RuCl2(PPh3)3 in solution is high (33% mole), it is not suitable to mix using gas bubbles. Internal impeller is a versatile mixing method. Therefore, internal impeller is chosen. Inlet and outlet holes for T05: (Please also refer to the discription of the block scheme for the need to design the inlet and outlet holes for T02) There are two inlet holes on the dissolving tank (refer to figure 8.5 below)

- No1 is the hole, which connects T02 with TEMPO input stream;

- No2 is designed to fit the hopper for adding solid (catalyst RuCl2(PPh3)3) into the tank. It is added manually into the reactor.

There are two outlet holes on the dissolving tank (refer to figure 8.5)

- No 3 is designed as a hole connecting to atmosphere;

- No 4 is designed as the hole for liquid outlet to mixer M04.

Approach: - Liquid volume is calculated

- Liquid volume is used to calculate design volume - Chosen diameter (for good tank size) is used to calculate the height - Universal mixing time relation is used to calculate mixing time and

power needed: 2 21/3

2/3 4 /3 1/3m s

mix

t L HN

D H D

[29]

Table 8.5: Design parameters of dissolving tank T05 Volume (m3) 0.7

Remark: refer to appendix A5.2 (dissolving tank design) for detailed calculation.

Diameter (m) 0.8Height (m) 1.4Mixing time (s) 86.0Power (kW) 0.4

8.6 Reactor (R01) Reactor is the most important equipment in the plant. There are several steps need to be followed during the reactor design. in the process of reactor design, mass transfer, which

Figure 8.5 Dissolving tank T05



49

is very important to a gas-liquid multiphase reactor, is on the base of experimental data of Arie DIJKSMAN’ thesis (2001). Type of reactor chosen: Agitated gas-liquid reactor is chosen. For agitated gas-liquid reactor the following assumptions are usually made:

The gas flow is perfectly mixed; The pressure is constant for low liquid height; Mass transfer coefficients, the gas hold-up and the specific interfacial are constant

over the whole reaction volume; The liquid phase is completely mixed.

Mixing method chosen: Internal impeller is chosen. (Please refer to mixing method chosen in T05 selection for reasons). Heat exchange equipment chosen: Because alcohol oxidation is an exothermic reaction, the heat balance of an agitated gas-liquid reactor depends on both the reactor vessel is cooled by an outside cooling medium and the liquid vaporizes. The system in the reactor is multiphase system. If internal heat transfer surface is used there might be a possibility in which the mixing condition is affected. Also internal heat transfer surface will increase the size of reactor, and leads to increase in the reactor’s cost. Due to these reasons, internal heat transfer surface will not be chosen [32]. Inlets and outlet holes of reactor: (Please also refer to chapter 3 of the discription of the block scheme for the need to design the inlet and outlet holes for the reactor) There are three inlet holes on the reactor (please also refer to figure 8.6 below) No 1 is the hole, which connects with stream 6 for mixture of raw material and recycle alcohol stream. No 2 is also a hole, which connects with stream 9, mixture of solution from dissolving tank with recycle catalyst stream. No 3 is designed as gas (oxygen, nitrogen) inlet. There is only one outlet hole on the reactor (refer to figure 8.6 below).

- No 4 is designed as the hole for product with gas outlet (for nitrogen and air)

Figure 8.6: Reactor design

Approach: - Liquid volume is calculated - Liquid volume is used to calculate design volume - Chosen diameter (for cheaper reactor cost) is used to calculate the height



50

- Universal mixing time relation is used to calculate mixing time and

power needed: 2 21/3

2/3 4 /3 1/3m s

mix

t L HN

D H D

[29]

Table 8.6: Design parameters of reactor

Vessel

Volume (m3) 1.44 Remark: refer to appendix A5.3 (reactor design) for detailed calculation.

Diameter (m) 1.00 Height (m) 1.84 Heat transfer surface (m2) 4.21

Stirrer

Mixing time (s) 104.00 Power (kW) 1.44 Impeller diameter (m) 0.33 Impeller height (m) 0.04

8.7 Gas-Liquid Separator (S01) A flash drum with demister pad is used to separate gas from liquid phase (product mixture). The layout of a vertical liquid-gas separator is shown in figure 8.7 Approach: The diameter of the vessel must be large enough to slow the gas down to below the velocity at which the liquid will settle out. Table 8.7: Design parameters of Gas-liquid separator S01 Volume (m3) 0.42

Remark: refer to appendix A5.4 for detailed calculation.

Diameter (m) 0.60Height (m) 1.50Residence time (min) 10.00

8.8 Decanter (D01) The decanter is the simplest form of equipment used to separate two liquid phases, immiscible or partially miscible liquids, where there is a sufficient difference in density between the liquids for the droplets to settle readily. Decanters are essentially tanks which give sufficient residence time for the droplets of the dispersed phase to rise (or settle) to the interface between the phases and coalesce. A great variety of vessel shapes is used for decanters, but for most applications a cylindrical vessel will be suitable, and will be cheapest shape (refer to figure 8.6). Figure 8.8 Decanter

L1 + L2 L1

L2

Figure 8.7



51

Table 8.8: Design parameters of decanter D01 Volume (m3) 0.66


Diameter (m) 0.75Height (m) 1.50Residence time of droplets (min) 4.60

8.9 Flash drum (S02) A flash drum is used to separate vapor from liquid phase (product mixture). The layout of a vertical liquid-gas separator is shown in figure 8.9 (Please refer to section 8.7). Table 8.9: Design parameters of Flash drum S02 Volume (m3) 1.73


Diameter (m) 1.00Height (m) 2.20Residence time (min) 10.00

8.10 Evaporator system (S05) Evaporation is the removal of a solvent by vaporization, from solids that are not volatile. It is normally used to produce a concentrated liquid, and a solid product can be obtained. The primary components of an evaporator are usually a heat exchanger, and a separation vessel. There are a number of different types of evaporators, selected depending on the type of product to be handled. Most evaporators will fall into two categories: (figure 8.10) film evaporators, where boiling is allowed to take place within the heat exchanger, and suppressed-boiling evaporators, where the liquid is heated in the heat exchange. Figure 8.10 thin film evaporator (a) and suppressed boiling evaporator The selection of the most suitable evaporator type for a particular application will depend on the following factors: throughput required, viscosity, nature of the product required (solid, slurry, or concentrated), heat sensitivity of the product, etc. In general, film evaporators are less expensive, from both a capital cost and operating cost standpoint

(a) (b)

Figure 8.9



52

view. Film evaporators are usually the preferred equipment for handling material that has little suspended solid. The thin film evaporator has high capacity, suitable for low viscosity. Forced-circulation evaporator is suitable for use with materials which tend to foul the heat transfer surfaces, (viscosities of up to 100 cP to 500 cP)and where crystallization can occur in the evaporator. Wiped-film evaporators are used for very viscous materials and for producing solid products. The main advantage of the suppressed evaporator is its ability to handle moderately fouling products. It is a better choice when handling products with high levels of suspended solids. In this case, suppressed-boiling evaporator is chosen for solid product. Table 8.10: Design parameters of evaporator system Vessel

Volume (m3) 0.4 Remark: refer to appendix A5.6 for detailed calculation.

Diameter (m) 0.6 Height (m) 1.4 Residence time (min) 10.0

Heat exchange

Area of heat-exchange (m2) 0.3 Duty (kJ/s) 7.3 Temperature inlet (0C) 162.0 Temperature outlet (0C) 172.0

Vacuum pump P08

Mass flow rate (kg/h) 11.8 Head (m) -8.8 Power (kw) 1.0

Recycle pump P07

Mass flow rate (kg/h) 1177.0 Head (m) 26.0 Power (kw) 2.5

8.11 Compressor (C01) The type of equipment best suited for the pumping of gases in pipelines depends on the flow-rate, the differential pressure required, and the operating pressure. In general, fans are used for the pressure drop is small (<0.03bar); axial flow compressors for high flow rates and moderate differential pressures; centrifugal compressors for high flow rates and high differential pressures. Reciprocating compressors can be used over a wide range of pressures and capacities, but normally only specified in preference to centrifugal compressors where high pressures are required at relatively low flow-rates. In this case, flow rate 723 kg/h (560.46 m3/h) is smaller than 1000 m3/h; differential pressure is 9.5 bars. So, a reciprocating compressor is chosen. Table 8.11: Design parameters of compressor C01 Volumetric flow rate (m3/h) 560.46


Discharge pressure(bar) 10.50Power (kw) 56.30stage 2.00Interstage pressure (bar) 3.30



53

8.12 Heat exchangers Heat exchangers are used as a cooler (H01) for feed to gas-liquid separator, a heater (H02) for feed to flash dram, a heater (H03) in distillation column S03, a cooler (H04) for feed air, a reboiler (H11), a condenser (H12) and in distillation column S04 as a rebolier (H13), a condenser (H14). Type of heat exchangers chosen: The most commonly used type of heat-transfer equipment, which is shell and tube exchanger is chosen due to its many advantages such as:

- The configuration gives a large surface area in a small volume - Good mechanical layout: a good shape for pressure operation - Well-established fabrication techniques - Can be constructed from wide range of materials - Easily cleaned - Well-established design procedures

Type of tube and shell heat exchangers chosen: There are different kinds of tube and shell heat exchanger are used in the industry, including baffle spacers and tie rods, fixed –tube plate, internal floating head without clamp ring, U-tube, single tube etc. Because the temperature difference for flows in the plant is not very much and also the amount of flows is not much, which is only between 60-1000 kg/h (please refer to mass balance and heat balance in chapter 7), so total amount of heat needed to be transferred is not so much and heat transfer surface required is quite small. Therefore, single tube heat exchanger is chosen.

Figure 8.12 Single tube and shell heat exchanger Approach: - Using known heat transfers power and mean temperature difference

(calculated from temperature difference using equation:

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

[8]) heat transfer surface is calculated

Table 8.12: Design parameters of heat exchangers H01-03, H11-H14

H01 H02 H03 H04 H11 H12 H13 H14

Hot utility

MP steam

MP

steam

MP steam

Cold utility Cooling water*

Cooling water*

Cooling water*

Cooling water

Cooling water*

Temperature range oC

22-100 25-220

22-157 22-227 20-25 215-220

22-74 119-220

exchange area (m2) 10.9 1.4 5.2 4 81.3 94 0.1 0.2

CW or Steam

Liquid needs to be heated or cooled



54

Remark: - Refer to appendix A5.8 for heat exchangers design - Medium pressure steam is superheated steam, 220oC, 10bar * Recycle cooling water from outlet of condenser in column S03 as cool media

8.13 Distillation column (S03) There are 2 types of distillation column: plate and packed column. The choice can usually be made based on experience. Normally, plate column can be designed with more assurance than packed columns. There is always some doubt that good liquid distribution can be maintained throughout a packed column under the operating conditions. In our case, not only is the purity requirement for the product quite high, more than 99%, but also because of nature character of the product, which is food additive, hence the quantity of the product should be kept high and stable. So, hereby the plate column is chosen.

Cross-flow plates are the most common type of plate contactor used in distillation columns. In a cross-flow plate the liquid flows across the plate and the vapor up through the plate. The flowing liquid is transferred from plate and a pool of liquid is retained on the plate by an outlet weir. Three principal types of cross-flow tray are used which are classified according to the method used to contact to the vapor and liquid:

- Sieve plates - Bubble-cap plates - 3 valve plates

The principal factors to be considered when comparing the performance of bubble-cap, sieve and valve plates are: cost, capacity, operating range, efficiency and pressure drop.

Cost: Bubble-cap plates are appreciably more expensive than sieve or valve plates. For mild steel the ratios of cost of bubble-cap to valve to sieve are approximately 3.0:1.5:1.0

Capacity: There is little difference in the capacity rating of the three types; however the ranking is sieve, valve and bubble-cap

Operating range: Bubble-cap plates have a positive liquid seal and can operate efficiently at very low vapor rates. Sieve plates rely on the flow of vapor through the holes to hold the liquid on the plate, and cannot operate at very low vapor rates; Valve plates are intended to give greater flexibility than sieve plates at a lower cost than bubble-caps.

Efficiency: The efficiency of the three types of plate will virtually the same when operating over their design flow range, and no real distinction can be made between them.

Pressure drop: The pressure drop over the plates can be an important design consideration, particularly for vacuum columns. In general sieve plates give the lowest pressure drop, followed by valves, then bubble-caps.

Therefore, from the point of view of cost and capacity, sieve plate is the first choice. In this case, one distillation column is operated under vacuum condition (0.1bar) and another is done under high pressure (6.5bar), sieve plate can deal with them. According to the calculation, (see appendix A5.7) the pressure drop along the whole column is less



55

than 0.01bar that is corresponding to the performance of sieve plate. In all, sieve plate is the choice in this case. McCabe-Thille method is the basic way to design distillation column. The dimensioning design procedure is given by J. M. Coulson & J. F. Richarson (1979) [8]. In Delft University of Technology, another effective method is used, so-called Delft Method [35]. In our design, integration by process simulation is done first for the calculation, which is shown in the appendix A5.7, and then Delft Method is used for dimensioning the column.

Delft method is an iterative procedure: - Internal flows and physical properties along the column - Choice of column type, i.e. vapor/liquid contacting device - Determination of column diameter and column height - Check pressure drop and adjust all calculation accordingly.

Table 8.13: The key dimensioning result of distillation column

Diameter (m) 0.650Total height (m) 32.100Tray spacing (m) 0.300Weir length (m) 0.480Weir height (m) 0.030Hole diameter (m) 0.005Remark: - Refer to appendix A5.7 for detailed calculation

8.14 Distillation column (S04) (Please refer to section 8.13)

Table 8.14: the key dimensioning result of distillation column S04

Diameter (m) 0.400Total height (m) 8.800Tray spacing (m) 0.150Weir length (m) 0.300Weir height (m) 0.025Hole diameter (m) 0.004Remark: - Refer to appendix A5.7 for detailed calculation

8.15 Mixing equipment (M01- M07) Specialized equipment is seldom needed for mixing gases, which because of their low viscosities mix easily. The mixing given by turbulent flow in a length of pipe is usually sufficient for most purposes. Also, for the continuous mixing of low viscosity fluids inline mixers can be used. 8.16 Pumps and pipes

Pumps are required on pipes that carry liquids because of pressure drop. Pumps can be classified into two general pumps:



56

- Dynamic pumps, such as centrifugal pumps - Positive displacement pumps, such as reciprocating, diaphragm and rotary pumps

The single-stage, horizontal, overhung, centrifugal pump is the most commonly used in the chemical process industry. Other types are used where special process considerations are specified, such as vacuum.

Pumps selection is made based on the flow rate and head required, together with other process considerations, such as corrosion, etc.

In order to determine the power of the pump, first the pipe size has to be determined. The most economic pipe diameter would be the one which gives the lowest operating costs. The capital cost of pipe run increases with diameter while the pumping cost decrease with diameter increasing. The formula is given below [8]:

d, optimum = 260 G0.52 ρ-0.37

Where, G = mass flow rate in kg/s

ρ = density of fluid in kg/m3

To calculate the pressure drop the pipe friction factor needs to be known. This is a function of Reynolds Number.

Pump P01: Raw material transport pump This pump is designed to pump 3,3-dimethylbutanol from storage tank T01 to reactor R01. The choice of pump will depend on the capacity and head requirement. Centrifugal pumps will normally be the first choice for pumping process fluids, the other types only being used for special applications, such as the use of reciprocating and gear pumps for metering. The capacity range of centrifugal pumps is 0.25-1000m3/h and maximum head is 150m of water for single stage, 1650m of water for multistage. Efficiency 45% at 22.7 m3/h, 70% at 113.5 m3/h, 80% at 2270 m3/h. In this case, the flow rate of raw material is 158 l/h, head required is 95m. So, centrifugal pump can be used.

Pump P02: Recycle catalyst transport pump This pump is designed to pump liquid from mixer M04 to reactor R01. The flow rate of liquid is 7.2m3/h and maximum head required is 36m so centrifugal pump can be used as liquid transport pump. Pump P03: Liquid transport pump This pump is designed to recycle the bottom product of distillation column S03 to reactor R01, bypassing mixer M01.

Pump P04: Column’s feed pump

This pump is designed to pump the liquid from flash drum S02 to distillation column S03.

Pump P05: TEMPO pump

This pump is designed to pump the liquid TEMPO from storage tank T02 to dissolving tank T05.

Pump P06: mixture of catalyst feed pump This pump is designed to pump the liquid from dissolving tank 05 to mixer M04.



57

Pump P07: evaporator recycle pump This pump is designed to recycle the liquid in evaporator system. Pump P08: Vacuum pump in evaporator system

To create vacuum, rotary pump or steam jet ejector can be used to create 0.13bar [34]. In rotary pumps the liquid is displaced by rotation of one or more members within a stationary housing. The selection of materials for designing rotary pumps is critical. To create 0.13bar, Rotary pump is used, according to Perry’s chemical engineering handbook ([34] refer to figure 10-105: Vacuum levels attainable with various types of equipment. P10-59). Approach: Firstly, the suction pressure and discharge pressure are calculated separately. The pressures differential across the pump and power of centrifugal pump are calculated.

Table 8.16: Design parameters of transport pumps P01-P07

Pumps

Head (kPa)

Flow rate

pQ (m3/s) Pump efficiency

p (%) Power (kW)

P01 1190.0 0.195 22 0.32P02 1462.0 0.436 22 0.80P03 908.0 0.520 30 0.45P04 904.0 0.670 22 0.76P05 217.3 0.014 20 0.10P06 217.3 0.014 20 0.10P07 253.3 0.763 50 1.10P08 -90.0 0.003 90 1.00Remarks: - Refer to 5.2, chapter 5 of PFS for the name and position of the pumps - Refer to appendix A5.9 for detailed calculation of pumps - Pump efficiency is determined according to J. M. Coulson & J. F. Richarson, 1979, Chemical Engineering, volume 6 (p435, figure 10.62)

8. 17 Equipment data sheets

Equipment Data Summary Sheets and Equipment Data Specification Sheets are given in appendix A5.11.



58

Chapter 9 Wastes 9.1 Introduction of wastes All the processes produce wastes that include indirect and direct wastes. In this chapter, only direct wastes are considered, that are waste is gaseous, liquids or solids produced directly from the process design plant. All the streams exiting the site that are not the main product are considered waste streams. The following information is about the direct wastes produced and how they are handled in the design. When waste is produced, processes must be incorporated in the design for its treatment and safe disposal. The following techniques can be considered [8]:

1. Dilution and dispersion; 2. Discharge to foul water sewer (with the agreement of the appropriate authority); 3. Physical treatments: scrubbing, settling, absorption and adsorption; 4. Chemical treatment: precipitation, neutralization; 5. Biological treatment: activated sludge and other processes; 6. Incineration on land, or at sea; 7. Landfill at controlled sites; 8. Sea dumping – (now subject to tight international control).

9.2 Classification of direct wastes Table 9.1: Wastes produced in the plant

Stream number Waste type Phase Amount of waste (tons/year) <13> Off-gas Gas 975.88<16> Wastewater Liquid 189.69<33> Waste TEMPO Solid 188.11

Remarks: (1) All the phases mentioned above should be in the room temperature. Assume the room temperature is 250C; (2) Stream <13> is the gaseous wastes stream from flash drum S01. This stream mainly contains oxygen, nitrogen and other gases in the air. Pure oxygen and nitrogen are generally neither toxic nor hazardous; hence treatment for this stream is not necessary; (3) Stream <16> contains pure water which is no toxic or hazardous to the environment; (4) Stream <33> is gas TEMPO after it comes out from the evaporator, but immediately after cooling down it becomes solid. So this stream should be dealt with as solid.

9.3 Methods of wastes treatment 9.3.1 Treatment of stream <13> Table 9.2: Treatment of waste off-gas

Waste Off-gas Effect on environment Basically no effect on environment Treatment Discharging into the environment by stack



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9.3.2 Treatment of stream <16> Table 9.3: Treatment of wastewater

Waste Wastewater Effect on environment Basically no toxic or hazardous to the environment Treatment Discharging into the environment

9.3.3 Treatment of stream <33> Table 9.4: Treatment of waste TEMPO

Waste Waste TEMPO Effect on environment TEMPO is a flammable and hazardous compound as well as

3,3-dimethylbutanol Treatment Burning Remarks: (1) It mainly contains deactivated TEMPO and 3,3-dimethylbutanol which are organic compounds, burning is chosen because it is an easy and cheapest method; (2) This waste stream can also be dealt with by regeneration, but the cost will be high.



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Chapter 10 Health, Safety and Environment 10.1 Introduction of safety Safety aspect in the design of plant relies on the application of various codes and standards of design. These represent the knowledge and experience of both experts and industry. Such application is always supported by the experience of the engineers involved in the plant. In addition, most companies admit to the fact that design personnel are under pressure to keep the project on schedule for a new plant. In this chapter, safety aspects vis à vis operating personnel from a process design point of view by using two tools: Dow Fire and Explosion Index (FEI) assessment and Hazard and Operability study (HAZOP). 10.2 Dow Fire and Explosion Index (FEI) The Dow Fire and Explosion Index is a very valuable tool in the project design stage as it makes clear which process units are to be considered as hazardous and where alternative or protective measures have to be considered. It is the most wildly used hazard index. The evaluation of the 3,3-dimethylbutyraldehyde-production plant is carried out. The calculation is in the Table 10.1. Material factor: for 3,3-dimethylbutanol MF=10 [24]; Note: 3,3-dimethylbutanol is considered the most flammable/explosive material on the plant [25], and its properties are given as below [26]: Table 10.1: 3,3-dimethylbutanol properties for FEI Compound MF Hc

(BTU/lb) NFPA Classification Flash

Point (oC) Boiling

Point (oC) Nf Nr Nh

3,3-dimethylbutanol

10 839.3723 2 0 2 41 142

10.2.1 General process hazards:

A. Exothermic chemical reactions: This is oxidizing reaction, hence factor = 0.5; B. Endothermic processes: not applicable; C. Material handling and transfer: Since for 3,3-dimethylbutanol Nf = 2, a penalty of

0.85 is applied, hence factor = 0.85; D. Enclosed or indoor process units: not applicable; E. Access: Adequate access is provided, hence factor = 0.0; F. Drainage and spill control: Adequate drainage is provided, factor = 0.0.

10.2.2 Special process hazards:

A. Toxic material: 3,3-dimethylbutanol is a toxic substance that can cause moderate irritation for eye, skin, ingestion and inhalation and Nh=2, a penalty of 0.2 Nh is used, hence factor = 0.4;

B. Sub-atmospheric pressure: The absolute pressure is more than 500 mmHg, hence this penalty not applicable;

C. Operation in or near flammable range: not applicable; D. Dust explosion: not applicable; E. Pressure: Using equation from Dow’s Fire and Explosion Index Hazard

Classification Guide [25]



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2 3

1.615030.16109 1.42879 0.5172

1000 1000 1000

X X XY

Where: Y is penalty factor; X is pressure in psig;

In this case, 10 bar is used, then X = 145.04 psig; It yields Y = 0.37; F. Low temperature: Since no unit operation is operated below ductile/transition

temperature (10C), this penalty not applicable; G. Quantity of flammable/unstable material:

The largest quantity of 3,3-dimethylbutanol in the process will be the liquid in the storage tank, which is 72385.1lbs (fill in every 10 days); The heat of combustion = 839.3723 BTU/lb; Hence the potential energy release = 839.3723 × 72385.1 = 0.061×109 BTU;

Which is too small to register from Dow Guide, factor = 0.0; H. Corrosion and erosion: Corrosion resistant materials of construction would be

specified, but external corrosion is possible, hence allow minimum factor = 0.1; I. Leakage-joints and packing: Welded joints will be used on toluene service and

mechanical seals on pumps. Use minimum factor 0.1 as full equipment details are not known at the flow sheet stage, hence factor = 0.1;

J. Use of fired heaters: There are no fired equipments in a process; hence this penalty is not applicable;

K. Hot oil heat exchange system: Heat exchange system only uses steam as heating medium, hence this penalty is not applicable;

L. Rotating equipment: not applicable. The index works out at 46: classified as “Light”. 3,3-dimethylbutanol is considered a dangerously flammable material (MF=10); the danger of material handling and transfer, and internal explosion in the storage tank and the reactor are the main process hazards. Toxicity of 3,3-dimethylbutanol would also need to be considered in a full hazard evaluation.



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Table 10.2: Dow Fire and Explosion Index Analysis

DOW FIRE AND EXPLOSION INDEX (FEI) LOCATION DATE

The Netherlands 12-Dec-02

PLANT PROCESS UNIT EVALUATED BY REVIEWED BY

3,3-dimethylbutyraldehyde Whole plant

MATERIALS AND PROCESS

MATERIALS IN PROCESS UNIT

3,3-dimethylbutyraldehyde, 3,3-dimethylbutanol, RuCl2(PPh3)3, TEMPO (2,2,6,6-tetramethyl-piperidinyl-1-oxy),

oxygen, nitrogen, water.

STATE OF OPERATION Basic Material For MF

[ ] START UP [ ] SHUT DOWN [ ] NORMAL OPERATION 3,3-dimethylbutanol

MATERIAL FACTOR ( MF ) 10

1. GENERAL PROCESS HAZARD Penalty Penalty

Used

BASE FACTOR 1.00 1.00

A. EXOTHERMIC CHEMICAL REACTIONS (FACTOR 0.30 to 0.125) 0.50

B. ENDOTHERMIC PROCESSES (FACTOR 0.20 to 0.40)

C. MATERIAL HANDLING AND TRANSFER (FACTOR 0.25 to 1.05) 0.85

D. ENCLOSED OR INDOOR PROCESS UNITS (FACTOR 0.25 to 0.90)

E. ACCESS 0.35

F. DRAINAGE AND SPILL CONTROL (FACTOR 0.25 to 0.50)

GENERAL PROCESS HAZARDS FACTOR (F1) 2.35

2. SPECIAL PROCESS HAZARDS

BASE FACTOR 1.00 1.00

A. TOXIC MATERIALS (FACTOR 0.20 to 0.80) 0.40

B. SUB-ATMOSPHERIC PRESSURE (<500 mmHg) 0.50

C. OPERATION IN OR NEAR FLAMMABLE RANGE

[ ] INERTED [ ] NOT INERTED

1. TANK FARMS STORAGE FLAMMABLE LIQUIDS 0.50

2. PROCESS UPSET OR PURGE FAILURE 0.30

3. ALWAYS IN FLAMMABLE RANGE 0.80

D. DUST EXPLOSION (FACTOR 0.25 to 2.00)

E. PRESSURE: OPERATING PRESSURE: 145.04 psig 0.37

F. LOW TEMPERATURE (FACTOR 0.20 to 0.30)

G. QUANTITY OF FLAMMABLE / UNSTABLE MATERIAL:

QUANTITY: 7238.51lbs Hc = 839.3723BTU/lb

1. LIQUIDS, GASES AND REACTIVE MATERIALS IN PROCESS

2. LIQUIDS OR GASES IN STORAGE

3. COMBUSTIBLE SOLIDS IN STORAGE DUST IN PROCESS

H. CORROSION AND EROSION (FACTOR 0.10 to 0.75) 0.10

I. LEAKAGE-JOINTS AND PACKING (FACTOR 0.10 to 1.50) 0.10

J. USE OF FIRED HEATERS

K. HOT OIL HEAT EXCHANGE SYSTEM (FACTOR 0.15 to 1.50)

L. ROTATING EQUIPMENT 0.50

SPECIAL HAZARD FACTOR (F2) 1.97

UNIT HAZARD FACTOR (F1 x F2 = F3) 4.63

FIRE AND EXPLOSION INDEX (F3 x MF = F&EI) 46



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10.3 Hazard and Operability study (HAZOP) Hazard and Operability study is a procedure for the systematic, critical, examination of the operability of a process. When applied to a process design or an operating plant, it indicates potential hazards that may arise from deviations from the intended design conditions. In this plant, the critical equipment that requires the most attention is reactor R01, because it is operated at high operating temperature, the reaction is exothermic, and 3,3-dimethylbutanol a flammable substance is present. Main material in: 3,3-dimethylbutanol, air, Ru-100, TEMPO.

Guide Word Deviation Possible Causes Consequences Action Required Line number: No. 6 Intention: Transfers 3,3-dimethylbutanol to reactor No

No Flow 1) No 3,3-dimethylbutanol available in storage tank 2) Pump P01 fails (motor fault, loss of drive, impeller corroded away etc) 3) Line blockage, isolation, FC valve closes in error or fails shut 4) Line fracture

Loss of feed to reactor and reduce output As for 1) As for 1), pump P01 overheats As for 1), 3,3-dimethylbutanol leaks to the ground

Ensure good communications with storage tank operator; Install low level alarm at storage tank As for 1) Install kickback on pump P01; Check design of pump strainers Institute regular patrolling and inspection of transfer line

More

More Flow 5) FC valve opens too wide in error

Reactor overfills; Incomplete oxidation of 3,3-dimethylbutanol

Install high level alarm

More Temperature

6) Thermal expansion in an isolated valve section due to fire or strong sunlight

Line fracture or flange leak; High pressure in transfer line and reactor

Install temperature control TC and cooling jacket for the reactor; Install a TI on line No. 6



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7) High temperature of incoming 3,3-dimethylbutanol (from recycle)

High reactor temperature

As for 6)

More Pressure

8) High pressure of incoming 3,3-dimethylbutanol (from recycle)

High reactor pressure

Install a PI on line No. 6

Less

Less Flow 9) Leaking flange or valve

Low rate of production

As for 3)

Line number: No. 32 Intention: Transfers compressed air to reactor No

No Flow 1) Compressor failure 2) FC valve fails close

Possible dangerous 3,3-dimethylbutanol concentration; Loss of feed to reactor and reduce conversion As for 1)

Install low level alarm PAI interlocked to shut down 3,3-dimethylbutanol flow As for 1)

More More Flow

3) FC valve opens too wide

It might bring to explosion

Check the valves periodically to ensure that it works as expected

More Temperature

4) No cooling water at storage

High reactor temperature

Ensure enough cooling water supplied for line No.32

Less Less Flow 5) Compressor partial failure

As for 1) As for 1)

Reverse Reverse Flow

6) Fall in line press (compressor fails); high pressure at reactor

3,3-dimethylbutanol in compressor – explosion hazard

Fit non-return valve (NRV1)

Line number: No. 9 Intention: Transfers mixed catalysts (Ru-100 and TEMPO) to reactor No

No Flow 1) No catalysts available in storage tank 2) Pump P06 fails

Lack of catalysts causes no reaction in the reactor As for 1)

Ensure good communications with storage tank operator; Install low level alarm at storage tank As for 1)



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3) Line blockage, isolation, FC valve closes in error or fails shut 4) Line fracture

As for 1) As for 1), 3,3-dimethylbutanol leaks to the ground

Install kickback on pump P02; Check design of pump strainers Institute regular patrolling and inspection of transfer line

Less Less Flow 5) Leaking flange or valve

Low rate of production

As for 3)

Less pressure

6) Pump P02 fails, low pressure of incoming catalysts

Low reactor pressure

Install a PI on line No. 9

Line number: Cooling water line Intention: Maintains temperature for line No. 32 No

No Flow 1) No cooling water at storage

Reactor temperature will increase and it might lead to explosion

Ensure good communication with operator

More More Flow 2) Valve opens too wide

Operating costs will be increased because of too much cooling water

Ensure the temperature control is working properly; Check the valves periodically to ensure that it works as expected

For separation part (distillation columns), everything is under control (please refer to chapter 6 of process control for explanation why a specific controller is put there). If the controllers are working properly then it will cause no damage. If some controllers are broken, then they should be replaced to keep the process running safely. Recommendation Basically, the main cause of hazard is whether or not the valve and controllers are working properly. Hence in order to keep the process inherently safe, checking those equipments regularly is recommended.



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Chapter 11 Economy

This chapter provides economic evaluation of a conceptual process design of the plant producing 3,3-dimethybutyraldehyde.

11.1 Capital investment costs The table below presents the capital investment costs of the plant Table 11.1: Capital investment cost

Capital

Description % IN (Dfl) IN (Euro) ( 106) ( 106) Fixed Capital (FC) 89.0 7.79 3.54 Working Capital (Approximate as11% of total investment) 11.0 0.96 0.44

Total investment 100.0 8.75 3.98 Remarks: - Refer to appendix A6.3 for details of fixed capital calculation - The working capital is assumed to be 11% of total cost. - The calculation is in Dutch gulden as well as Euro as standard currency. The ROE

between Dutch gulden and Euro is 2.2. 11.2 Annual income Table 11.2:Annual income of sales of products and sellable by-products

Products

Description OUT Sales IN (Euro/a) (kg/a) (Euro/kg) ( 106) Product

3,3-dimethybutyraldehyde 1,067,900 143.5 153.24By products

Deactivated Ru-catalyst 201,485 692.0 139.43

Total / Gross Income 292.67 Remarks: - Refer to tables 3.6 and 3.7 (chapter 3) for the amount of product and by-products; - Refer to table 3.11 (chapter 3) for the prices of products .the price of deactivated Ru catalyst is estimated to 70% of that of Ru catalyst; - The waste gas produces only when there are leaks in the distillation column; - Assume the leaks in distillation column seldom found, and if there is some waste gas is produced, the treatment is to burn it to CO2 and H2O and then release them to the atmosphere. The cost of this treatment can be negligible (refer to chapter 9 for waste treatment).



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11.3 Operation costs 11.3.1 Raw material costs Table 11.3: Costs of raw materials

In going streams

Description IN Purchase cost OUT (Euro/a)

(kg/a) (Euro/kg) ( 106)

Raw materials 3,3-dimethylbutanol 1,123,200 24.4 27.41 Catalysts

Ru-catalyst 201,485 989.0 199.27 TEMPO 98,496 30.8 3.03

Total 1,423,181 229.71

Remarks: - Refer to part 2.3, chapter 2.

11.3.2 Utilities costs Table 11.4: Costs of utilities

Utilities

Description UnitsIN

(Units/a) Cost

(Dfl/Unit) OUT

(Dfl/a) OUT

(Euro/a)

Steam t 2,329 30.00 69,880 31,763Cooling water t 211,706 0.10 21,171 9,623Electrical power kWh 542,333 0.22 119,313 54,233

Total 210,364 95,620Remarks:

- Refer to part 5.5, chapter 5 for the quantity of utilities



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11.3.3 Summary of production costs Table 11.5: Summary of annual production (manufacturing) costs

Summary of annual production (manufacturing) costs (kEuro/a)

Cost type

kEuro/a

%

Remarks

Direct Variable (A)

1 Raw materials 229,710 99.94% 79.14% 2 Miscellaneous materials 40 0.02% 0.01% (5): kEuro 3980 10%3 Utilities 95 0.04% 0.03% Refer to part 5.5, chapter 5 4 Shipping & packaging Can be always negligible Sub-total 229,845 100% 79.19% Fixed (B)

5 Maintenance 398 17% 0.14% Fixed Cap.: kEuro 3980 10%

6 Operating labour 600 25% 0.21%9 oper.+2QC+1 for other job= 12 operators

7 Laboratory 138 6% 0.05% (6): kEuro 600 23%8 Supervision 120 5% 0.04% (6): kEuro 600 20%9 Plant overhead 300 13% 0.10% (6): kEuro 600 50%

10 Capital charges 597 25% 0.21% Fixed cap.: kEuro 3980 15%11 Insurance 40 2% 0.01% Fixed cap.: kEuro 3980 1%12 Local taxes 80 3% 0.03% Fixed cap.: kEuro 3980 2%13 Royalties 80 3% 0.03% Fixed cap.: kEuro 3980 1% Sub-total 2,353 100% 0.80% Total 232,198 80.00% (A)+(B)

Other (C) 14 Sales expenses 15 General overhead 58,050 20.0% (A)+(B): kDfl 232,198 25 %16 Research & Dev. Total Production Costs Annual [MEuro/a] 290.25 100% (A)+(B)+(C) Per ton P [Euro/kg] 271.80 1067900kg/a

Note: - This table follows J. M. Coulson & J. F. Richarson’s Chemical Engineering

textbook (vol.6, p232, table 6.6)

There are 12 operators needed: 9 operators, 2 quality control engineers and one worker. The responsibility of the 12 workers is to operate the production line including reaction part and separation part. The reasons for having 3 operators are: - Firstly, at the beginning of the process raw material and catalysts are added

manually (refer part 3.2.2, chapter 3 for process description). It is supposed that this step requires two operators. At the same time it is still necessary to have at least another operator working in the control room. Therefore, the minimum number of operators required is 3 per shift. In Europe normal working hour is 8 hours per day, two quality control engineers’ responsibility is to ensure the quality of the product is under control. Another worker is needed to do other jobs like cleaning or maintaining equipment.



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11.4 Gross income, net cash flow and economic criteria The following assumptions are made to determine maximum allowed investment: - The economic plant life of project is 11 years - The design/building period is 1 year and the working period is 10 years - The interest rate (r) is 8% per year

Table 11.6: Gross income; net cash flow, economic criteria

Gross income, net cash flow, economic criteria

Item Unit Value kEuro/a Remarks Gross Income 292,670

Production Costs 290,250

0 Net Cash Flow (Before Tax) 2,420 = (A) Economical Plant Life & Depreciation - Total Investment kEuro 3,980 = (B) - Econ. Plant Life

year

11

Incl. 1 yrs Des.& Con.

- Annual Depreciation over 11 years 1,120 Net Cash Flow (After Depreciation) 12,256 - Income Tax @ 45% 5,515 Net Cash Flow (After Tax) 6,741 Pay-Out Time (Before Tax) year 1.64 = (B) / (A) Rate of Return (Before Tax) % 60.8% = (A) / (B) DCF Rate of Return (Before Tax) % 60.3% nil DCFROR Net Present Value (Before Tax) % 8.0% 12,256 Appendix A6.3

Net Future Value (Before Tax) kEuro 97,280 Interest = 0 Note: - Refer to appendix A6.3 for the details of Discounted Cash Flow Rate On Return, Net Present Value and Net Future Value’s calculation.

11.5 Cost review

Table 11.5 shows that variable cost is 79%, and other cost is 21% because during the calculation other cost is calculated as 20-30% of the direct production cost [15]. It is also shown in table 11.5 that raw material costs 99% of total variable cost. Catalyst cost account for the most part among that. There are two possibilities to reduce the cost of catalyst:

- Firstly, the possibility to reduce the cost is dependent on the progress of laboratory work. Better ratio between Ru-catalyst and raw materials needed might be a solution. New catalyst, which has cheaper price, can be studied as an alternative.

- Secondly, the reaction time can be prolonged to reduce the amount of catalyst needed. However, justification is needed for this reasoning. Also prolonging the time of the reaction might be not a wise solution since time is money.



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11.6 Sensitivities The most sensitive factor is catalyst price among investment, utilities cost, raw material cost and product price. If the cost price increases 10% the plant will have deficit even NCF has negative value. The second sensitive factor is product price among investment, utilities cost, raw material cost and product price. If the price decreases 20% the plant will have deficit even NCF has negative values. For other factors fluctuations are can be disregarded as sensitive factors. 11.7 Discussion As it shows in table 11.6, the net present value is 12.56 million Euro with the interest of 8% and net future value is 97.28 millions. The total investment of the plant is 3.6 millions. The pay out time is only after 2 year. It can be seen that the process in very promising in term of profits and it is quick to get the money back. The values obtained for economic evaluation in this chapter can be said to be reliable because the highest costs for raw materials are taken for the calculation. However, this process is very sensitive with price of catalyst and product price.



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Chapter 12 Creativity & Group Process Methods 12.1 What is creativity? Creativity is generally regarded to involve breaking the kind of rigid rules standing at the heart of logic. Creativity, on the one hand, of the artistic variety, is commonly identified with the emotions and the ''irrational". Freud, whose specific claims are today a bit tenuous, remains a seminal figure often getting at least the tenor of things right. Freud believed that creativity is the link between art and play, and requires the ''the suspension of rational principles". He wrote that '"The creative writer does much the same as the child at play. He creates a world of phantasm which he takes very seriously that is, which he invests with large amounts of motion-while separating it sharply from reality". Creativity, on the other hand, within the technical field, must be consistent with the reality. It is based on intuition, and maybe seems unstructured or visionary, but indeed, should coordinate well with consistency, which is based on various numbers and clear concepts. Process and plant design is the creative activity whereby ideas are generated and then translated into equipment and processes for producing new materials or for significantly upgrading the value of existing materials. Thus, creativity in design is essential for technical progress. A lack of creativity will at best only result in "run of the mill" solutions. Therefore creativity needs to be an integral part of the design process. 12.2 Creativity and process design Creativity is a quality that is highly valued, but not always well understood. Those who have studied and written about it stress the importance of a kind of flexibility of mind. Studies have shown that creative individuals are more spontaneous, expressive, and less controlled or inhibited. They also tend to trust their own judgment and ideas-- they are not afraid of trying something new.

A common misunderstanding equates creativity with originality. In point of fact, there are very few absolutely original ideas. Most of what seems to be new is simply a bringing together of previously existing concepts in a new way. Psychologist and author Arthur Koestler referred to this merging of apparently unrelated ideas as dissociation. The fact that creative thinking is based on knowledge of previous work in one's field is the justification for teaching the history and foundations of a given field as a resource for future research and creative work. It is possible to develop ones ability to think intuitively and creatively.

Thus creativity is the ability to see connections and relationships where others have not. The ability to think in intuitive, non-verbal, and visual terms has been shown to enhance creativity in all disciplines. It has also been shown that the creative process is very similar in all fields.

Essentially the design process is a problem-solving process, and the designer, just like the laboratory scientist, will be most successful if the problem is approached in a systematic manner. Successful fine artists generally follow the same pattern in developing their



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creative ideas, though they may be less conscious of the process they are following. Initially the researcher or designer/artist will tend to experiment in a rather random manner, collecting ideas and skills through reading or experimentation. Gradually a particular issue or question will become the focus of the reading and experimentation. The next step is to formulate a tentative problem, and begin to explore that topic. Eventually the problem is refined into a research question or design problem that the person will then pursue through repeated experimentation. In design or fine arts production, this takes the form of works created in a series. Each effort solves certain problems, and suggests issues to be dealt with in the next work (or experiment). Working in a series is the most important stage of the design process. The ability to experiment, to value and learn from mistakes, and build on the experience achieved is the hallmark of a truly successful and creative individual, whatever the field. 12.3 Why should creativity be utilized in the design? Our design problem is greatly under defined, i.e. only a very small fraction of the information needed to define a design problem is available from the problem statement. To supply the missing information, assumptions must be made:

----what is our target product(s), and what will be the production per year? ----what types of process units should be used? ----how those units will be interconnected? ----what temperatures, pressures, and process flow rates will be required? ----what is the component of the waste? ----where should the plant be located at?

All those assumption can only be made by utilizing great amount creativity, otherwise, it would be much more difficult to fill in missing information. 12.4 What kind of creativity is utilized in the design? In order to made the assumption and realize the design, two types of creative activity can be distinguished: adaptive & inventive. Adaptive creativity involves the adept ion and extension of existing knowledge to a new situation. On the contrary, inventive creativity, which is purely original, involves radical new fundamental principles or methods in order to achieve an existing or new function. In our design, both of the two types of creativity are utilized, but mainly adaptive. There are many creativity methods can be and have been used in the design activity, such as:

a. brainstorming b. +/- evaluation c. diary with ideas, solutions, new solutions d. discussion

----about contradictory elements ----on improvements of mistakes, on more direct communication, on

productivity and on participation e. making more conscious one’s latent knowledge



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f. methods for improved group work (co-evolution of problem solution, task distribution, group management, group rules)

g. symbolizing the main elements of the process-installation (i.e. for reactors, pumps, valves, cooling, molecule- and gas- combinations); analogy of the installation with human organization, patterns of interaction, competitive games etc.

h. multiple solutions: ----exploring alternative solutions, and values of different approaches ----exploring solutions applied in different, but comparable situations

i. reporting how mistakes have been found and improved j. use of more outside- or inside-information; and of more experience in practice k. use of incubation-phase in problem solving: after thorough exploring and

searching, leaving the problem for some time without strong control (and without feeling of dissatisfaction)

l. exploration (and reporting) of critical moments or periods during the work-process; their meaning; and how they were solved

m. explanation why quotes, statements and formulas were used n. use adequate and safe way of capturing, storing and reusing information and

knowledge o. employing computer-based simulations(i.e. Aspen).

There seems to be some cross-links between the different items. Actually, it is the truth that no single creativity method is far from the others, and is capable to solve the problem solely. Only good combination of various creativities can lead to good design. During the design, all the creative methods were used except the first one — brainstorming. How these methods were used is reported in 12.7. 12.5 Rules to utilize creativity in the design To design a chemical plant utilizing a “green” technology, some rules [28] can be followed:

1. A design methodology delivers a structural framework of design methods to make the designer's life easier.

2. The process designer needs a lot of bounded free space in the beginning of the design to ensure innovation and creativity.

3. Least information must be necessary to decide in the first stage of design. 4. The hardest part of the conceptual design of sustainable chemical plants is the

implementation of soft criteria. 5. The basic chemistry of sustainable plants is Green Chemistry.

12.6 How to improve creativity during the design? At the first stage of the design, it was quite difficult to utilize many creativity tools because none of the group member had this kind of experience. Meetings were hold by the group members and the creativity coach to improve the creativity, understanding of creativity and ability to utilize creativity into the design of the group members. Discussions on the understandings of creativity methods helped the members to get further knowledge on what these methods are and how to use them.



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12.7 How creativity is utilized in the design? A selection of these activities and results regarding creativity, group process tools, will be reported here as follows. 12.7.1 Literature search This work was done in the first week after the kick-off meeting. The main objective of literature searching is:

1. to find similar process in the industry as the design, and try to find their disadvantages

2. to look for information on the present synthesis method of the candidate products

3. to look for physical and chemical properties of the chemicals involved in the process

4. to look for economic data for the process investment 5. to find any information related to background of the design

In order to get as much useful information as possible in very short time duration (one week), high work efficiency must be achieved. Thus it was done in a stepwise way.

1. list the objective information that is needed 2. distribute the tasks to individuals 3. useful information found (i.e. related data, summary of related papers, etc.) be

uploaded to BLACKBOARD-File Exchange 4. group discussion on the information found 5. delete useless ones upon the agreement achieved during the group discussion 6. compile summary on literature

The literature searching was done by searching Internet, by looking into Encyclopedia, as well as contacting with experts in industry. By the stepwise processing, much enough information was successfully found to the next step (Preliminary Basic of Design). Of course, some information still has to be found, but the literature searching will not be a major task in the following weeks. In this stage, some creativity methods such as, +/- evaluation (on usability of the literature), task distribution, adequate way of capturing and storing information, etc., were used. 12.7.2 Determine the target product This task was done during the literature searching. The truth was identified that it was difficult to say which aldehyde is the most profitable product by using this Ru-based catalytic oxidation method without getting insight knowledge on the synthesis process of the candidate products. So it was decided to determine the product (or product distribution) after enough information on the present synthesis method was found. So during the week for literature searching, based on the knowledge got from the present synthesis process (the utilization of the products, economic issues, environmental concerns, and availability of the feedstock, etc.), 3,3-dimethylbutyraldehyde was selected as the target product. And after that, more information focus on the target was searched.



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In this stage, the creativity methods such as, group discussion, use of more outside- or inside- information, use of incubation-phase in problem solving (after reading some literature, not to make decision immediately, but leave it for one or two days without strong control), etc., were used. 12.7.3 Determine the production After determination of the target product and before the definition of the process, the production must be fixed. Since it was not easy to find a similar industrial process as the design, it was not possible to determine the production by industrial example. Then another way was tried. Our product, 3,3-dimethylbutyraldehyde, is an important intermediate in the synthesis of Neotame, which is a nonnutritive sweeter. So from the market demand of sweeter, the market demand of 3,3-dimethylbutyraldehyde was deduced. In this stage, exploring alternative solutions are one of the creativity methods that were utilized. 12.7.4 Process concept chosen This is the most important part of the BOD. From literature, provided documentation and/or the use of creative methods various process options are (made) available, all having pro’s and con’s. Creative methods were used to create additional options by searching for solutions for the negative aspects of some options. From these options, one option was selected as a basis for further process design. Decision criteria for the design were determined to evaluate all the options. Discussions were hold to develop the adequate chemical routes, options between batch and continuous process, and the reaction conditions according to kinetics and thermodynamics of the reaction. Controversies were emerged, on many aspects, such as the residence time in the reactor. It was eventually decided to use a residence time of 3 hours according to the economic and environmental criteria. Advice from expertise was also taken into account. In this stage, creativity methods such as, diary with ideas and solutions, discussion, reporting mistakes, use of more outside- or inside- information, etc., were used. 12.7.5 Developing PFS The PFS (Process Flow Scheme) is an essential document, not only for “Process Administrator”, but also for transfer of know-how about the process design to others. It is the heart of the design. It was developed from the basic block scheme and represents the line-up of the various unit operations (equipment) and their connection. The development of PFS is not a stage which is separated from other design activities, but an essential clew to solve the whole design problem. At the earlier stage of the design, a simple PFS was developed by putting some unit operations into the basic block scheme. It can be used as a model for mathematic simulation. According to some results of the simulation, some unit operations were modified. An update PFS was developed and was used as the basic for the design of process equipment and process control. Then all the information including equipments, heating and control system and the stream data were put into the PFS. This PFS is the final Process Flow Scheme for the whole design. It was a highly integrated work of the whole group and the whole design.



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In this stage, creativity method such as, co-evolution of problem solving, making more conscious one’s latent knowledge, symbolizing the main elements of the process-installation (i.e., for the reactor, symbolizing it as an cooking pot, a tree, or a human body), analogy the installation with human activity (i.e., analogy the chemical process design with a chair design), were used. 12.8 Group rules Group work can only be processed smoothly and successfully when everything is in good order and some rules are always followed by the group members. Below are our group rules:

1. Working 8 hours for the design project everyday; 2. English should be the only working language to be used between the group

members for discussion and communication; All the exchanged files, report components and information should be written in English;

3. Holding group meeting at least twice per week; meeting with coach is proposed once per week;

4. Trying to show up in all the meetings and contribute good idea; 5. Finishing individual tasks on time; 6. Trying all the possible ways to solve when problems are encountered; 7. Working cooperatively and encouraging each other to overcome difficulties; 8. Go ahead, never give up (the most important).

During the design, all the members of our group complied with the group rules strictly, which is a good base for the success of the design.



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Chapter 13 Conclusions and Recommendations 13.1 Conclusions The plant of production scale of 1000 tons/year of 3,3-dimethylbutyraldehyde from the oxidation of 3,3-dimethylbutanol is successfully designed. The chemical route of the process has been developed by Biocatalysis and Organic Chemistry section of Delft University of Technology (chapter 2). This route is considered as a novel catalytic method for alcohol oxidation in term of economical potential and yields (chapter 1). The plant in design consists of two parts: reaction part and separation part (chapter 2,3). The design of the plant is based on the assumptions, key data for the design (chapter 3), thermodynamic models and properties (chapter 4). It results in the process structure (chapter 5) and design of equipment of the plant (chapter 8). The methods of treatment for direct wastes produced in the plant are selected to make the plant environmental more friendly. Since off-gas stream mainly contains oxygen, nitrogen and other gases in the air, it almost has no effect on environment, discharging is chosen. In wastewater stream only a small amount of organic compounds is produced, harmful smoke will not generated after burning. Since the cost for regeneration of the catalyst is very high, burning is considered for the treatment of deactivated TEMPO. The safety aspects are assessed (chapter 10) and the Dow Fire and Explosion Index is found to be 46, which means that the degree of hazard of the plant is only light. Several companies are found to produce 3,3-dimethylbutyraldehyde by using different ways from what we used. Therefore the product is currently available in the market. However, it is expected to have a bigger market for this aldehyde due to its industrial use and economical potential (chapter 1). It can be seen from the economic analysis (chapter 11) that total capital investment of this plant is 3.98 million euros, and rate of return is 60.8% with pay out time is equal to 2 years. In term of rate of return and pay out time, this project is very promising. It is believed that the product of the plant in the design will be able compete in the market of 3,3-dimethylbutyraldehyde since the market is expected to be bigger. Therefore, it can be concluded that this plant is very potentially and economically feasible to implement. 13.2 Reliability of the design The process can be considered robust and reliable because appropriate controllers are designed to maintain the desired process conditions (chapter6). However some deviations might occur during operation since perfect operation can be rarely achieved. From Hazard and Operability study (HAZOP) analysis, checking valves and controllers regularly is recommended to ensure that the process is running safely. 13.3 Recommendation and suggestions 1) Ru catalyst has a very complex structure. It cannot be easily implemented by ASPEN simulation. For this reason, some extra calculations need to be done by hand. To get a sensible result, much computation work has been done. This can be avoided if we could master more knowledge about ASPEN.



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2) There are several hot streams and cool streams. From the energy and economical point of view. These streams can be combined each other in pairs. 3) The Ru-based catalyst is considerable expensive. If we can decrease the demand of this catalyst and on the other hand, the production is not decrease so much. This process will be more profitable. 4) Regarding to the relationship between reactor residence time and the operating cost, it’s possible to find an optimal point to get the maximum production. 5) The deactivated rate of Ru catalyst and TEMPO is deferent [2] (Ru catalyst deactivates faster than TEMPO). In order to guarantee the sufficient inlet of Ru catalyst, in fact, the TEMPO we used is more than the stoichiometric number (Ru-100: TEMPO=1:3 mol/mol) [2]. Unfortunately, the price of TEMPO is not cheap, either. To prevent from wasting TEMPO, the inlet TEMPO can be programming-injected by a in-line concentration detector and computer. In this case, we can adjust the Ru-100/TEMPO ratio, and eventually, reach the maximum usage of TEMPO. 6) It is very attractive if we can change the homogeneous catalyst into heterogeneous one. In this case, a much convenient technology can be adopted such as easy catalyst separation, easy catalyst regeneration. Moreover, the price of catalyst can be less expensive.


Lin Luo (1113259) Appendix 1- 1 Weimin Wang (1118668) Shuang Zhao (1118447) Zhengjie Zhu (1113658)

Appendix 1

Block Scheme And Process Options

Figure A1.1 Block Scheme Option 1



Figure A1.2 Block Scheme Option 2



Figure A1.3 Block Scheme (Chosen)



Appendix 2

ASPEN Simulation A2.1 Introduction In this design, all the units in block scheme (except Evaporator separating Ru catalyst from TEMPO) have been implemented through ASPEN PLUS. After simulating, the generated data are used for designing the process equipment. In this section, the detailed specifications of ASPEN simulation and generated results are provided. All the results can also be seen in the attached ASPEN files as well. A2.2 Component specification There’re ten components in the design (see the table below). Five of them, namely hexane, hexane, hexanal, water, oxygen and nitrogen can be easily found from the databank, while the rest are not. However, the properties of 3,3-DIMETHYLBUTANOL, 3,3-DIMETHYLBUTYRALDEHYDE and TEMPO can be estimated by ASPEN. With the required property parameters, the User-Defined Component Wizard is used. Unfortunately, the Ru-100 catalyst could neither be found from databank nor defined by the User-Defined Component Wizard. It’s excluded from the ASPEN simulation. The below table indicates the components we used along with their IDs, types and formula.

Table A2. 1 Specification of components in ASPEN PLUS

Component ID Type Component Name Formula HEXANE Conventional N-HEXANE C6H14-1

HEXENE Conventional 1-HEXENE C6H12-3

HEXANAL Conventional 1-HEXANAL C6H12O-D2

WATER Conventional WATER H2O

TEMPO Conventional TEMPO C9H18NO

E Conventional 3,3-DIMETHYLBUTANOL C6H14O

A Conventional 3,3-DIMETHYLBUTYRALDEHYDE C6H12O

O2 Conventional OXYGEN O2

N2 Conventional NITROGEN N2

RU Solid RU-100 RuCl2(PPh3)3 According to Chapter 6 of ASPEN handbook (ASPEN PLUS 11.1 User Guide), we need to specify the unknown components by means of giving information of molecular structure. The follows tables are relevant to that.



TEMPO

Table A2. 2 Molecular Structure of TEMPO

Molecular structure Atom 1 Atom 2 Bond type

1 C 2 C Single bond2 C 3 C Single bond3 C 4 N Single bond4 N 5 C Single bond5 C 6 C Single bond6 C 1 C Single bond3 C 8 C Single bond3 C 9 C Single bond5 C 10 C Single bond5 C 11 C Single bond4 N 7 O Single bond

3,3-DIMETHYLBUTANOL

Table A2. 3 Molecular Structure of 3,3-DIMETHYLBUTANOL


1 C 2 C Single bond2 C 3 C Single bond3 C 4 C Single bond4 C 5 O Single bond2 C 6 C Single bond



2 C 7 C Single bond 3,3-DIMETHYLBUTYRALDEHYDE

Table A2. 4 Molecular Structure of 3,3-DIMETHYLBUTYRALDEHYDE


1 C 2 C Single bond 2 C 3 C Single bond 3 C 4 C Single bond 4 C 5 O Double bond2 C 6 C Single bond 2 C 7 C Single bond



A2.3 Thermodynamic methods ASPEN PLUS distinguishes three methods of calculating phase equilibria: The equation of state method, the activity coefficient method and a special application method.

The equation of state (EOS) method

The equation of state method is used to describe both liquid and vapor phase behavior. This method is applicable for systems where the interaction of the components in the liquid phase is assumed to be minimal. This is for most hydrocarbon systems at high and moderate pressure and temperature. Typical equation of state are: Redlich-Kwong (RK) Soave-Redlich-Kwong (SRK) Peng Robinson (PR)

The activity coefficient method

The activity coefficient method is a combination of two different equations: one to calculate vapor phase properties (j

v, EOS/IG), methods Nothnagel and Hayden-O’Connell with Henry are effective. The other one is to describe liquid phase behavior (j in the unsymmetrical form). This method is widely used in the highly non-ideal mixtures, which are comprised of acetone, alcohol, ester, ether, etc. Especially, high accuracy would be obtained for (chemical) systems with dissimilarity in size, shape and intermolecular forces. Since the chemicals involved in the system are raw material, 3,3-dimethylbutanol with impurity, n-hexane and 1-hexene, and products, 3,3-dimethylbutyraldehyde and water, it is a non-ideal mixture with intermolecular forces, the activity coefficient method should be applicable to model vapor/liquid equilibrium. Thirdly, the special application method is available for systems that for some reasons cannot easily fall into one of the other two methods. In the activity coefficient method, typical equations model liquid phase behaviors are: Wilson NRTL Uniquac Van Laar UNIFAC

Wilson and Van Laar are only suitable for binary vapor-liquid systems which interaction is small, while NRTL and Uniquac can also be used in cases that there are two liquid phases present. NRTL and Uniquac have better accurate in systems containing water than Wilson, however contrary result is showed for vapor-liquid system. None of these models should be used under critical conditions. A special, widely used activity coefficient model is UNIFAC. This method can predict phase equilibrium for systems for which no experimental data are available. It doesn’t depend on binary interaction parameters, instead bases upon the so-called



group contribution method. And a modified NRTL method, NRTL-HOC is used Hayden-O’Connell method to calculate fugacity coefficient.

Special application method Special application methods are available for systems that for some reasons can not easily fall into one of the other two methods

In this design, since the thermodynamic data for raw material-3, 3-dimethylbutanol 3,3-dimethybutyraldehyde and TEMPO are hardly available in database. Due to these components are not hydrocarbons and all of them are polar molecules, the Wilson thermodynamic method is chosen to predict vapor/liquid equilibrium.



A2.4 Equipments A2.4.1 Compressor Compressor simulates a: Polytropic compressor Polytropic positive displacement compressor Isentropic compressor Isentropic turbine

Use Compressor to change stream pressure when energy-related information, such as power requirement, is needed or known. Compressor can handle single-phase as well as two- and three-phase calculations. Compressor can also calculate compressor shaft speed, but cannot handle performance curves for a turbine One can use Compr to rate a single stage of a compressor or a single wheel of a compressor, by specifying the related performance curves. Compressor allows to specify either: Dimensional curves, such as head versus flow or power versus flow Dimensionless curves, such as head coefficient versus flow coefficient

C01

The vapor flow (stream 4) has a pressure of 1 bar that is not high enough to be fed into the reactor (10 bar). The compressor C01 is used to increase the pressure.

Figure A2. 1

Table A2. 5 Inputs of C01

model ASME-POLYTROP discharging pressure 10.5bar

Table A2. 6 Results of C01

Compressor model: 5 Phase calculations: 9 Indicated horsepower: 90180.70Watt



Brake horsepower: 90180.70Watt Net work required: 90180.70Watt Power loss: 0Watt Efficiency: 0.72 Mechanical efficiency: Outlet pressure: 1050000N/sqm Outlet temperature: 714K Isentropic outlet temperature: 565K vapor fraction 1



A2.4.2 Decanter Decanter models knock-out drums, decanters, and other single-stage separators with sufficient residence time for separation of two liquid phases but without a vapor phase. It determines the thermal and phase conditions of a mixture with one or more inlet streams, at the specified temperature or heat duty. Decanter can calculate liquid-liquid distribution coefficients from:

Physical property method User supplied distribution correlation User supplied Fortran subroutine

Decanter can perform: Liquid-liquid equilibrium calculations Liquid-free-water calculations

Use Decanter to model knock-out drums, decanters, and other single-stage separators without a vapor phase. When you specify outlet conditions, Decanter determines the thermal and phase conditions of a mixture of one or more inlet streams. Decanter can calculate liquid-liquid distribution coefficients using: An activity coefficient model An equation of state capable of representing two liquid phases A user-specified Fortran subroutine A built-in correlation with user-specified coefficients

D01 The stream 15 from the flash drum S01 is a mixture of many components. Because the water generated by the oxidation does not solute with the Alcohol or Aldehyde. It separates into two phases, namely the aqueous phase and the organic phase. This decanter is used to separate these two phases. The stream 16 is water to be discharged and the stream 17 is the organic phase.

Figure A2. 2



Table A2. 7 Inputs of D01

Temperature 298K Pressure 1Bar key components Water, E component molar fraction 0.0001 Separation Efficiency for each componentsHEXANE HEXENE HEXANAL WATER 99 TEMPO 90 E 90 A 90 O2 N2 Valid phase Liquid - water

Table A2. 8 Results of D01

Outlet Temperature 298 K Outlet Pressure 1 bar Heat duty -0.0096 Watt Net duty -0.0096 Watt 1st/ total liquid 0.85



A2.4.3. Heat exchanger Heater performs these types of single phase or multiphase calculations: Bubble or dew point calculations Add or remove any amount of user specified heat duty Match degrees of superheating or subcooling Determine heating or cooling duty required to achieve a certain vapor

fraction Heater produces one outlet stream, with an optional water decant stream. The heat duty specification may be provided by a heat stream from another block. Use Heater to model: Heaters or coolers (one side of a heat exchanger) Valves when you know the pressure drop Pumps and compressors whenever you do not need work-related results

When one specifies the outlet conditions, Heater determines the thermal and phase conditions of a mixture with one or more inlet streams. H01

Use the cooler H01 to cool the hot stream 10 down to a lower temperature so as to be fed in the flash drum S01.

Figure A2. 3

Table A2. 9 Inputs of H01

Temperature 298K Pressure 1bar Valid phase V-L

Table A2. 10 Results of H01

Outlet Temperature 298 K Outlet Pressure 1 bar Vapor fraction 0.75 Heat duty -51791.83 Watt Net duty -51791.83 Watt 1st/total liquid 1 Pressure drop correlation parameter 22862608.6



H02 Heat the stream 17 from 298 K to 400 K

Figure A2. 4




Outlet Temperature 400 K Outlet Pressure 1 bar Vapor fraction 0.00286 Heat duty 39142.43 Watt Net duty 39142.43 Watt 1st/total liquid 1 Pressure drop correlation parameter 0 H03 Cool the stream 26 from 428K to 333K as a feed of distillation column S03

Figure A2. 5


Temperature 333K Pressure 1bar Valid phase L-only




Outlet Temperature 333 K Outlet P 1 bar Vapor fraction 0 Heat duty -93002.37 Watt Net duty -93002.37 Watt 1st/total liquid 1 Pressure drop correlation parameter 0 H04 Cool the stream 7 to 298K

Figure A2. 6


Temperature 298K Pressure 10.5bar Valid phase V-L


Outlet Temperature 298 K Outlet Pressure 10.5 bar Vapor fraction 0.99 Heat duty -92505.86 Watt Net duty -92505.86 Watt 1st/total liquid 1 Pressure drop correlation parameter 0



A2.4.4 Mixer & Splitter Mixer combines material streams (or heat streams or work streams) into one outlet stream. If material streams are mixed, one can use an optional water decant stream to decant free water from the outlet. One can specify an outlet pressure or pressure drop for material streams. The mixer model determines the combined outlet stream temperature and phase condition by performing an adiabatic phase equilibrium flash calculation on the composite feed streams. Mixer can be used to model mixing tees, or other types of stream mixing operations. Select the Heat (Q) and Work (W) Mixer icons from the Model Library for heat and work streams respectively. A single Mixer block cannot mix streams of different types (material, heat, work). FSplit combines material streams (or heat streams or work streams) and divides the resulting stream into two or more outlet streams. All outlets have the same composition and properties. Use FSplit to model flow splitters and purges or vents. One must provide specifications for all but one outlet stream. FSplit calculates the flowrate of the unspecified stream. M01 Combine the recycle stream 31 with the alcohol feedstock.

Figure A2. 7

Table A2. 17 Inputs of M01

Pressure 10.5 bar Valid phase V-L Temperature estimate K

Table A2. 18 Results of M01

Outlet Temperature 454K Outlet Pressure 1050000N/sqm Vapor fraction 0



1st/total liquid 1 M02

Combine the air feed flow with the recycled vapor flow in which contains much less concentration of oxygen.

Figure A2. 8


Pressure 1atm Valid phase V-L Temperature estimate 298K


outlet Temperature 298K outlet Pressure 101325N/sqm vapor fraction 1 1st/total liquid 1 M03 The stream 12 is split in two ways, one is purged out as the gas waste and another go back to the Mixer M02.

Figure A2. 9




Split fraction (stream 13) 0.16


13 0.16 114 0.84 2

M04 Combine the recycled catalyst rich stream 20 with the raw catalyst stream 5 into stream 8

Figure A2. 10


Pressure 1bar Valid phase V-L Temperature estimate K


Outlet Temperature 429.5K Outlet Pressure 100000N/sqm Vapor fraction 0 1st/total liquid 1



M05 Combine the stream 25 (from the flash drum S02) and stream 24 (from the distillation column S04) into stream 26

Figure A2. 11


Pressure 1bar Valid phase V-L Temperature estimate K


Outlet Temperature 428K Outlet Pressure 100000N/sqm Vapor fraction 0.99 1st/total liquid 1 M06 Split the stream 19 into stream 20 (90% m/m) and stream 21 for the further separation

Figure A2. 12




Split fraction(stream 21) 0.1


21 0.1 120 0.9 2

M07 The M07 performs like a compressor. It increases the pressure of stream 23 in order to avoid the stream 25 (high pressure) flow back.

Figure A2. 13


Pressure 1bar Valid phase V-L Temperture estimate K

Table A2. 30 Results M07

Outlet Temperature 347K Outlet Pressure 100000N/sqm Vapor fraction 0 1st/total liquid 1



A2.4.5 Pump Pump simulates a pump or hydraulic turbine. This model calculates either the power requirement or the power produced, given an outlet pressure specification. Pump can calculate the outlet pressure, given a power specification.

Pump is designed to handle a single liquid phase. For special cases, one can specify two- or three-phase calculations to determine the outlet stream conditions and to compute the fluid density used in the pump equations. The accuracy of the results depends on a number of factors, such as the relative amounts of the phases present, the compressibility of the fluid, and the efficiency specified. Use Pump to change pressure when the power requirement is needed or known.

Use the Pump block to rate a pump or a turbine by specifying scalar parameters or by specifying the related performance curves. To use the performance curves, one can specify either: Dimensional curves such as head versus flow or power versus flow Dimensionless curves such as head coefficient versus flow coefficient

P01 Increase the pressure of feedstock (stream 1) to 10.5 bar

Figure A2. 14

Table A2. 31 Inputs

Model pump Discharging pressure 10.5 bar

Table A2. 32 Results

fluid power 32.14 watt brake power 108.42 watt electricity 108.42 watt volume flow rate 3.37E-05 cum/sec pressure change 950000 N/sqm NPSH availabe 92.77 J/kg NPSH required Head developed 887.71 J/kg Pump efficiency used 0.296



Net work required 108.42 Watt P02 Increase the pressure of the TEMPO feed (stream 5) to 10.5 bar

Figure A2. 15

Table A2. 33 Inputs of P02

model pump discharging pressure 10.5 bar

Table A2. 34 Results of P02

fluid power 55.5 watt brake power 187.93 watt electricity 187.93 watt volume flow rate 5.85E-05 cum/sec pressure change 950000 N/sqm NPSH availabe 5.84 J/kg NPSH required Head developed 987.4 J/kg Pump efficiency used 0.296 Net work required 187.93 Watt P03 Increase the pressure of recycle (stream 30)

Figure A2. 16






fluid power 60.04 watt brake power 203.08 watt electricity 203.08 watt volume flow rate 1.50E-04 cum/sec pressure change 400000 N/sqm NPSH availabe 0 J/kg NPSH required Head developed 508.74 J/kg Pump efficiency used 0.296 Net work required 203.08 Watt P04 Decrease the pressure of stream 27 to stream 28 which is favorable to the distillation column S03

Figure A2. 17






fluid power 86.92 watt brake power 294.00 watt electricity 294.00 watt volume flow rate 1.58E-04 cum/sec pressure change 550000 N/sqm NPSH availabe -106.34 J/kg NPSH required Head developed 570.54 J/kg Pump efficiency used 0.296 Net work required 294.00 Watt



A2.4.6 Reactor RStoic models a reactor when: Reaction kinetics are unknown or unimportant. Stoichiometry is known. You can specify the extent of reaction or conversion.

Rstoic can handle reactions that occur independently in a series of reactors. It can also perform product selectivity and heat of reaction calculations.

R01 In the reactor, two parallel reactions take place. The first is the oxidation of 3,3-DIMETHYLBUTANOL to 3,3-DIMETHYLBUTYRALDEHYDE. The other is the oxidizing reaction of 1-HEXENE to 1-HEXANAL. The condition is 10bar/ 373K and the conversion of alcohol is 23%. Stream 6 is the alcohol feed, Stream 37 is the oxygen feed and the stream 9 is the catalyst feed. The Effluent of the reactor is stream 10.

Figure A2. 18

Table A2. 39 Inputs of R01

T 373 K P 10 bar Valid phase V-L

Rxn No Specification type Stoichiometry 2 CONVERSION 2 E + O2 --> 2 A + 2 WATER 1 CONVERSION 2 HEXENE + O2 --> 2 HEXANAL



Table A2. 40 Results of R01

Outlet Temperature 373.15 K Outlet Pressure 1000000 N/sqm Heat duty -85317.73 Watt Net duty 0 Watt 1st/total liquid 1 Vapor frac 0.769

Rxn No Reaction extent

(kmol/sec) Stoichiometry 1 6.48E-08 2 HEXENE + O2 --> 2 HEXANAL2 0.000172 2 E + O2 --> 2 WATER + 2 A



A2.4.7 Separator Flash2 performs rigorous 2 (vapor-liquid) or 3 (vapor-liquid-liquid) phase equilibrium calculations. Flash2 produces one vapor outlet stream, one liquid outlet stream, and an optional water decant stream. One can use Flash2 to model flashes, evaporators, knock-out drums, and any other single-stage separators, with sufficient vapor disengagement space. Optionally, one can specify a percentage of the liquid phase to be entrained in the vapor stream. Flash2 performs vapor-liquid or vapor-liquid-liquid equilibrium calculations. When one specify the outlet conditions, Flash2 determines the thermal and phase conditions of a mixture of one or more inlet streams. DSTWU performs a Winn-Underwood-Gilliland shortcut design calculation for a single-feed, two-product distillation column, with a partial or total condenser. For the specified recovery of the light and heavy key components, DSTWU estimates the minimum for either: Reflux ratio Number of theoretical stages

DSTWU estimates one of the following requirements: Reflux ratio given the number of theoretical stages Number of theoretical stages given the reflux ratio

DSTWU also estimates: Optimum feed stage location Condenser and reboiler duties

DSTWU assumes constant molal overflow and constant relative volatilities. DSTWU uses this method/correlation To estimate Winn Minimum number of stages Underwood Minimum reflux ratio

Gilliland

Required reflux ratio for a specified number of stages or the required number of stages for a specified reflux ratio

DSTWU also estimates the optimum feed stage location and the condenser and reboiler duties. DSTWU can produce tables and plots of reflux ratio versus number of stages.



RadFrac is a rigorous model for simulating all types of multistage vapor-liquid fractionation operations. In addition to ordinary distillation, it can simulate: Absorption Reboiled absorption Stripping Reboiled stripping Extractive and azeotropic distillation

RadFrac is suitable for: Three-phase systems Narrow-boiling and wide-boiling systems Systems exhibiting strong liquid phase nonideality

RadFrac can detect and handle a free-water phase or other second liquid phase anywhere in the column. One can decant free water from the condenser.

RadFrac can model columns where chemical reactions are occurring. Reactions can have fixed conversions, or they can be: Equilibrium Rate-controlled Electrolytic

RadFrac is a rigorous model for simulating all types of multistage vapor-liquid fractionation operations. These operations include: Ordinary distillation Absorption Reboiled absorption Stripping Reboiled stripping Extractive and azeotropic distillation



S01 Flash The flash drum S01 is used to separate the nitrogen and oxygen vapor flow from liquid mixture. Stream 12 is the vapor flow to be recycled to the reactor and the stream 15 is the liquid mixture

Figure A2. 19

Table A2. 41 Inputs of S01


Table A2. 42 Results of S01

Outlet Temperature 298K Outlet Pressure 100000N/sqm Heat duty 3.49E-11Watt Net duty 3.49E-11Watt 1st/total liquid 1 Vapor frac 0.751 S02 Flash This flash drum makes a rough separation of catalyst and product. The reason why a distillation column has not been chose is the catalyst concentration of stream 18 is only about 3%. It’ll cost lots of energy if it’s separated by a distillation column.



Figure A2. 20




Outlet Temperature 430K Outlet Pressure 100000N/sqm Heat duty 7.24E+04Watt Net duty 7.24E+04Watt 1st/total liquid 1 Vapor frac 0.756 S03 Distillation Column This distillation column aims to the product purification. 3,3-DIMETHYLBUTANOL and the product, 3,3-DIMETHYLBUTYRALDEHYDE are separated. The bottoms (stream 30) is mainly alcohol while the distillates (stream 29) is 99% pure aldehyde.

Figure A2. 21




reflux ratio -1.3 condensor P 6.5bar reboiler P 6.5bar condenser total light component A recovL 0.9812 heavy component E recovH 0.0002621


Minimum reflux ratio: 3.014 Actual reflux ratio: 3.918 Minimum number of stages: 61.2 Number of actual stages: 108.9 Feed stage: 71.7 Number of actual stages above feed: 70.7 Reboiler heating required: 117027.26Watt Condenser cooling required: 71291.43Watt Distillate temperature: 298.0K Bottom temperature: 488.5K Distillate to feed fraction: 0.234

Table A2. 47 RADFRAC profile of S03

K P(N/sqm) heat duty Lflow kg/hr Vflow kg/hr

1 301.4 650000 -133639 501.94 0.00 2 479.6 650000 0 1261.47 630.04 3 480.5 650000 0 1255.10 1389.57 4 480.9 650000 0 1244.67 1383.20 5 481.3 650000 0 1234.98 1372.77 6 481.7 650000 0 1226.24 1363.08 7 482.0 650000 0 1218.49 1354.34 8 482.3 650000 0 1211.69 1346.58 9 482.5 650000 0 1205.79 1339.79

10 482.7 650000 0 1200.71 1333.89 11 482.9 650000 0 1196.36 1328.81



12 483.1 650000 0 1192.67 1324.46 13 483.2 650000 0 1189.54 1320.77 14 483.3 650000 0 1186.90 1317.64 15 483.4 650000 0 1184.69 1315.00 16 483.5 650000 0 1182.84 1312.79 17 483.6 650000 0 1181.28 1310.94 18 483.6 650000 0 1179.98 1309.38 19 483.7 650000 0 1178.90 1308.08 20 483.7 650000 0 1178.01 1307.00 21 483.7 650000 0 1177.26 1306.10 22 483.8 650000 0 1176.64 1305.36 23 483.8 650000 0 1176.12 1304.74 24 483.8 650000 0 1175.69 1304.22 25 483.8 650000 0 1175.32 1303.79 26 483.8 650000 0 1175.03 1303.42 27 483.8 650000 0 1174.79 1303.13 28 483.9 650000 0 1174.59 1302.89 29 483.9 650000 0 1174.43 1302.69 30 483.9 650000 0 1174.29 1302.52 31 483.9 650000 0 1174.18 1302.39 32 483.9 650000 0 1174.08 1302.27 33 483.9 650000 0 1173.99 1302.18 34 483.9 650000 0 1173.93 1302.09 35 483.9 650000 0 1173.88 1302.03 36 483.9 650000 0 1173.83 1301.97 37 483.9 650000 0 1173.79 1301.93 38 483.9 650000 0 1173.76 1301.89 39 483.9 650000 0 1173.74 1301.86 40 483.9 650000 0 1173.71 1301.83 41 483.9 650000 0 1173.70 1301.81 42 483.9 650000 0 1173.68 1301.79 43 483.9 650000 0 1173.67 1301.78 44 483.9 650000 0 1173.66 1301.77 45 483.9 650000 0 1173.65 1301.76 46 483.9 650000 0 1173.64 1301.75 47 483.9 650000 0 1173.63 1301.74 48 483.9 650000 0 1173.63 1301.73 49 483.9 650000 0 1173.62 1301.73 50 483.9 650000 0 1173.62 1301.72 51 483.9 650000 0 1173.61 1301.72 52 483.9 650000 0 1173.61 1301.71 53 483.9 650000 0 1173.61 1301.71 54 483.9 650000 0 1173.61 1301.70 55 483.9 650000 0 1173.60 1301.70 56 483.9 650000 0 1173.60 1301.70 57 483.9 650000 0 1173.60 1301.70



58 483.9 650000 0 1173.60 1301.70 59 483.9 650000 0 1173.60 1301.70 60 483.9 650000 0 1173.60 1301.70 61 483.9 650000 0 1173.61 1301.70 62 483.9 650000 0 1173.61 1301.70 63 483.9 650000 0 1173.60 1301.70 64 483.9 650000 0 1173.60 1301.70 65 483.9 650000 0 1173.60 1301.70 66 483.9 650000 0 1173.60 1301.69 67 483.9 650000 0 1173.60 1301.69 68 483.9 650000 0 1173.60 1301.69 69 483.9 650000 0 1173.58 1301.69 70 483.9 650000 0 1173.42 1301.68 71 483.9 650000 0 1170.90 1301.51 72 484.2 650000 0 2400.74 1298.99 73 484.3 650000 0 2401.43 1980.38 74 484.3 650000 0 2401.10 1981.06 75 484.3 650000 0 2400.66 1980.73 76 484.3 650000 0 2400.15 1980.30 77 484.3 650000 0 2399.59 1979.79 78 484.4 650000 0 2398.97 1979.22 79 484.4 650000 0 2398.28 1978.60 80 484.4 650000 0 2397.53 1977.92 81 484.4 650000 0 2396.71 1977.17 82 484.4 650000 0 2395.82 1976.35 83 484.4 650000 0 2394.79 1975.45 84 484.5 650000 0 2393.72 1974.42 85 484.5 650000 0 2392.55 1973.35 86 484.5 650000 0 2391.42 1972.18 87 484.6 650000 0 2390.05 1971.06 88 484.6 650000 0 2388.57 1969.69 89 484.6 650000 0 2386.95 1968.20 90 484.7 650000 0 2385.21 1966.59 91 484.7 650000 0 2383.31 1964.84 92 484.8 650000 0 2381.27 1962.95 93 484.8 650000 0 2379.06 1960.90 94 484.9 650000 0 2376.68 1958.69 95 484.9 650000 0 2374.12 1956.32 96 485.0 650000 0 2371.38 1953.76 97 485.1 650000 0 2368.44 1951.01 98 485.1 650000 0 2365.29 1948.07 99 485.2 650000 0 2361.95 1944.93

100 485.3 650000 0 2358.40 1941.59 101 485.4 650000 0 2354.63 1938.03 102 485.5 650000 0 2350.65 1934.27 103 485.6 650000 0 2346.46 1930.28



104 485.7 650000 0 2342.05 1926.09 105 485.8 650000 0 2337.44 1921.69 106 485.9 650000 0 2332.58 1917.07 107 486.0 650000 0 2326.67 1912.21 108 486.2 650000 0 2305.57 1906.31 109 487.8 650000 178674.3 420.37 1885.20



S04 Distillation Column Some alcohol and aldehyde go out of flash drum S02 along with the catalyst. It is not economical to treat them as the waste. Therefore a further separation is done by distillation column S04. The distillates are the mixture of alcohol and aldehyde and the bottoms are TEMPO and Ru-100.

Figure A2. 22


reflux ratio -1.2 condensor P 0.1bar reboiler P 0.1bar condenser total light component E recovL 0.8227 heavy component TEMPO recovH 0.02642


Minimum reflux ratio: 0.0091 Actual reflux ratio: 0.0109 Minimum number of stages: 1.5 Number of actual stages: 59.6 Feed stage: 44.6 Number of actual stages above feed: 43.6 Reboiler heating required: 502.60Watt Condenser cooling required: 1042.91Watt Distillate temperature: 347.0K



Bottom temperature: 392.6K Distillate to feed fraction: 0.479

Table A2. 50 RadFrac Profile of S04

K P(N/sqm) heat duty Lflow kg/hr Vflow kg/hr

1 348.0 10000 -1392.94 0.111 0.000 2 373.9 10000 0.00 0.083 10.285 3 375.1 10000 0.00 0.083 10.257 4 375.1 10000 0.00 0.083 10.256 5 375.1 10000 0.00 0.083 10.256 6 375.1 10000 0.00 0.083 10.256 7 375.1 10000 0.00 0.083 10.256 8 375.1 10000 0.00 0.083 10.256 9 375.1 10000 0.00 0.083 10.256

10 375.1 10000 0.00 0.083 10.256 11 375.1 10000 0.00 0.083 10.256 12 375.1 10000 0.00 0.083 10.256 13 375.1 10000 0.00 0.083 10.256 14 375.1 10000 0.00 0.083 10.256 15 375.1 10000 0.00 0.083 10.256 16 375.1 10000 0.00 0.083 10.256 17 375.1 10000 0.00 0.083 10.256 18 375.1 10000 0.00 0.083 10.256 19 375.1 10000 0.00 0.083 10.256 20 375.1 10000 0.00 0.083 10.256 21 375.1 10000 0.00 0.083 10.256 22 375.1 10000 0.00 0.083 10.256 23 375.1 10000 0.00 0.083 10.256 24 375.1 10000 0.00 0.083 10.256 25 375.1 10000 0.00 0.083 10.256 26 375.1 10000 0.00 0.083 10.256 27 375.1 10000 0.00 0.083 10.256 28 375.1 10000 0.00 0.083 10.256 29 375.1 10000 0.00 0.083 10.256 30 375.1 10000 0.00 0.083 10.256 31 375.1 10000 0.00 0.083 10.256 32 375.1 10000 0.00 0.083 10.256 33 375.1 10000 0.00 0.083 10.256 34 375.1 10000 0.00 0.083 10.256 35 375.1 10000 0.00 0.083 10.256 36 375.1 10000 0.00 0.083 10.256



37 375.1 10000 0.00 0.083 10.256 38 375.1 10000 0.00 0.083 10.256 39 375.1 10000 0.00 0.083 10.256 40 375.1 10000 0.00 0.083 10.256 41 375.1 10000 0.00 0.083 10.256 42 375.1 10000 0.00 0.083 10.256 43 375.1 10000 0.00 0.083 10.256 44 375.1 10000 0.00 0.099 10.256 45 378.8 10000 0.00 17.464 6.375 46 410.6 10000 0.00 18.696 6.403 47 435.3 10000 0.00 19.194 7.635 48 440.2 10000 0.00 19.268 8.132 49 440.7 10000 0.00 19.276 8.206 50 440.8 10000 0.00 19.277 8.215 51 440.8 10000 0.00 19.277 8.216 52 440.8 10000 0.00 19.277 8.216 53 440.8 10000 0.00 19.277 8.216 54 440.8 10000 0.00 19.277 8.216 55 440.8 10000 0.00 19.277 8.216 56 440.8 10000 0.00 19.277 8.216 57 440.8 10000 0.00 19.277 8.216 58 440.8 10000 0.00 19.277 8.215 59 440.8 10000 0.00 19.277 8.215 60 440.8 10000 841.57 11.061 8.215



Appendix 4.1 Process Stream Summary

Table A4. 1 Process Stream Summary (Stream 1 Stream 4)


Stream Nr. <1> IN

Alcohol Feed <2>

Pump P01outlet <3> IN

Air Feed <4>

Compressor C01 Feed

Component Flow Flow Flow Flow

kg/h kmol/h kg/h kmol/h kg/h kmol/h kg/h kmol/h

HEXANE 0.390 4.53 0.390 4.53 0.000 0.00 0.187 2.17

HEXENE 0.260 3.09 0.260 3.09 0.000 0.00 0.127 1.51

HEXANAL 0.000 0.00 0.000 0.00 0.000 0.00 0.088 0.88

WATER 0.000 0.00 0.000 0.00 0.000 0.00 1.753 97.32

TEMPO 0.000 0.00 0.000 0.00 0.000 0.00 0.025 0.16

3,3-dimethylbutanol 129.350 1265.87 129.350 1265.87 0.000 0.00 7.035 68.85 3,3-dimethylbutyraldehyde 0.000 0.00 0.000 0.00 0.000 0.00 6.539 65.28

O2 0.000 0.00 0.000 0.00 27.300 853.13 66.483 2077.67

N2 0.000 0.00 0.000 0.00 102.700 3666.06 640.746 22872.78

Ru-100 0.000 0.00 0.000 0.00 0.000 0.00 0.000 0.00

Total 130.000 1273.48 130.000 1273.48 130.000 4519.19 722.984 25186.62

Enthalpy : kW -135.964 -135.855 -0.005 -18.229 Temperature : K 298 300 298 298

Pressure : N/sqm 100000 1050000 100000 101325

Phase L L V V

Stream Nr. <5>

TEMPO Feed <6>

Reactor Inlet <7>

Compressor C01 outlet<8>

Pump P02 inlet



HEXANE 0.000 0.00 0.390 4.53 0.187 2.17 0.012 0.14

HEXENE 0.000 0.00 0.260 3.09 0.127 1.51 0.006 0.07

HEXANAL 0.000 0.00 2.094 20.91 0.088 0.88 0.285 2.84

WATER 0.000 0.00 0.000 0.00 1.753 97.32 0.001 0.05

TEMPO 11.400 72.96 25.123 160.79 0.025 0.16 116.740 747.11


O2 0.000 0.00 0.000 0.00 66.483 2077.67 0.000 0.00

N2 0.000 0.00 0.000 0.00 640.746 22872.78 0.000 0.00

Ru-100 24.140 25.17 0.000 0.00 0.000 0.00 241.400 251.70

Total 35.540 98.13 554.882 5347.45 722.984 25186.62 443.913 1841.27

Enthalpy : kW -1.228 -500.609 71.951 -92.291 Temperature : K 298 454 714 429

Pressure : N/sqm 100000 1050000 1050000 100000

Phase L L V L





Stream Nr. <9>

Pump P02 outlet <10>

Reactor Effluent <11>

Flash S01 inlet <12>

Tops from S01



HEXANE 0.012 0.14 0.589 6.84 0.589 6.84 0.223 2.59 HEXENE 0.006 0.07 0.353 4.20 0.353 4.20 0.151 1.79 HEXANAL 0.285 2.84 2.513 25.09 2.513 25.09 0.105 1.05 WATER 0.001 0.05 24.059 1335.49 24.059 1335.49 2.087 115.86 TEMPO 116.740 747.11 141.888 908.06 141.888 908.06 0.029 0.19 3,3-dimethylbutanol 70.741 692.32 475.929 4657.72 475.929 4657.72 8.375 81.97 3,3-dimethylbutyraldehyde 14.728 147.05 147.629 1473.91 147.629 1473.91 7.784 77.72 O2 0.000 0.00 46.667 1458.36 46.667 1458.36 46.647 1457.75 N2 0.000 0.00 640.747 22872.78 640.747 22872.78 640.532 22865.11 Ru-100 241.400 251.70 241.400 251.70 241.400 251.70 0.000 0.00

Total 443.913 1841.27 1721.773 32994.15 1721.773 32994.15 705.933 24604.02


Pressure : N/sqm 1050000 1000000 100000 100000 Phase L V-L V-L V

Stream Nr. <13> OUT

Discharge <14>

Nitrogen recycle <15>

Feed to Decanter D01 <16> OUT

Discharge



HEXANE 0.036 0.41 0.187 2.17 0.366 4.25 0.000 0.00

HEXENE 0.024 0.29 0.127 1.51 0.203 2.41 0.000 0.00

HEXANAL 0.017 0.17 0.088 0.88 2.408 24.05 0.000 0.00

WATER 0.334 18.54 1.753 97.32 21.972 1219.64 21.955 1218.67

TEMPO 0.005 0.03 0.025 0.16 141.858 907.88 0.000 0.00


O2 7.463 233.24 39.183 1224.50 0.020 0.61 0.000 0.00

N2 102.485 3658.39 538.046 19206.68 0.215 7.67 0.000 0.00

Ru-100 0.000 0.00 0.000 0.00 241.400 251.70 0.000 0.00

Total 112.949 3936.62 592.984 20667.36 1015.840 8390.19 21.955 1218.67


Pressure : N/sqm 100000 100000 100000 100000

Phase V V L L





Stream Nr. <17>

Decanter D01 outlet <18>

Feed to Flash S02 <19>

Bottoms of Flash S02 <20>

Catalyst recycle




Total 993.885 7171.51 993.885 7171.51 453.747 1936.79 408.373 1743.16


Pressure : N/sqm 100000 100000 100000 100000 Phase L V-L L L

Stream Nr. <21>

Feed to dist. Col S04 <22>

Bottoms of S04 <23>

Distillates of S04 <24>

Pressurized flow



HEXANE 0.001 0.02 0.000 0.00 0.001 0.02 0.001 0.02

HEXENE 0.001 0.01 0.000 0.00 0.001 0.01 0.001 0.01

HEXANAL 0.032 0.32 0.003 0.03 0.029 0.29 0.029 0.29

WATER 0.000 0.01 0.000 0.00 0.000 0.01 0.000 0.01

TEMPO 11.704 74.91 11.395 72.93 0.309 1.98 0.309 1.98


O2 0.000 0.00 0.000 0.00 0.000 0.00 0.000 0.00

N2 0.000 0.00 0.000 0.00 0.000 0.00 0.000 0.00

Ru-100 24.140 25.17 24.140 25.17 0.000 0.00 0.000 0.00

Total 45.375 193.69 37.050 112.94 8.325 80.74 8.325 80.74


Pressure : N/sqm 100000 10000 10000 100000

Phase L L L L





Stream Nr. <25>

Tops of the Flash S02 <26>

Feed to the Cooler H03<27>

Feed to Pump P04 <28>

Feed to column S03




Total 540.138 5234.68 548.463 5315.42 548.463 5315.42 548.463 5315.42


Pressure : N/sqm 100000 100000 100000 650000 Phase V V-L L L

Stream Nr. <29> OUT

Distillate of S03 <30>

Bottoms of S03 <31>

Alcohol Recycle <32>

Bottoms from S04



HEXANE 0.354 4.11 0.000 0.00 0.000 0.00 0.187 2.17

HEXENE 0.197 2.34 0.000 0.00 0.000 0.00 0.127 1.51

HEXANAL 0.027 0.27 2.094 20.91 2.094 20.91 0.088 0.88

WATER 0.017 0.92 0.000 0.00 0.000 0.00 1.753 97.32

TEMPO 0.000 0.00 25.123 160.79 25.123 160.79 0.025 0.16


O2 0.020 0.61 0.000 0.00 0.000 0.00 66.483 2077.67

N2 0.215 7.67 0.000 0.00 0.000 0.00 640.746 22872.78

Ru-100 0.000 0.00 0.000 0.00 0.000 0.00 0.000 0.00

Total 123.581 1241.44 424.882 4073.93 424.882 4073.93 722.984 25186.62


Pressure : N/sqm 650000 650000 1050000 1050000

Phase L L L V-L




Stream Nr. <33> <35> <36> <37>




Total 21.772 25.17 11.400 72.96 24.140 25.17 25.28 32.47

Enthalpy : kW -1.012 -0.520 -0.708 -1.686

Temperature : K 298 298 298 298

Pressure : N/sqm 100000 100000 100000 100000

Phase L L S S-L

Conceptual Process DesignDesign of a plant utilizing novel catalytic method for alcohol oxidation

AppendixCPD-3284

A4.2 Utilities summary

EQUIPMENTHeating Cooling Power

Nr. Name Load Consumption (t/a) Load Consumption (t/a) Actual Consumption (t/a, kWh/a)Steam Hot Cooling Air Refrig. Load Steam (t/h) Electricity Remarks

kW LP MP HP Oil kW Water kW HP MP kWh/aT05 Stirrer of dissolving tank 0 0 0,3 2592R01 Stirrer of reactor 0 0 1,44 12441,6 refer to appendix A3.2 for the summary of utilitiesR01 Reactor 0 21142080 0 refer to calculations in Appendix A5.2C01 compressor 0 0 56,3 486432 refer to calculations in Appendix A5.3S05 evaporator 103680 0 0H01 Heat exchanger 0 20217600 0 refer to calculations in appendix A5.8H02 Heat exchanger 552960 0 0H03 Heat exchanger 0 9953280 0H04 Heat exchanger 0 7361280 0H11 Condenser 0 1,51E+08 0 refer to calculation in Appendix A5.8H12 Reboiler 1658880 0 0H13 Condenser 0 1866240 0H14 Reboiler 13824 0 0P01 Pump 0 0 0,32 2764,8 refer to Appendix A5.9 of pump design P02 Pump 0 0 0,8 6912P03 Pump 0 0 0,45 3888P04 Pump 0 0 0,76 6566,4P05 Pump (into dissolving tank) 0 0 0,1 864P06 Pump(out of dissolving tank) 0 0 0,1 864P07 Pump (in the evaporator) 0 0 1,1 9504P08 Vaccum Pump 0 0 1 8640

TOTAL 2329344 2.12·10862,67 541468,8

Project ID Number : CPD 3284Completion date : 14 Jan 2002

SUMMARY OF UTILITIESUTILITIES

Lin Luo (1113259)Shuang Zhao (1118447) Appendix 4-6

Weimin Wang (1118668)Zhengjie Zhu (1113658)

Conceptual Process DesignDesign of a plant utilizing novel catalytic method for alcohol oxidation

AppendixCPD-3284

Lin Luo (1113259)Shuang Zhao (1118447) Appendix 4-6

Weimin Wang (1118668)Zhengjie Zhu (1113658)



Appendix 5

Equipment Design A5.1 Design of storage tanks A5.1.1 3,3-dimethylbutanol storage tank T01 & product storage tank T04 Assumption: - each of the two storage tanks can store the amount 10 days. Because

the amount required for 10 days of raw material is a little larger than that of product. So, the size of the two tanks is designed to same according to the amount of raw material

The amount of 3,3-dimethylbutanol required for 10 days: 130kg/h 10 day 24 hour = 3120 kg (chapter 5) The density of 3,3-dimethylbutanol is 0.817 103 kg/m3 (Please refer to part 3.2.4, chapter 3) The total volume of 3,3-dimethylbutanol required for 10 days 3120/(0.817 103) = 38.19 m3 It is designed that the liquid only occupies 80% of the volume of the tank (for safety reason), so the total design volume is 47.73 m3. The diameter of the tank is chosen to be 3.8m and by using the following

equation 2

4V D H

, the height of tank is calculated to be 4.2m.

Table A5. 1 Calculation of tank dimension (T01)

V= π/4 D2H (m3)

D(m)

H(m)

47.73 3.2 5.947.73 3.4 5.247.73 3.6 4.647.73 3.8 4.247.73 4.0 3.847.73 4.2 3.4



A5.1.2 Co-catalyst (TEMPO) storage tank T02 & By-product (mainly deactivated TEMPO) tank T04 Assumption: - Sizes of storage tank is designed according to requirement of By-

product, because amount of by-product is slight large than that of TEMPO. And each of them can store amount required for 30 days

The amount of TEMPO required for 30 days: 11.4 kg/h 30 day 24 hour = 8208 kg (chapter 5) The density of TEMPO is estimated 0.8 103 kg/m3 The total volume of TEMPO required for 30 days 8208/(0.8 103) = 10.26 m3 It is designed that the liquid only occupies 80% of the volume of the tank (for safety reason), so the total design volume is 12.82m3. The diameter of the tank is chosen to be 2.4m and by using the following

equation 2

4V D H


Table A5. 2 Calculation of tank dimension (T04)

V= π/4 D2H (m3)

D(m)

H(m)

12.9 2.0 4.1112.9 2.2 3.3912.9 2.4 2.8512.9 2.6 2.4312.9 2.8 2.10



A5.2 Design of dissolving tank T05 A5.2.1 Dimensioning of the dissolving tank Assumption: this tank is designed to amount required to 1 day (24 hour). The volume of mixture increases 60 percent since the solid dissolved into liquid.

Table A5. 3 The volume of the components in dissolving tank T02

Components

Density (kg/l)

Weight (kg)

Volume (m3)

TEMPO 0.8 273.6 0.342Ru-catalyst 559.7

Total 833.3 0.547

The volume of the mixture is 0.547m3. It is designed that the liquid only occupies 80% of the volume of the tank (for safety reason), so the total design volume of the tank is 0.684 m3. The diameter of the tank is chosen to be 0.8m and by using the following equation

2

4V D H


Figure A5. 1 Configuration of the dissolving tank



A5.2.2 Design of the stirrer and mixing time calculation

2 21/3

2 /3 4 /3 1/3m s

mix

t L HN

D H D

(1) [29]

Where: Nmix Mixing number tm Mixing time [s] Total specific power input [W/kg] D Diameter of dissolve tank [m] Homogeneity factor Using a homogeneity of 95% and maximum distance

=0.374

Primary eddy size sH

D ( Hs is blade height)

Average viscosity index 35 Ls Flow path length [m] H Reactor height [m] LS = H + 2D = 0.9D + 2D = 2.9D

22.39.0

9.2

D

D

H

LS

4 3 53 5

2 2

1' '2

4 4

'

I LS I P

LL

S

N c N DP N N N D

D DVH H

Hc

D

Where: Ps Power input per stirrer [W]

L Density for the liquid [kg/m3] V Reactor volume [m3] D’ Stirrer diameter [m] D’=1/3D assume it is Rushton stirrer. D Reactor diameter [m] c Ratio Np Power number 6 (Rushton) N Angle speed

44 4

1 2 2 60.123

2P

P

Nc N c

' 0.123 ' ' 8S SH cD D D H 1

0.04163 ' 3 8 24

S S S

S

H H H

D D H

5.669.022.3

350416.0

374.0 22

3/13/4

2

3/13/4

D

H

H

LN S

mix

(2) [29] (3)



From equation (1):1/3 2 /3

2 /3 1/3m mix

mix m

t N DN t

D

(4)

From calculation in A5.2.1: V = 0.684m3 and D = 0.8 m, hence:

3/1

3/2

3/1

3/2 8.05.66

DN

t mixm

for suspending light solids, blending high-viscosity liquid, the input power of stirred tank is chosen for 0.4kW/m3 [31]. Therefore, for the dissolving tank T05 (V = 0.684 m3), Ps = 0.4kW/m3 0.684m3 = 0.3 kW

Total specific power is can be determined from the formula:

sPower P W

Mass kg ,

which is 273.6W/833.3kg = 0.3283W/kg. Using equation (4) the mixing time is calculated to be 96s.



A5.2.3 Design of attached pump P06 The same pump as pump P05 is chosen, because they have the analogous working- condition. (Volumetric flow rate, head)



A5.3 Design of reactor R01 A5.3.1 Dimensioning of the reactor

Table A5. 4 The volume of the liquid in reactor R01

Component

Input (kg/h)

Density (kg/l)

Volume (m3)

Remarks

Alcohol 602.7 0.817 0.7377(1) Refer to part 5.3 chapter 5 for the input flow rate

(2) Refer to part 3.2.4, chapter 3 for the density of the components

(3) Volume of Ru-catalyst is assumed the volume increased by dissolution. The value is 60% volume of TEMPO.

(4) Gas hold-up volume is assumed to account for 10% total liquid volume

(stream 6 +7 +9) TEMPO 142.1 0.800 0.1776 (stream 6 +7 +9) Dimethybutyraldehe 23.6 0.798 0.0296 Ru-catalyst 233.2 1.541 0.1066* Total liquid 1001.6 1.0515Gas hold-up 0.1052

Total 1.1567

The volume of the mixture is 1.1567m3. It is designed that the liquid only occupies 80% of the volume of the reactor (for safety reason), so the total design volume of the reactor is 1.44 m3. Diameter and height chosen: The following formulas are used to calculate the diameter, height and heat exchange area of the reactor:

2 2 4 4liquid liquidV D H V D H

A DH

4

4

V D VD

A A useful liquidA DH

Table A5. 5 Comparison of the heat exchange area for the decision on the dimensions of the reactor

Diameter (m)

Height (m)

Heat exchange area (m2)

Liquid height (m)

Useful heat exchange area(m2)

0.6 5.11 9.64 3.72 7.010.8 2.88 7.23 2.09 5.260.9 2.27 6.43 1.65 4.671.0 1.84 5.78 1.34 4.211.1 1.52 5.26 1.11 3.821.2 1.28 4.82 0.93 3.51



In order to mix perfectly, the diameter for normal reactor is always less than 1 m. As long as heat exchange area is enough to control the temperature of reactor, the smallest exchange area is chosen because it leads to lower cost for construction of reactor. The minimum heat exchange area calculates: General equation for heat transfer across a surface is:

Q = UATm

where: Q: heat transferred per unit time [kW] U: the overall heat transfer coefficient [kW/m2 oC]

A: heat-transfer area [m2] Tm: the main temperature difference, the temperature driving force: From the ASPEN simulation, duty of reactor is -85.32 kw Cooling water is used, which design value of inlet temperature is 20 0C, maximum allowed outlet temperature is 400C. Average of temperature is 300C For jacket water as cooling media, U is 0.23 to 0.3 according to process design principle. In this case, 0.3 [kW/m2 oC] is chosen. The temperature of reactor is controlled to 1000C.

The minimum area 06.4301003.0

23.85

mTU

QA m2

The flow rate of cooing water Mw = 85.23/(4.1830) = 0.68 kg/s = 2447 kg/h Therefore the dimension of the reactor is: diameter D = 1m, height H = 1.84m. Useful heat exchange area is 4.21 m2



A5.3.2 Design of the stirrer and mixing Dimension of stirrer: Stirrer diameter is calculated from the formula D’ = D/3,[31] in which D is reactor diameter, hence D’ = 0.333m.

44 4

1 2 2 60.123

2P

P

Nc N c

[29] (1)

' 0.123 ' ' 8S SH cD D D H [29] (2)

c Ratio Np Power number 6 (Rushton) Stirrer height is calculated from equation (2), Hs = 0.04m Power input: Because the low power input (blending low-viscosity liquid at 1000C, extra agitation by gas flow) for stirred tank is 1-2kW/m3 according to process design principle.

Therefore, for the reactor R101 (V = 1.44m3) the power needs to be input : Ps = 1kW/m3 1.44m3 = 1.44kW Mixing time:

2 21/3

2 /3 4 /3 1/3m s

mix

t L HN

D H D

[29] (3)

Nmix Mixing Number tm Mixing time [s] Total specific power input [W/kg] D Diameter of reactor [m] Homogeneity factor Using a homogeneity of 95% and maximum distance

=0.374

Primary eddy size sH

D ( Hs is blade height)

Average viscosity index 35.013 Ls Flow path length [m] H Reactor height [m] LS = H + 2D = 1.84D + 2D = 3.84D

09.284.1

84.3

D

D

H

LS



4 3 53 5

2 2

1' '2

4 4

'

I LS I P

LL

S

N c N DP N N N D

D DVH H

Hc

D

Ps Power input per stirrer [W]

L Density for the liquid [kg/m3] V Reactor volume [m3] D’ Stirrer diameter [m] D’=1/3D assume it is Rushton stirrer. D Reactor diameter [m] N Angle speed

10.0416

3 ' 3 8 24S S S

S

H H H

D D H

3.11784.109.2

013.350416.0

374.0 22

3/13/4

2

3/13/4

D

H

H

LN S

mix

From equation (1):1/3 2 /3

2 /3 1/3m mix

mix m

t N DN t

D

(6)

From calculation in A5.3.1: V = 1.44m3 and D = 1m, hence:

3/1

3/2

3/1

3/2 13.117

DN

t mixm

Total specific power can be determined from the formula:

sPower P W

Mass kg ,

which is 1.44KW/1001.6kg = 1.44W/kg (refer to part 5.3, chapter 5 for the mass in the reactor) Using equation (6) the mixing time is calculated to be 104 seconds. Ideal CSTR (continuous stirred tank reactor) behavior is approached when the mean residence time is 5-10 times the length of time needed to achieve homogeneity according to process design principle. In this design, the ratio of residence time to mixing time is 35. It is much larger than the desire. So the assumption that mixing perfectly is reasonable.

(4) (5)



A5.4 Gas-liquid separator & Flash drum A5.4.1 Gas-liquid separator The diameter of the vessel must be large enough to slow the gas down to below the velocity at which the liquid will settle out. So the minimum allowable diameter will be given by[8]:

2/1/07.0

4

vvLt

t

vv

u

u

VD

Where Dv = minimum vessel diameter, m

ut = the settling velocity of the liquid droplets, m/s Vv = gas or vapor volumetric flow-rate, m3/s ρL = liquid density, kg/m3 ρv = gas density, kg/m3

0.07 = coefficient for vessel which has demister. Calculated value of Dv is 0.33 m. and the design diameter D of 0.6 m is chosen. The height of the height of the vessel out let above the inlet should be sufficient to allow for disengagement the liquid drops. Typical value of height H: H = 1.5 Dv + 0.4 + HL

Where HL, the liquid level will depend on the hold-up time necessary for smooth operation and control; typically 10 minutes would be allowed.

2

4

16

1

D

VH L

where 1/6V = volume of 10 minutes residence time, m3 D = design diameter, m Calculated value of H is 1.5 m



A5.4.2 Flash drum

2/1/07.015.0

4

vvLt

t

vv

u

u

VD

Where Dv--- minimum vessel diameter, m

ut --- the settling velocity of the liquid droplets, m/s Vv--- gas or vapor volumetric flow-rate, m3/s ρL--- liquid density, kg/m3 ρv--- gas density, kg/m3

0.07 * 0.15 = coefficient for vessel without a demister. Calculated value of Dv is 0.85 m. and the design diameter of 1 m is chosen. Typical value of height H: H = 1.5 Dv + 0.4 + HL

Where HL, the liquid level will depend on the hold-up time necessary for smooth operation and control; typically 10 minutes would be allowed. Calculated value of H is 2.2 m



A5.5 Design of decanter A rough estimate of the decanter volume required can be made by taking a hold-up time of 5 to 10 minutes, which is usually sufficient where emulsions are not likely to form. The decanter vessel is sized on the basis that the velocity of the continuous phase must be less than settling velocity of the droplets of the dispersed phase. Plug flow is assumed, and the velocity of the continuous phase calculated using the area of the interface. And Strokes’s law is used to determine the settling velocity of the droplets.

C

Cddd

di

CC

gdu

uA

Lu

18

2

where dd = droplet diameter, m, ud , uC = velocity of the dispersed phase and continuous phase, respectively, m/s ρd , ρC = density of the dispersed and continuous phase, respectively, kg/m3 µC = viscosity of continuous phase, N s/m2 g = gravitational aceeleration, 9.81 m/s2 LC = continuous phase volumetric flow rate, m3/s Ai = area of the interface, m2 Design a decanter to separate a light product from water.

Table A5. 6 Component in the decanter

phase flow rate,

kg/h viscosity,

cP density, kg/m3

water dispersed 21.95 1 1000 light product continuous 985.70 3 870

Take dd = 150 10-6 m, as the flow rate is small, use a vertical, cylindrical vessel. Table A5. 7 Data in cylindrical vessel

ud m/s 0.0005317

LC m3/s 0.0003147

Ai m2 0.60

D,m 0.75 Take the height as twice the diameter, a reasonable value for a cylinder: H = 1.5 m Take the dispersion band as 10 percent of the height, 0.15m Check the residence time of the droplets in the dispersion band = 0.15/ 0.0005317 = 282 s = 4.6 min This satisfactory, a time of 2 to 5 minutes is normally recommended.



A5.6 Evaporator design A5.6.1 Vessel for evaporation Operating condition of the vessel is 162 0C, 0.13 bar From the stream 22, there are two components entering evaporator: Tempo,11.4 kg/h And Ru-Catalyst, 23.32 kg/h. assume that 90% of tempo is vaporized and the rest goes out from the stream of solid Ru-catalyst. Solid with tempo rate = 23.32 + 11.410% = 24.46 kg/h Vapor rate = 36.23 – 24.46 = 11.77 kg/h The recirculation rate will need to be 100 times the evaporation rate. Recirculation rate = 100 11.77 = 1177 kg/h

2/1/07.015.0

4

vvLt

t

vv

u

u

VD

Where Dv--- minimum vessel diameter, m

ut --- the settling velocity of the liquid droplets, m/s Vv--- gas or vapor volumetric flow-rate, m3/s ρL--- liquid density, kg/m3 ρv--- gas density, kg/m3

Calculated value of Dv is 0.2 m. and the design diameter of 0.6 m is chosen. Typical value of height H: H = 1.5 Dv + 0.4 + HL

Where HL, the liquid level will depend on the hold-up time necessary for smooth operation and control; typically 10 minutes would be allowed. Calculated value of H is 1.4 m

TEMPO vapour, 11.4kg/h *90%

Tempo + Ru-Catalyst

Ru-Catalyst(s) 23.32kg/h + 1.14kg/h Tempo



A5.6.2 Area of Heat-exchanger The heat input to the evaporator is absorbed by the liquid[34]: q = Mz Cp (T2 - T1) where the subscript z refers to the recycling liquid, and T1 and T2 are the liquid inlet and exit temperatures, respectively. To limit the temperature rise through the heat exchanger to 10 K, the recirculation rate will need to be 100 times the evaporation rate. Evaporate rate = 11.77 kg/h (refer to mass balance) Recirculation rate = 100 11.77 = 1177 kg/h Overall heat-transfer coefficient is assumed 1000 w/m2K T1 = 436 K and T2 = 446K Boiling point of TEMPO (0.13bar) is (436K), Cp = 350.335 J/mol K The recirculating solution is at the final product concentration of 50% solids, Mz = 11.4 100 = 1177 kg/h q = 1177 350.335/156.23 10 = 26393.4 kJ/h = 7.3 kw steam needed = 26393/2200 = 12 kg/h Assuming the temperature difference is 30K at the product inlet, and 20 K at outlet, so the log-mean temperature difference is 24.7 k A = 7.3/(24.7*1) = 0.3 m2



A5.6.3 Vacuum pump 08

To create vacuum, rotary pump or steam jet ejector can be used to create 0.13bar [35]. In rotary pumps the liquid is displaced by rotation of one or more members within a stationary housing. The selection of materials for designing rotary pumps is critical.

To create 0.13bar, Rotary pump is used, according to Perry’s chemical engineering handbook ([35] refer to figure 10-105: Vacuum levels attainable with various types of equipment. P10-59)

The capacity is 0.0033m3/s and the difference pressure is 314.2kPa. (Please refer to appendix 5.9 pipes and pumps design)

Theoretical pressure = v( pd - ps)102= 1.0 kW. (v is capacity of pump, m3/s)

The efficiency of reciprocating pumps is usually around 90% ([8] p435) Assume that the efficiency of rotary pump is 90%.

Therefore the Power at shaft is 1.2kW



A5.6.4 Recycle pump P07

Refer to appendix 5.9 (pipes and pumps calculation)



A5.7 Design of distillation column A5.7.1 Design of distillation column S03 Useful data from ASPEN Simulation

Table A5. 8 Useful data from ASPEN Simulation

Column S03 Minimum reflux ratio: 3.014161 Actual reflux ratio: 3.91841 Minimum number of stages: 61.17365 Number of actual stages: 108.859 Feed stage: 71.73564 Number of actual stages above feed: 70.73564 Reboiler heating required: 117027.3 Watt Condenser cooling required: 71291.43 Watt Distillate temperature: 298.0158 K Bottom temperature: 488.506 K Distillate to feed fraction: 0.233558

Plate hydraulic design is on the base of methods from Coulson & Richardson chemical Engineering, Volume 6. A mass balance: Feed, F = 548.5 kg/h Top product, D = 123.6 kg/h Vapour rate, V = D(1+3.91841) = 607.9 kg/h Bottom product, B = F – D = 424.9 kg/h Number of real stages = 109 (take reboiler as equivalent to one stage) Estimate physical properties: Column top: 250C, 99% 3,3-dimethylbutyraldehyde Vapor density = ρv 6.5 = 1.29 6.5 = 8.4 kg/m3 Liquid density = 798 kg/m3 Surface tension 5710-3 N/m Bottom, bottom temperature is 205 0C Vapor density = ρv 6.5 = 1 6.5 = 6.5 kg/m3 Liquid density = 7980.8 = 638 kg/m3 Surface tension 2310-3 N/m



Column diameter

The liquid-vapor flow factor L

vLV Vw

LwF

Where Lw = liquid mass flow rate, kg/s Vw = vapour mass flow rate, kg/s Bottom,

6.5

5 0.5638

vLV

L

LwF

Vw

top,

06.0798

4.857.0

L

vLV Vw

LwF

take tray spacing as 0.3m, From the figure 11.27, Vol. Coulson & Richardson Chemical Engineering, Bottom constant K1 = 0.028, Top constant K1 = 0.05 Correction for surface tensions Bottom K1 = (57/20)0.2 0.028 = 0.035

top K1 = (23/20)0.2 0.05 = 0.052 bottom uf = 0.032 [(638-6.5)/6.5]1/2 = 0.32 m/s top uf = 0.057 [(798-8.4)/8.4]1/2 = 0.63 m/s

design for 85% flooding at maximum flow rate bottom uv = 0.35 m/s 0.85 = 0.30 m/s

top uv = 0.52 m/s 0.85 = 0.44 m/s maximum volumetric flow-rate

bottom = 424.9/36005/6.5 = 0.09 m3/s top = 123.6/36004.92/8.4 = 0.02 m3/s

Net area required Bottom = 0.09/0.3 = 0.30 m2 Bottom = 0.02/0.44 = 0.05 m2 Column diameter Bottom = (0.3 4 / 3.14 )1/2 = 0.62 m Top = (0.05 4 / 3.14 )1/2 = 0.25 m Use same diameter above and below feed, reducing the perforated area for plates above the feed. The nearest standard pipe size is with an inside-diameter of 650mm.



Column height Column heght Hc = (109 – 1) 0.3 = 32.1 m Weir size Downcomer area Ad = 0.120.30 = 0.036 m2 Net area An = Ac – Ad = 0.3-0.036 = 0.254 m2 Active area Aa = Ac – 2Ad = 0.3-0.0362 = 0.218 m2 Hole area Ah take 6 % percent of active area as first trial = 0.0126 Weir length lw ( from figure 11.31) = 0.74 0.65 = 0.48 m Take weir height hw 30 mm Hole diameter dh 5 mm Plate thickness 5 mm Check weeping

Maximum liquid rate = (424.9/36005) = 0.59 kg/s Minimum liquid rate, at 70% turn-down = 0.590.7 = 0.41kg/s

3/2

750

lw

Lwh

Low

Where lw = weir length, m Lw = liquid flow rate, kg/s how = weir crest, mm liquid maximum how = 15 mm minimum how = 11 mm at minimum rate how+hw = 41 mm

from figure 11.30 , K2 = 29.5 the minimum design vapor through the hole uh

smdhK

uv

h /4.4)]4.25(9.02[

2/1

actual minimum vapour velocity = 0.7*0.09/0.0126 = 5m/s so minimum operating rate will be well above weep point.



A5.7.2 Design of distillation column S04 Useful data from ASPEN Simulation

Table A5. 9 Useful data from ASPEN Simulation

Column S04 Minimum reflux ratio: 0.009109 Actual reflux ratio: 0.010931 Minimum number of stages: 1.451957 Number of actual stages: 59.5581 Feed stage: 44.56951 Number of actual stages above feed: 43.56951 Reboiler heating required: 502.5792 Watt Condenser cooling required: 1042.906 Watt Distillate temperature: 346.9678 K Bottom temperature: 392.6399 K Distillate to feed fraction: 0.479131

Plate hydraulic design is on the base of methods from Coulson & Richardson chemical Engineering, Volume 6. A mass balance: Feed, F = 44.56 kg/h Top product, D = 8.325 kg/h Vapour rate, V = D(1+0.01) = 8.41 kg/h Bottom product, B = F – D = 36.23 kg/h Number of real stages = 60 (take reboiler as equivalent to one stage) Estimate physical properties: Column top: 250C, 99% 3,3-dimethylbutyraldehyde Vapor density = ρv 6.5 = 1.29 6.5 = 8.4 kg/m3 Liquid density = 798 kg/m3 Surface tension 5710-3 N/m Bottom, bottom temperature is 205 0C Vapor density = ρv 6.5 = 1 6.5 = 6.5 kg/m3 Liquid density = 7980.8 = 638 kg/m3 Surface tension 2310-3 N/m



Column diameter

The liquid-vapor flow factor L

vLV Vw

LwF

Where Lw = liquid mass flow rate, kg/s Vw = vapour mass flow rate, kg/s Bottom,

6.5

5 0.5638

vLV

L

LwF

Vw

top,

06.0798

4.857.0

L

vLV Vw

LwF

take tray spacing as 0.15m, From the figure 11.27, Vol.6 Coulson & Richardson Chemical Engineering, Bottom constant K1 = 0.018, Top constant K1 = 0.031 Correction for surface tensions Bottom K1 = (57/20)0.2 0.018 = 0.022

top K1 = (23/20)0.2 0.031 = 0.032 bottom uf = 0.022 [(638-6.5)/6.5]1/2 = 0.22 m/s top uf = 0.032 [(798-8.4)/8.4]1/2 = 0.33 m/s

design for 85% flooding at maximum flow rate bottom uv = 0.22 m/s 0.85 = 0.19 m/s

top uv = 0.33 m/s 0.85 = 0.28 m/s maximum volumetric flow-rate

bottom = 36.23/3600/0.8= 0.02 m3/s top = 8.41/36001.01/0.13 = 0.02 m3/s

Net area required Bottom = 0.02/0.19 = 0.11 m2 Bottom = 0.02/0.28 = 0.07 m2 Column diameter Bottom = (0.11 4 / 3.14 )1/2 = 0.37 m Top = (0.07 4 / 3.14 )1/2 = 0.27 m Use same diameter above and below feed, reducing the perforated area for plates above the feed. Nearest standard pipe size, inside diameter 40mm.



Column height Column heght Hc = (60 – 1) 0.15 = 8.8 m Weir size Downcomer area Ad = 0.120.1256 = 0.015 m2 Net area An = Ac – Ad = 0.1256-0.015 = 0.11 m2 Active area Aa = Ac – 2Ad = 0.1256-0.0152 = 0.096 m2 Hole area Ah take 10 % percent of active area as first trial = 0.01 Weir length lw ( from figure 11.31) = 0.74 0.4 = 0.3 m Take weir height hw 25 mm Hole diameter dh 4 mm Plate thickness 4 mm



A5.8 Design of heat exchangers A5.8.1 Cooler H01 for feed of gas-liquid separator Table A5. 10 Useful data for the design of heat exchange H01

H01 Tin Tout Heat transfer Shell/Tube (oC) (oC) (kJ/s)

Hot stream (process steam) 100 (t1) 25 (t2) 51.794 tube

Cold stream 22 (T1) 40(T2) Shell Flow rate of Cooling water desired Heat duty: Q = 51.794 kJ/s (table A5.8.1) Logarithmic Mean Temperature Difference (Tlm)

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

= 19.02 oC

where Tlm: log mean temperature difference. T1: inlet shell-side fluid temperature. T2: outlet shell-side fluid temperature. t1: inlet tube-side temperature. t2: outlet tube-side temperature. Cp = 4.18 kJ/kgoC (water) Cooling water follow rate M1 = Q/CpTm = 51.794/(4.18*19.02) = 0.65kg/s = 2340 kg/h Heat Exchange area (A) For “single pass”-tube/”single pass”-shell combination. General equation for heat transfer across a surface is:

Q = UACTm

where: Q: heat transferred per unit time [kW] U: the overall heat transfer coefficient = 0.250 [kW/m2 oC]

A: heat-transfer area [m2] Tm: the main temperature difference, the temperature driving force: Tm = FtTlm

Ft = the temperature correction factor. For “single pass”-tube/“single pass”-shell combination: Ft = 1 [30] hence Tm = 19.02 A = 10.9 m2 A5.8.2 heater H02 for feed of flash drum S02



Table A5. 11 Useful data for the design of heat exchange H02


Hot stream (MP steam) 220(t1) 220 (t2) 39.142Shell

Cold stream(process) 25(T1) 157(T2) tube Flow rate of medium pressure superiheated steam Heat duty: Q = 39.142 kJ/s (table A5.8.2) Steam condensed heat = 2200kJ /kg Flow rate of medium pressure supeiheated steam: Ms = 39.142/2200 3600 = 64 kg/h Heat Exchange area (A) For “single pass”-tube/”single pass”-shell combination. General equation for heat transfer across a surface is:

Q = UACTm Logarithmic Mean Temperature Difference (Tlm)

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

= 116.8 oC

where Tlm: log mean temperature difference. T1: inlet shell-side fluid temperature. T2: outlet shell-side fluid temperature. t1: inlet tube-side temperature. t2: outlet tube-side temperature. Q: heat transferred per unit time [kW]

U: the overall heat transfer coefficient = 0.250 [kW/m2 oC] A: heat-transfer area [m2] Tm: the main temperature difference, the temperature driving force: Tm = FtTlm

Ft = the temperature correction factor. For “single pass”-tube/“single pass”-shell combination: Ft = 1 [30] hence Tm = 116.8 A = 1.4 m2



A5.8.3 Cooler H03 for feed of distillation column S03 Table A5. 12 Useful data for the design of heat exchange H03


Hot stream (process stream) 157 (t1) 60 (t2) 93.002 tube


1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

= 70.2 oC

Cp = 4.18 kJ/kgoC (water) Cooling water follow rate M1 = Q/CpTm = 93.002/(4.1870.2) = 0.32 kg/s = 1152 kg/h Heat Exchange area (A) General equation for heat transfer across a surface is:

Q = UACTm

For “single pass”-tube/“single pass”-shell combination: Ft = 1 [30] hence Tm = 71.7 A = 5.2 m2



A5.8.4 Cooler H03 for compressed gas H04



Hot stream (process stream) 227 (t1) 60 (t2) 92.505 tube


1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

= 93.5 oC


Q = UACTm

For “single pass”-tube/“single pass”-shell combination: Ft = 1 [30] hence Tm = 93.5 A = 4 m2



A5.8.5 Condenser H11 of Column S03



Hot stream (process steam) 25 (t1) 25 (t2) 71.291 tube


1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

= 3.9 oC


Q = UACTm

For “single pass”-tube/“single pass”-shell combination: Ft = 1 [30] hence Tm = 3.9 oC A = 81.3 m2



A5.8.6 Reboiler H12 of Column S03




Cold stream(process) 215(T1) 215(T2) tube Flow rate of medium pressure superiheated steam Heat duty: Q = 117.56 kJ/s (table A5.8.6) Steam condensed heat = 2200kJ /kg Flow rate of medium pressure supeiheated steam: Ms = 117.56/2200 3600 = 192 kg/h Heat Exchange area (A) General equation for heat transfer across a surface is:

Q = UACTm Mean Temperature Difference (Tlm) = 5 oC A = 94 m2



A5.8.7 Condenser H13 of Column S04



Hot stream (process steam) 74(t1) 74 (t2) 1.04 tube


1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

= 42.4 oC

Cp = 4.18 kJ/kgoC (water) Cooling water follow rate M1 = Q/CpTm =1.04 /(4.18742.4) = 0.06 kg/s = 216 kg/h Heat Exchange area (A) General equation for heat transfer across a surface is:

Q = UACTm

For “single pass”-tube/“single pass”-shell combination: Ft = 1 [30] hence Tm = 42.4 oC A = 0.1 m2



A5.8.8 Reboiler H14 of Column S03




Cold stream(process) 119(T1) 119(T2) tube Flow rate of medium pressure superiheated steam Heat duty: Q = 0.5025 kJ/s (table A5.8.8) Steam condensed heat = 2200kJ /kg Flow rate of medium pressure supeiheated steam: Ms = 0.5025/2200 3600 = 1.6 kg/h Heat Exchange area (A) General equation for heat transfer across a surface is:

Q = UACTm Mean Temperature Difference (Tlm) = 101 oC A = 0.2 m2



A5.9 Design of pipes and pumps

In order to determine the power of the pump, first the pipe size has to be determined. the most economic pipe diameter would be the one which gives the lowest operating costs. The capital cost of pipe run increases with diameter while the pumping cost decrease with diameter increasing. The formula is given below:[8]

d,optimum = 260 G0.52 ρ-0.37

Where, G = mass flow rate in kg/s

ρ = density of fluid in kg/m3

To calculate the pressure drop the pipe friction factor needs to be known. This is a function of Reynolds Number. Friction loss (use fanning pressure drop equation):

10 1.84 0.16 1 4.84

2

3

-2

4.13 10

where pressure drop,kN/m ( ),

flow rate, kg/s,

density, kg/m ,

viscosity,mNm ,

pipe diameter, mm.

P G d

P kPa

G

s

d

(turbulent flow)

Darcy equition was used to calculate the pressure drop in laminar flow regime.

dg

uLfP

281.9

where f = friction factor, 64/Re 9.81 = conversion factor for kPa/m L = length in meter u = velocity in m/s g = acceleration due to gravity in m/s2 d = pipe diameter in m



A5.9.1 Design of raw material transfer pump P01 3,3-dimethylbutanol: Density (at 250C) = 817 kg/m3 (refer to part 3.2.4 for pure component properties) Viscosity: 4.578 mNs/m2 (4.578 cp) Estimation of pipe diameter required:

Typical velocity for liquid 2 m/s ([30], p.191) Mass flow = 130 kg/h = 0.03611 kg/s (refer to stream <1>, table 5.2, chapter 5) Volumetric flow = 0.03611/817 = 4.42e-5 m3/s Diameter of pipe = [(4.42e-5)/2*4/3.14]1/2 = 5.3 mm Take diameter as 12.5mm. Cross-sectional area = ¼ * 3.14 * (0.0125)2 = 1.2266e-4 m2 Actual velocity = 0.4 m/s Pressure –drop calculation:

Reynold’s number 3048105.12*10578.4*14.3

03611.0*44Re

33

d

G

It is a laminar flow Take the higher value, and design for a maximum flow rate of 20 per cent above the normal (average) flow:

kPaL

dg

uLfP 32.0

81.9*0125.0*2

2.1*4.0*817**3048/64*81.9

281.9

The loss through the bends and block values can be included in line pressure-loss calculation as an “equivalent length of pipe” [30] All the bends will be taken as 900 elbows of standard radius, equivalent length =30d, and the valves as plug valves, fully open, equivalent length = 18d.

Assumed length of pipes connected to pumps P01

Line to pump suction: Length = 2 m Bend, 1 30 12.5 10-3 = 0.4m Valve, 1 18 12.5 10-3 = 0.2m Total line length is 2.6m

Entry loss =2

2

u [30] at maximum design velocity = 0.1kPa

(where maximum velocity = 0.4 flow velocity) Control valve pressure drop, allows normal 140kPa (1.22) maximum 200 kPa Heat exchanger, allows normal 70 kPa

(1.22) maximum 200 kPa



Orifice, allows normal 15 kPa (1.22) maximum 22 kPa Line from pump discharge: Length: 44m Bend: 50.4 = 2 m Valve: 10.2 = 0.2 m Total 46.2m See table A5.18 for line calculation sheet for pump P101 on next page.

Table A5. 18 Pump P01 and line calculation sheet

PUMP P01 AND LINE CALCULATION SHEET

Job no. Sheet no. By: Checked:

Fluid 3,3-dimethylbutanol Discharge calculation Temperature© 25 Line size mm Density kg/m3 817 Flow Norm. Max. Units

Viscosity mNs/m2 4.578 Velocity 0.4 0.5 m/s Normal Flow kg/s 0.03611 Friction loss 0.27 0.32 kPa/m

Design Max. Flow kg/s 0.04332 Line length 46.2 m

20% above the normal flow Line loss 12.5 15.0 kPa

Suction Calculation Orifice 15 22 kPa Line size mm Control Valve 140 200 kPa

Flow Norm. Max. Units Equipment Velocity 0.4 0.5 m/s (a) Heat ex. kPa Friction loss 0.27 0.32 kPa/m (b) kPa Line length 2.6 m (c) kPa Line loss 0.7 0.8 kPa (6) Dynamic Loss 167.5 237 kPa Entrance 0.1 0.1 kPa

Strainer kPa Static head 1 m (1) Sub-total 0.8 0.9 kPa 9.8 9.8 kPa

Static head z1 0.7 0.7 m Equip Press (Max) 1050 1050 kPa

Contingency None None kPa Ρg Z1 6.8 6.8 kPa (7) Subtotal 1060 1060 kPa Equip. Press 101.3 101.3 kPa (7)+(6) Discharge Press (pd) 1228 1297 kPa (2) Sub-total 108.1 108.1 kPa Suction Press 107 107 kPa (2)-(1) (3) Suction Press 107.4 107.3 kPa

(8) Diff. Press 1121 1190 kPa

(4) VAP.PRESS kPa 114 121 m (3)-(4) (5) NPSH 107.4 107.3 kPa

Valve/(6)Control Valve

(5)/Ρg 13.4 13.4 m %Dyn. Loss

In this case, the capacity is 0.195 m3/h and the head is 121 m of water. Therefore, multi centrifugal pump is chosen (capacity range is 0.25-1000m3/h, typical head is 10-50m of water (single stage) 300m of water (multistage))



Theoretical Power = v( pd - ps)102 = 0.07 kW.

The efficiency is found to be 22% (from figure 10.62, p435, [6]) Therefore the power at shaft is 0.32kW



A5.9.2 Design of liquid transport pump P02

The detail calculation is the same as the calculation in section A5.9.1 for pump P01




Fluid TEMPO+Ru-Cat Discharge calculation

Temperature© 157 Line size mm

Density kg/m3 1000 Flow Norm. Max. Units

Viscosity mNs/m2 1 Velocity 0.99 1.2 m/s

Normal Flow kg/s 0.99 Friction loss 4.3 6 kPa/m

Design Max. Flow kg/s 1.2 Line length 5 5 m

20% above the normal flow Line loss 12.5 21.5 30Suction Calculation Orifice 15 15 22

Line size mm Control Valve 140 200 kPa Flow Norm. Max. Units Equipment Velocity 0.99 1.2 m/s (a) Heat ex. kPa Friction loss 4.3 6 kPa/m (b) kPa

Line length 2.6 m (c) kPa Line loss 11.2 15.6 kPa (6) Dynamic Loss 176.5 252 kPa Entrance 0.1 0.1 kPa



Contingency None None kPa Ρg Z1 6.8 6.8 kPa (7) Subtotal 1227 1302 kPa Equip. Press 101.3 101.3 kPa (7)+(6) Discharge Press (pd) 1403 1554 kPa (2) Sub-total 108.1 108.1 kPa Suction Press 107 107 kPa

(2)-(1) (3) Suction Press 96.8 92.4 kPa

(8) Diff. Press 1306 1462 kPa (4) VAP.PRESS kPa 133.1 149.0 m

(3)-(4) (5) NPSH 96.8 92.4 kPa

Valve/(6)

Control Valve

(5)/Ρg 9.9 9.4 m %Dyn. Loss

In this case, the capacity is 0.4357m3/h and the head is 149.0 m of water. Therefore, multi centrifugal pump is chosen (capacity range is 0.25-1000m3/h, typical head is 10-50m of water (single stage) 300m of water (multistage)) Theoretical Power = v( pd - ps)102 = 0.18 kW.

The efficiency is found to be 22% (from figure 10.62, p435, [8])

Therefore the power at shaft is 0.8kW



A5.9.3 Design of recycle pump P03 Table A5. 20 Pump P03 and line calculation sheet



Fluid Recycle liquid Discharge calculation Temperature© 215 Line size mm Density kg/m3 817 Flow Norm. Max. Units


Design Max. Flow kg/s 0.1416 Line length 5 m



Flow Norm. Max. Units Equipment Velocity 1.12 1.34 m/s (a) Heat ex. kPa Friction loss 4.7 5.3 kPa/m (b) kPa Line length 50 m (c) kPa Line loss 235 265 kPa (6) Dynamic Loss 178 248 kPa Entrance 0.1 0.1 kPa

Strainer kPa Static head 1 m (1) Sub-total 235 265 kPa 9.8 9.8 kPa


Contingency None None kPa Ρg Z1 6.8 6.8 kPa (7) Subtotal 1060 1060 kPa Equip. Press 658 658 kPa (7)+(6) Discharge Press (pd) 1238 1308 kPa (2) Sub-total 665 665 kPa Suction Press 430 400 kPa (2)-(1) (3) Suction Press 430 400 kPa


(4) VAP.PRESS kPa 82 93 m (3)-(4) (5) NPSH 430 400 kPa


(5)/Ρg 43.9 40.8 m %Dyn. Loss In this case, the capacity is 0.52 m3/h and the head is 93 m of water. Therefore, multi centrifugal pump is chosen (capacity range is 0.25-1000m3/h, typical head is 10-50m of water (single stage) 300m of water (multistage)) Theoretical Power = v( pd - ps)102 = 0.131 kW.




A5.9.4 Design of distillation feed pump P04




Fluid 3,3-Dimethylbutyral Discharge calculation Temperature© 60 Line size mm Density kg/m3 817 Flow Norm. Max. Units



20% above the normal flow Line loss 89 125 kPa


Flow Norm. Max. Units Equipment Velocity 1.5 1.8 m/s (a) Heat ex. kPa Friction loss 8.9 12.5 kPa/m (b) kPa Line length 2.6 m (c) kPa Line loss 23.1 32.4 kPa (6) Dynamic Loss 229 325 kPa Entrance 0.1 0.1 kPa



Contingency None None kPa Ρg Z1 6.8 6.8 kPa (7) Subtotal 668 668 kPa Equip. Press 101.3 101.3 kPa (7)+(6) Discharge Press (pd) 887 980 kPa (2) Sub-total 108.1 108.1 kPa Suction Press 85 76 kPa (2)-(1) (3) Suction Press 84.9 75.6 kPa


(4) VAP.PRESS kPa 82 92 m (3)-(4) (5) NPSH 84.9 75.6 kPa


(5)/Ρg 8.6 7.7 m %Dyn. Loss In this case, the capacity is 0.6696m3/h and the head is 92 m of water. Therefore, multi centrifugal pump is chosen (capacity range is 0.25-1000m3/h, typical head is 10-50m of water (single stage) 300m of water (multistage)) Theoretical Power = v( pd - ps)102 = 0.17 kW.




A5.9.5 Design of Tempo transferring pump P05 & P06




Fluid TEMPO Discharge calculation Temperature© 25 Line size mm Density kg/m3 800 Flow Norm. Max. Units

Viscosity mNs/m2 1 Velocity 0.1 0.3 m/s Normal Flow kg/s 0.0032 Friction loss 0.1 0.4 kPa/m

Design Max. Flow kg/s 0.01 Line length 2.6 m





Static head z1 0.7 0.7 m Equip Press (Max) 101.3 101.3 kPa

Contingency None None kPa Ρg Z1 6.8 6.8 kPa (7) Subtotal 101.3 101.3 kPa Equip. Press 101.3 101.3 kPa (7)+(6) Discharge Press (pd) 255.1 324.3 kPa (2) Sub-total 108.1 108.1 kPa Suction Press 107.7 107 kPa (2)-(1) (3) Suction Press 107.7 107 kPa

(8) Diff. Press 147.4 217.3 kPa

(4) VAP.PRESS kPa 15 22 m (3)-(4) (5) NPSH 107.7 107 kPa


(5)/Ρg 11.0 10.9 m %Dyn. Loss In this case, the capacity is 0.014m3/h and the head is 22m of water. Therefore, multi centrifugal pump is chosen (capacity range is 0.25-1000m3/h, typical head is 10-50m of water (single stage) 300m of water (multistage)) Theoretical Power = v( pd - ps)102 = 0.01 kW.




A5.9.6 Design of pump P07 in evaporator system




Fluid 3,3-Dimethylbutyral Discharge calculation Temperature© 162 Line size mm Density kg/m3 800 Flow Norm. Max. Units

Viscosity mNs/m2 1 Velocity 1.78 2.14 m/s Normal Flow kg/s 0.3178 Friction loss 1.73 2.4 kPa/m


20% above the normal flow Line loss 8.6 12 kPa




Static head z1 0.7 0.7 m Equip Press (Max) 10.1 10.1 kPa

Contingency None None kPa Ρg Z1 6.8 6.8 kPa (7) Subtotal 29.7 29.7 kPa Equip. Press 10.13 10.13 kPa (7)+(6) Discharge Press (pd) 193.3 263.7 kPa (2) Sub-total 16.8 16.8 kPa Suction Press 12.2 10.4 kPa (2)-(1) (3) Suction Press 12.2 10.4 kPa

(8) Diff. Press 181.1 253.3 kPa

(4) VAP.PRESS kPa 18.5 26.2 m (3)-(4) (5) NPSH 12.2 10.4 kPa


(5)/Ρg 1.2 1.0 m %Dyn. Loss In this case, the capacity is 0.7627m3/h and the head is 26m of water. Therefore, multi centrifugal pump is chosen (capacity range is 0.25-1000m3/h, typical head is 10-50m of water (single stage) 300m of water (multistage)) Theoretical Power = v( pd - ps)102 = 0.54 kW.




A5.9.7 Design of vacuum pump P08 in evaporator system




Fluid Tempo vapor Discharge calculation Temperature© 162 Line size mm Density kg/m3 1.2 Flow Norm. Max. Units

Viscosity Ns/m2 0.00002 Velocity 10 15 m/s Normal Flow kg/s 0.00327 Friction loss 0.07 0.10 kPa/m

Design Max. Flow kg/s 0.02 Line length 44 44 m Line loss 3.08 4.4 kPa


Flow Norm. Max. Units Equipment Velocity 10 15 m/s (a) Heat ex. kPa Friction loss 0.07 0.07 kPa/m (b) kPa Line length 5 5 m (c) kPa Line loss 0.35 0.50 kPa (6) Dynamic Loss 158.1 226.4 Entrance 0.06 0.14 kPa

Strainer kPa Static head 2 2 m (1) Sub-total 0.4 0.6 kPa 0.024 0.02 kPa

Static head 2 2 m Equip Press (Max) 100 100 kPa

Contingency None None kPa 0.0 0.0 kPa (7) Subtotal 100 100.0 kPa Equip. Press 15 15 kPa (7)+(6) Discharge Press 258.1 326.5 kPa (2) Sub-total 13.0 13.0 kPa Suction Press 12.5 12.3 kPa (2)-(1) (3) Suction Press 12.6 12.4 kPa

(8) Diff. Press 245.6 314.2 kPa

(4) VAP.PRESS 0.1 0.1 kPa m (3)-(4) (5) NPSH 12.5 12.3 kPa


1.27 1.25 m %Dynamic loss 89%

To create 0.13bar, Rotary pump is used, according to Perry’s chemical engineering handbook [34] The capacity is 0.0033m3/s and the difference pressure is 314.2kPa.

Theoretical pressure = v( pd - ps)102= 1.0 kW. (v is capacity of pump, m3/s)

The efficiency of reciprocating pumps is usually around 90% (from figure 10.62, p435, [6]) Assume that the efficiency of rotary pump is 90%.

Therefore the Power at shaft is 1.2kW



Appendix 5.10 compressor design In this case, flow rate 723 kg/h (560.46 m3/h) is smaller than 1000 m3/h; differential pressure is 9.5 bars. So, a reciprocating compressor is chosen[36]

the total pressure ratio is 10.5/1=10.5 maximum compression ratios of 3-4.5 per stage with a maximum of 8-12 per machine are commonly used. [37]

a two stage compressor will have a compression ratio of 3.2 and an inter-stage pressure pi of 3.3 bar. The total power Ptot = 154 Q ps lgΠ Where Ptot = total power, kw Q = volume flow m3/s Ps = absolute pressure inlet Π = pressure ratio Q = 0.2008 kg/s = 0.2008kg/s /1.29 kg/m3 = 0.1556 m3/s Ptot = 1540.15561ln10.5 = 56.3 kw Heat exchange H04 Refer to appendix 5.8.4 Heat transport surface is 4 m2, duty 92.5 kJ/s, cooling water 852 kg/h



A5.11 Equipment data sheets A5.11.1 Equipment data summary sheets Table A5. 25 Data summary sheet of storage tanks and vessels T01-T04

STORAGE TANKS AND VESSELS – DATA SUMMARY SHEET

EQUIPMENT NUMBER: NAME:

T01 Raw

material tank

T02 TEMPO

tank

T03 Product

tank

T04 De-active TEMPO

tank Vessel Vessel Vessel Vessel Pressure [bar] 1 1 1 1 Temperature [C] 20.0 20.0 20.0 30.0 Volume [m3] (1) 47.7 12.9 47.7 12.9 Diameter [m] 3.8 2.4 3.8 2.4 L or H [m] 4.2 2.9 4.2 2.9 Internals -Stirrer type -Stirrer diameter [m] -Stirrer height [m] -Power needed [kW] -Well-mixing time [s]

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

n.a

Number -Series: - Parallel:

1 -

1 -

1 -

1 -

Materials of construction (2)

CS

CS CS CS

Other:

Remarks: (1) The effect volume of vessel is 20% of total volume. (2) SS = Stainless Steel; CS = Carbon Steel



Table A5. 26 Data summary sheet of dissolving tank T05

STORAGE TANKS AND VESSELS - DATA SUMMARY SHEET


T05 dissolving

tank

D01

DecanterS01

Gas-liquid separator

S02 Flash drum

Vessel Vessel Vessel Vessel Pressure [bar] 1 1 1 1 Temperature [C] 50.0/20.0 25 25 25.0 Volume [m3] (1) 0.68 0.66 0.42 1.73 Diameter [m] 0.8 0.75 0.6 1.0 L or H [m] 1.4 1.5 1.5 2.2 Residence time (min) 4.6 10 10

Internals -Stirrer type -Stirrer diameter [m] -Stirrer height [m] -Power needed [kW] -Well-mixing time [s]

impeller

0.24

0.04

0.4

86

n.a

n.a

n.a.

n.a

n.a

n.a

n.a

n.a.

n.a

n.a

n.a

n.a

n.a.

n.a

n.a

Number -Series - Parallel

1 -

1 -

1 -

1 -


CS

CS CS CS

Other

Remarks: (1) The effect volume of vessel is 20% of total volume. (2) SS = Stainless Steel; CS = Carbon Steel



Table A5. 27 Data summary sheet of reactor R01

REACTOR R01 - DATA SUMMARY SHEET


R01 Stirred gas-liquid Jacketed Reactor

Vessel Jacket Pressure [bar] 10 1 Temperature [C] 30.0 20/25 Volume [m3] (2) 1.44 Heat transfer area [m2] 4.21 4.21 Diameter [m] 1.00 1.00 L or H [m] 1.84 1.84 Internals -Stirrer type: -Stirrer diameter [m] -Stirrer height [m] -Power needed [kW] -Well-mixing time [s]

Impeller

0.33

0.04

1.44

104


1 -

1 -


CS Al-Br

Other

Remarks: (1) The effect volume of vessel is 20% of total volume. (2) SS = Stainless Steel; CS = Carbon Steel, Al-Br = Aluminum Bronze



Table A5. 28 Data summary sheet of distillation column S03 & 04

DISTILLATION COLUMNs - DATA SUMMARY SHEET EQUIPMENT NUMBER: NAME:

S03 S04

Tray Column Tray Column Pressure [bar] 6.5 0.13 Temperature [C] 25/215 73/119 Volume [m3] 10.6 1.1 Diameter [m] 0.65 0.40 L or H [m] 32.1 8.8

Internals -Tray type -Tray number -Fixed packing Type : Shape : -Catalyst Type : Shape :

Sieve Trays

109

n.a. n.a

n.a. n.a.

Sieve Trays

60

n.a. n.a

n.a. n.a.

Number -Series : - Parallel :

1

1

Materials of construction (1) - Trays: SS314 Column: CS

- Trays: SS314 Column: CS

Other

Remarks: (1) SS = Stainless Steel; CS = Carbon Steel



Table A5. 29 Data summary sheet of heat exchangers H01-H04

HEAT EXCHANGERS - DATA SUMMARY SHEET

EQUIPMENT NUMBER NAME

H01 Separator S01 feed cooler

H02 Flash S02

feed heater

H03 Distillation column S03 Feed cooler

H04 Compressed

gas cooler

Single tube Sheet

Single tube Sheet

Single tube Sheet

Single tube Sheet

Substance -Tubes: -Shell:

mixture

cooling water

mixture

M.P. Steam

mixture

Cooling water

8% N2+92%O2

Cooling water

Duty [kW]

51.79 39.1 93.5 92.5

Heat exchange area [m2]

10.9 1.4 5.2 4

Number -Series - Parallel:

1 -

1 -

1 -

1 -

Pressure [bar] -Tubes -Shell

1.0

1.0

1

10

1.0

1.0

1.0

1.0

Temperature In / Out [oC]

- Tubes

- Shell

100.0 / 25.0

22.0 /40.0

15.0/157

220.0 / 220.0

157.0/ 25.0

22.0 /40.0

227/25

22.0/40.0

Special materials of construction (1)

Tubes : CS Shell : CS

Tubes : CS Shell : Al-Br



Other

Remarks: (1) CS = Carbon Steel; Al-Br = Aluminum Bronze



Table A5. 30 Data summary sheet of heat exchangers H11-H14

HEAT EXCHANGERS - DATA SUMMARY SHEET

EQUIPMENT NUMBER NAME

H11 S03 Condenser

H12 S03 Reboiler

H13 S04 Condenser

H14 S04 Reboiler

Single tubes Water cooled

Single tube Sheet

Single tubes Water cooled

Single tube Sheet

Substance -Tubes: -Shell:

product

Cooling water

alcohol

L.P. Steam

Alcohol

Cooling water

TEMPO

L.P. Steam

Duty [kW]

71.3 117.6 1.04 0.5

Heat exchange area [m2]

81.3 94.0 0.1 0.2

Number -Series - Parallel:

1 -

1 -

1 -

1 -

Pressure [bar] -Tubes -Shell

0.15

1.0

0.15

10

0.15

1.0

0.15

10


- Tubes

- Shell

25 / 25

20.0 / 22.0

215/215

220.0 / 220.0

73/ 73

20.0 / 30.0

119/119

220.0 / 220.0 Special materials of construction (1)





Other

Remarks: (1) CS = Carbon Steel; Al-Br = Aluminum Bronze



Table A5. 31 Data summary sheet of pumps P01-P04

PUMPS - DATA SUMMARY SHEET

EQUIPMENT NUMBER: NAME :

P01 Raw

material transport

P02 Catalyst transport

P03 Recycle pump

P04 S03 Feed transport

Type Number

Centrifugal

2

Centrifugal

2

Centrifugal

2

Centrifugal

2 Medium transferred

Raw material

Catalyst alcohol Alcohol/ aldehyde

Capacity [kg/s] [m3/s]

0.043

0.195

0.12

0.436

0.14

0.52

0.18

0.67

Density [kg/m3]

817 1500 817 817

Pressure [bar] Suction / Discharge

0.94 /10.5

0.94/10.5

0.91 /10.5

0.91 /6.5


25.0 / 25.0 157.0 / 157.0 215.0 / 215.0 60.0 / 60.0

Power [kW] -Theoretical: -Actual:

0.32

0.8

0.45

0.76

Number -Theoretical: -Actual (1):

2

2

2

2 Special materials of construction:

MS casing

MS casing

MS casing

MS casing

Other Double mechanical seals

Double mechanical seals



Remarks: (1) One installed spare included



Table A5. 32 Data summary sheet of pumps P05-P08

PUMPS - DATA SUMMARY SHEET


P05 TEMPO transport

P06 Catalyst

transport

P07 Evaporator

recycle

P08 Evaporator Top vacuum

Type Number

Centrifugal

2

Centrifugal

2

Centrifugal

2

Rotary

3 Medium transferred

TEMPO TEMPO TEMPO /Ru-catalyst

TEMPO


0.01

0.014

0.01

0.014

0.38

0.763

0.02

0.008

Density [kg/m3]

1000 1000 1500 1.2


0.9/ 2.2

0.9/ 2.2

0.10/ 2.63

0.14/ 3.27


25.0 / 25.0

25.0 / 25.0

162.0 / 162.0

162.0 / 162.0


0.1

0.1

1.0


2

2

2

3 Special materials of construction:

MS casing

MS casing

MS casing

MS casing





Remarks: (1) One installed spare included



Table A5. 33 Data summary sheet of compressor C01

COMPRESSOR - DATA SUMMARY SHEET


C01 Compressor

Type Number

reciprocating 1

Medium transferred

Mixture gas (8:92 N2:O2)


0.2 0.155

Density [kg/m3]

1.29


1/ 10.5


25.0 / 227.0


56.3


1

Special materials of construction: MS casing


Remarks:



Table A5. 34 Data summary sheet of evaporator S05

EVAPORATOR S05 - DATA SUMMARY SHEET


S05 evaporator

Vessel with heat-exchanger Pressure [bar] 0.10 Temperature [C] 162 Volume [m3] (2) 0.4

Diameter [m] 0.6

L or H [m] 1.4

Internals -Heat exchange type: -Heat exchange area [m2] -Heat duty [kw]

Single tube Sheet 0.3 7.3


1 -


CS

Other

Remarks: (1) The effect volume of vessel is 20% of total volume. (2) SS = Stainless Steel; CS = Carbon Steel, Al-Br = Aluminum Bronze



A5.11.2 Equipment data specification sheets

Table A5. 35 Equipment data specification sheet of toluene/product splitter C101



DISTILLATION COLUMN 01 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : S03 NAME : product distillation

General Data Service : - distillation/ extraction/ absorption/ Column type : - packed / tray / spray / Tray type : - cap / sieve / valve / Tray number (1) - Theoretical : 61 - Actual : 109 - Feed (actual) : 71 Tray Distance (HETP) [m] : 0.30 Tray Material: SS314 (2) Column Diameter [m] : 0.67 Column Material: CS (2) Column Height [m] : 32.1 Heating : - none /open steam /reboiler/ (3) Process Conditions

Stream Details Feed Top Bottom Reflux / Absorbent

Extractant / side stream

Temp. [oC] Pressure [bara] Density [kg/m3] Mass Flow [kg/s]

: 60 : 6.5 : 817 : 0.1524

: 25 : 6.5 : 798 : 0.03433

: 215 : 6.5 : 817 :0.1181

: 25 : 6.5 : 798 : 0.03433

Composition mol% wt% mol% wt% mol% wt% mol% wt% mol% wt% HEXANE HEXENE HEXANAL WATER TEMPO E A O2 N2

0.060.030.380.004.55

73.5721.370.000.03

0.280.160.020.010.000.09

99.250.010.16

0.280.160.020.010.000.09

99.250.010.16

0.28 0.16 0.02 0.01 0.00 0.09 99.2 0.01 0.16

Column Internals (4) Trays (5) Number of caps / sieve holes / : Active Tray Area [m2] : 0.218 Weir Length [mm] : 48 Diameter of chute pipe/hole/ [mm] : 5

Packing Not Applicable Type : Material : Volume [m3] : Length [m] : Width [m] : Height [m] :

Remarks: (1) Tray numbering from top to bottom. (2) SS = Stainless Steel; CS = Carbon Steel. (3) Reboiler is E103; operates with LP steam. (4) Sketch & measures of Column & Tray layout should have been provided. (5) Tray layout valid for whole column.



DISTILLATION COLUMN S04 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : S04 NAME : alcohol distillation

General Data Service : - distillation/ extraction/ absorption/ Column type : - packed / tray / spray / Tray type : - cap / sieve / valve / Tray number (1) - Theoretical : 15 - Actual : 60 - Feed (actual) : 44.5 Tray Distance (HETP) [m] : 0.15 Tray Material: SS314 (2) Column Diameter [m] : 0.40 Column Material: CS (2) Column Height [m] : 8.8 Heating : - none /open steam /reboiler/ (3) Process Conditions

Stream Details Feed Top Bottom Reflux / Absorbent

Extractant / side stream

Temp. [oC] Pressure [bara] Density [kg/m3] Mass Flow [kg/s]

: 157 : 1 : 800 : 0.0124

: 73 : 0.1 : 798 : 0.0023

: 157 : 0.1 : 800 :0.0101

: 73 : 0.1 : 798 : 0.0023

Composition mol% wt% mol% wt% mol% wt% mol%

wt% mol% wt%

HEXANE HEXENE HEXANAL WATER TEMPO E A O2 N2

0.000.000.070.00

26.2118.123.460.000.00

0.010.010.340.003.67

78.9717.000.000.00

0.000.000.010.00

31.463.960.310.000.00

0.01 0.01 0.34 0.00 3.67

78.97 17.00 0.00 0.00

Column Internals (4) Trays (5) Number of caps / sieve holes / : Active Tray Area [m2] : 0.1 Weir Length [mm] : 30 Diameter of chute pipe/hole/ [mm] : 4

Packing Not Applicable Type : Material : Volume [m3] : Length [m] : Width [m] : Height [m] :

Remarks: (2) Tray numbering from top to bottom. (3) SS = Stainless Steel; CS = Carbon Steel. (3) Reboiler is E103; operates with LP steam. (4) Sketch & measures of Column & Tray layout should have been provided. (5) Tray layout valid for whole column.



Table A5. 36 Equipment data specification sheet of batch jacketed reactor R101

BATCH JACKETED REACTOR R01– DATA SPECIFICATION SHEET

EQUIPMENT NR. : NAME:

R01 Operating: 1 stirred gas-liquid Jacketed Reactor

Number: 1 Reactor Jacket

Pressure [bar] 10 Pressure [bar]

1 (cooling water)

Temperature [C] 20/100 Temperature in/out [C]

20/25 (cooling water)

Volume [m3] (1) 1.44 Heat exchange area [m2]

3.334

Diameter [m] 1 Liquid height [m] 0.849 L or H [m] 1.44 Useful heat

exchange are [m2]

2.668

Internals -Stirrer type -Stirrer diameter [m] -Stirrer height [m] -Power needed [kW] -Well-mixing time [s]

Impeller 0.33 0.0416 0.168 115

Heat transferred [kJ] (2)

10843.6(b) 29146.1(c)

Time consuming [hrs] (2)

4.238 (b) 0.788 (c)


1 -


1 -


CS


Al-Br

Other : Remarks:

(1) The effect volume of vessel is 20% of total volume. (2) Duty (a) is for liquid heating, duty (b) is for temperature maintain, duty (c) is for liquid

cooling. (3) SS = Stainless Steel; CS = Carbon Steel (4) Al-Br = Aluminum Bronze



Table A5. 37 Equipment data specification sheet of distillation column’s feed heater E101

DISTILLATION COLUMN’S FEED HEATER E101 – DATA SPECIFICATION SHEET

QUIPMENT NUMBER : E101 In Series : 1 NAME : C101 Feed heater In Parallel : none

General Data Service : - Heat Exchanger - Vaporizer - Cooler - Reboiler - Condenser Type : - Fixed Tube Sheets - Plate Heat Exchanger - Floating Head - Finned Tubes - Hair Pin - Thermosyphon - Double Tube - Single Tube Position : - Horizontal - Vertical Capacity [kW] : 25.02 (1) (Calc.) Heat Exchange Area [m2] : 2.1 (2) (Calc.) Overall Heat Transfer Coefficient [W/m2oC] : 250 (Approx.) Log. Mean Temperature Diff. (LMTD) [oC] : 48.27 (3)

Passes Tube Side : 1 Passes Shell Side : 1

Correction Factor LMTD (min. 0.75) : 1.0 Corrected LMTD [oC] : 48.27

Process Conditions

Medium :

Mass Stream [kg/h] : Mass Stream to - Evaporize [kg/s] : - Condense [kg/h] :

Average Specific Heat [kJ/kgoC] : Heat of Evap. / Condensation[kJ/kg] :

Temperature IN [oC] : Temperature OUT [oC] :

Pressure [bar] : Material (4) :

Shell Side Tube Side LP steam

74.39 -

74.39

- 2200

110 110

1.5 CS

Toluene/precursor

751.1040

0 751.1040

1.628 -

20 88

1 CS

Remarks: (1) Capacity = Average Specific Heat Total Amount of Feed (2) Capacity = Area Overall Heat Transfer Coefficient LMTD

(3)

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

[15] (p598 equation 12.4)

(4) CS = Carbon Steel



Table A5. 38 Equipment data specification sheet of distillation column’s reboiler E103

DISTILLATION COLUMN’S REBOILER E103 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : E103 In Series : 1 NAME : C101 Reboiler In Parallel : none

General Data Service : - Heat Exchanger - Vaporizer - Cooler - Reboiler - Condenser Type : - Fixed Tube Sheets - Plate Heat Exchanger - Floating Head - Finned Tubes - Hair Pin - Thermosyphon - Double Tube - Single Tube Position : - Horizontal - Vertical Capacity [kW] : 102.202 (1) (Calc.) Heat Exchange Area [m2] : 37.63 (2) (Calc.) Overall Heat Transfer Coefficient [W/m2oC] : 250 (Approx.) Log. Mean Temperature Diff. (LMTD) [oC] : 10.87 (3)


Correction Factor LMTD (min. 0.75) : 1.0 Corrected LMTD [oC] : 10.87

Process Conditions

Medium :

Mass Stream [kg/h] : Mass Stream to - Evaporize [kg/h] : - Condense [kg/h] :

Average Specific Heat [kJ/kg oC] : Heat of Evap. / Condensation[kJ/kg] :

Temperature IN [oC] : Temperature OUT [oC] :


Shell Side Tube Side LP steam

167.24 -

167.24

1.88 2200

110.0 110.0

1.5 CS

Precursor

1205.93

1205.93 -

-

98.2 100.02

0.15 CS

Remarks: (3) Heat capacity simulate from ASPEN (4) Heat Capacity = Overall Heat Transfer Coefficient Heat transfer Area LMTD

(4)

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

[15] (p598 equation 12.4)

(4)CS = Carbon Steel



Table A5.39 Equipment data specification sheet of distillation column’s condenser E102

DISTILLATION COLUMN’S CONDENSER E102 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : E102 In Series : 1 NAME : C101 Overhead Condenser In Parallel : none

General Data Service : - Heat Exchanger - Vaporizer - Cooler - Reboiler - Condenser (Water cooled) Type : - Fixed Tube Sheets - Plate Heat Exchanger - Floating Head - Finned Tubes - Hair Pin - Thermosyphon - Double Tube - Single Tube Position : - Horizontal - Vertical Capacity [kW] : 121.022 (1) (Calc.) Heat Exchange Area [m2] : 26.27 (2) (Calc.) Overall Heat Transfer Coefficient [W/m2 oC] : 250 (Approx.) Log. Mean Temperature Diff. (LMTD) [oC] : 18.43 (3)


Correction Factor LMTD (min. 0.75) : 1.0 (4) Corrected LMTD [oC] : 18.43

Process Conditions

Medium :

Mass Stream [kg/s] : Mass Stream to - Evaporize [kg/s] : - Condense [kg/s] :

Average Specific Heat [kJ/kgoC] : Temperature IN [oC] : Temperature OUT [oC] :


Shell Side Tube Side Cooling water

2.90

n.a.

4.18 20.0 30.0

Atm. n.a.

Toluene / Precursor

0.2768

0.2768

- 53.7 34.0

0.15 CS

Remarks: (1) Heat capacity simulate from ASPEN (2) Capacity = Area Overall Heat Transfer Coefficient LMTD

(3)

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

[15] (p598 equation 12.4)

(4) [15] (p697, for single tube heat exchanger typical effectiveness 1) (5) CS = Carbon Steel



Table A5.40 Equipment data specification sheet of product cooler E104

PRODUCT COOLER E104 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : E104 In Series : 1 NAME : C101 Product cooler In Parallel : none

General Data Service : - Heat Exchanger(Water cooled) - Vaporizer - Condenser - Cooler - Reboiler Type : - Fixed Tube Sheets - Plate Heat Exchanger - Floating Head - Finned Tubes - Hair Pin - Thermosyphon - Double Tube - Single Tube Position : - Horizontal - Vertical Capacity [kW] : 1.55 (1) (Calc.) Heat Exchange Area [m2] : 0.26 (2) (Calc.) Overall Heat Transfer Coefficient [W/m2oC] : 250 (Approx.) Log. Mean Temperature Diff. (LMTD) [oC] : 24.19 (3)


Correction Factor LMTD (min. 0.75) : 1.0 (4) Corrected LMTD [oC] : 24.19

Process Conditions

Medium :

Mass Stream [kg/s] : Mass Stream to - Evaporize [kg/s] : - Condense [kg/s] :

Average Specific Heat [kJ/kgoC] : Temperature IN [oC] : Temperature OUT [oC] :


Shell Side Tube Side Cooling water

0.037

n.a.

4.18 20.0 30.0

Atm. n.a.

Toluene / Precursor

0.0177

0.0177

- 98.2 25

0.15 CS

Remarks: (1) Heat capacity simulate from ASPEN (2) Capacity = Area Overall Heat Transfer Coefficient LMTD

(3)

1 2 2 1

1 2

2 1

lnlm

T t T tT

T t

T t

[15] (p598 equation 12.4)

(4) [15] (p697, for single tube heat exchanger typical effectiveness 1) (5) CS = Carbon Steel



Table A5.41 Equipment data specification sheet of lipase B separation filter F101

FILTER F101 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER: F101 In Series : 1 NAME : Lipase B separation filter In Parallel : none

Service : Lipase B separation filter Type : Microfilter Number : 1

Operating Conditions & Physical Data Separation phases : Solid/Liquid Exit temperature [0C] : 30 Average filtrate flux [l/m2h] : 20 Max. concentration of solid [g/liter] : 600 Filtration mode : 1 ( Batch concentration =1, Feed and bleed=2) For each component rejection coefficient : 0

Unit power consumption [kW/m2] : 0.2 Membrane replacement cost [operating hrs] : 2000 Unit membrane cost [$/m2] : 200

Design Mode Concentration factor(Feed/Retentate) : 1 Max. Area [m2] : 80

Rating Mode

Membrane Area [m2] : 80 Number of units : 1

Remarks: - The filter chosen from SUPERPRO Design. Default design of Microfilter is used - Design the max. membrane area as 80 m2

Cost



Table A5.42 Equipment data specification sheet of lanthanum hydroxide separation filter F102

FILTER F102 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : F102 In Series : 1 NAME : La(OH)3 separation filter In Parallel : none

Service : La(OH)3 separation filter Type : Microfilter Number : 1

Operating Conditions & Physical Data Separation phases : Solid/Liquid Exit temperature [0C] : 30 Average filtrate flux [l/m2h] : 20 Max. concentration of solid [g/liter] : 600 Filtration mode : 1 ( Batch concentration =1, Feed and bleed=2) For each component rejection coefficient : 0

Unit power consumption [kw/m2] : 0.2 Membrane replacement cost [operating hrs] : 2000 Unit membrane cost [$/m2] : 200

Design Mode Concentration factor(Feed/Retentate) : 1 Max. Area [m2] : 80

Rating Mode

Membrane Area [m2] : 80 Number of units : 1

Remarks: - The filter chosen from SUPERPRO Design. Default design of Microfilter is used - Design the max. membrane area as 80 m2

Cost



Table A5.43 Equipment data specification sheet of toluene transport pump P101

PUMP P101 – DATA SPECIFICATION SHEET

EQUIPMENT NUMBER : P101 Operating : 1 NAME : Toluene Transport Pump Installed Spare : 1 Service : Toluene transfer pump Type : Centrifugal Number : 2

Operating Conditions & Physical Data Pumped liquid: Toluene Temperature (T) [oC] : 25.0 Density () [kg/m3] : 866 Viscosity () [Ns/m2] : 0.0001

Power Capacity (v) [m3/s] : 0.001 Suction Pressure (ps) [bar] : 0.939 Discharge Pressure (pd) [bar] : 4.91

Theoretical Power [kW] : 0.4 { = v( pd - ps)102 }

Pump Efficiency [-] : 0.27 [15] (p435) Power at Shaft [kW] : 1.5

Construction Details (1) RPM : 3000 Drive : Electrical Type electrical motor : Tension [V] : 380 Rotational direction : Clock / Counter Cl. Foundation Plate : Combined / two parts Flexible Coupling : Yes Pressure Gauge Suction : No Pressure Gauge Discharge : Yes Min. Overpressure above pv/pm [bar] : 0.1

Nominal diameter Suction Nozzle […] : Discharge Nozzle […] : Cooled Bearings : Yes / No Cooled Stuffing Box : Yes / No Smothering Gland : Yes / No If yes - Seal Liquid : Yes / No - Splash Rings : Yes / No - Packing Type : - Mechanical Seal : Yes / No - N.P.S.H. [m] : { = pmg }

Construction Materials (2) Pump House : MS Pump Rotor : HT Steel Shaft : HT Steel Special provisions : none

Operating Pressure [bar] : 2

Wear Rings : Shaft Box :

Test Pressure [bar] : Remarks:

(1) Double mechanical seals and seal fluid required for LPG service. Further details to be specified by Rotating Equipment specialist

(2) MS = Mild Steel; HT Steel = High Tensile Steel



Table A5.44 Equipment data specification sheet of liquid transport pump P102


EQUIPMENT NUMBER : P102 Operating : 1 NAME : Liquid Transport Pump Installed Spare : 1 Service : Liquid transfer pump Type : Centrifugal Number : 2

Operating Conditions & Physical Data Pumped liquid: Catalyst (lanthanum isopropoxide) dissolved in toluene Temperature (T) [oC] : 25.0 Density () [kg/m3] : 873.6 Viscosity () [Ns/m2] : 0.0001




Construction Details (1) RPM : 3000 Drive : Electrical Type electrical motor : Tension [V] : 380 Rotational direction : Clock / Counter Cl.Foundation Plate : Combined / two parts Flexible Coupling : Yes Pressure Gauge Suction : No Pressure Gauge Discharge : Yes Min. Overpressure above pv/pm [bar] : 0.1





Test Pressure [bar] : Remarks: (1) Double mechanical seals and seal fluid required for LPG service. Further details to be specified by Rotating Equipment specialist. (2) MS = Mild Steel; HT Steel = High Tensile Steel



Table A5.45 Equipment data specification sheet of liquid transport pump P103


EQUIPMENT NUMBER : P103 Operating : 1 NAME : Liquid Transport Pump Installed Spare : 1 Service : Liquid transfer pump Type : Centrifugal Number : 2

Operating Conditions & Physical Data Pumped liquid: Reaction mixture Temperature (T) [oC] : 25.0 Density () [kg/m3] : 904.5 Viscosity () [Ns/m2] : 0.0001




Construction Details (1) RPM : 3000 Drive : Electrical Type electrical motor : Tension [V] : 380 Rotational direction : Clock / Counter Cl.Foundation Plate : Combined / two parts Flexible Coupling : Yes Pressure Gauge Suction : No Pressure Gauge Discharge : Yes Min. Overpressure above pv/pm [bar] : 0.1








Table A5.46 Equipment data specification sheet of distillation column’s feed pump P104

PUMP P104 – DTA SPECIFICATION SHEET

EQUIPMENT NUMBER : P104 Operating : 1 NAME : C101 feed Pump Installed Spare : 1 Service : Column feed pump Type : Centrifugal Number : 2

Operating Conditions & Physical Data Pumped liquid: Reaction mixture Temperature (T) [oC] : 25.0 Density () [kg/m3] : 897 Viscosity () [Ns/m2] : 0.0001




Construction Details (1) RPM : 3000 Drive : Electrical Type electrical motor : Tension [V] : 380 Rotational direction : Clock / Counter Cl. Foundation Plate : Combined / two parts Flexible Coupling : Yes Pressure Gauge Suction : No Pressure Gauge Discharge : Yes Min. Overpressure above pv/pm [bar] : 0.1








Table A5.47 Equipment data specification sheet of vacuum pump P105


EQUIPMENT NUMBER : P105 Operating : 1 NAME : Vacuum Pump Installed Spare : 1 Service : create vacuum Type : rotary Number : 2

Operating Conditions & Physical Data Pumped medium: air Temperature (T) [oC] : 20.0 Density () [kg/m3] : 1.2 Viscosity () [Ns/m2] : 0.00002 Vapour Pressure (pv) [bar] : 0.15-1.0 at Temperature [oC] : 20.0



Pump Efficiency [-] : 0.9 Power at Shaft [kW] : 2.8

Construction Details (1) RPM : Drive : Electrical Type electrical motor : Tension [V] : 380 Rotational direction : Clock / Counter Cl.Foundation Plate : Combined / two parts Flexible Coupling : Yes Pressure Gauge Suction : No Pressure Gauge Discharge : Yes Min. Overpressure above pv/pm [bar] : 0.1



Operating Pressure [bar] : 0.15bar



(1) Double mechanical seals and seal fluid required for LPG service. Further details to be specified by Rotating Equipment specialist.




Table A5.48 Equipment data specification sheet of reflux pump P106


EQUIPMENT NUMBER : P106 Operating : 1 NAME : Toluene Reflux Pump Installed Spare : 1 Service : Toluene reflux Type : centrifugal Number : 2

Operating Conditions & Physical Data Pumped liquid: Toluene/Product Temperature (T) [oC] : 34.0 Density () [kg/m3] : 854.7 Viscosity () [Ns/m2] : 0.0001 Vapour Pressure (pv) [bara] : 0.15 at Temperature [oC] : 34.0

Power Capacity (v) [m3/s] : 1 10-4

Suction Pressure (ps) [bar] : 0.297 Discharge Pressure (pd) [bar] : 2.359













Table A5.49 Equipment data specification sheet of toluene transport pump P107


EQUIPMENT NUMBER : P107 Operating : 1 NAME : Toluene Transfer Pump Installed Spare : 1 Service : Toluene transfer Type : centrifugal Number : 2

Operating Conditions & Physical Data Pumped liquid: Technical grade toluene Temperature (T) [oC] : 34.0 Density () [kg/m3] : 854.7 Viscosity () [Ns/m2] : 0.0001 Vapour Pressure (pv) [bar] : 0.15 at Temperature [oC] : 34.0

Power Capacity (v) [m3/s] : 2.2 10-4

Suction Pressure (ps) [bar] : 0.201 Discharge Pressure (pd) [bar] : 3.99 Theoretical Power [kW] : 0.09 { = v( pd - ps)102 }









(2) (2) MS = Mild Steel; HT Steel = High Tensile Steel



Table A5.50 Equipment data specification sheet of product transport pump P108


EQUIPMENT NUMBER : P 108 Operating : 1 NAME : Product Transfer Pump Installed Spare : 1 Service : Product transfer Type : diaphragm Number : 2

Operating Conditions & Physical Data Pumped liquid: Product (precursor of Prozac) Temperature (T) [oC] : 100.0 Density () [kg/m3] : 1202.8 Viscosity () [Ns/m2] : 0.0003 Vapour Pressure (pv) [bar] : 0.15 at Temperature [oC] : 100.0

Power Capacity (v) [m3/s] : 1.5 10-5

Suction Pressure (ps) [bar] : 0.60 Discharge Pressure (pd) [bar] : 4.80



Construction Details (1) RPM : Drive : Type electrical motor : Tension [V] : Rotational direction : Clock / Counter Cl.Foundation Plate : Combined / two parts Flexible Coupling : Yes Pressure Gauge Suction : No Pressure Gauge Discharge : Yes Min. Overpressure above pv/pm [bar] : 0.1










Appendix 6.1


Table A6.1.1: Summary of purchase costs of dissolving tanks, storage vessels and distillation column

Name Equipment Type/ D (m) H (m) M or C Press. Costs Costs Costs Costs Figure/ Curve Costs (bar) equip. Intern. total total Quantity. (UK £) Factor Factor (UK £) (UK £) (UK £) (€)

T01 Vessel Vertical 3.8 4.2 CS 1 fig.6-4 C4- 9,400 1 1 £9,400 £9,400 €12,784

T02 Vessel Vertical 2.4 2.9 CS 1

fig.6-4 C2 5,000 1 1 £5,000 £5,000 €6,800


fig.6-4 C4- 9,400 1 1 £9,400 £9,400 €12,784


fig.6-4 C2 5,000 1 1 £5,000 £5,000 €6,800


fig.6-4 C1 3,000 1 1 £3,000 £3,000 £6,000 €8,160

D01 Vessel Vertical 0.75 1.5 CS 1

fig.6-4 C1 3,000 1 1 £3,000 £1,000 £4,000 €5,440

S01 Vessel Vertical 0.6 1.5 CS 1

fig.6-4 C1 2,000 1 1 £2,000 £2,000 €2,720

S02 Vessel Vertical 1.0 2.2 CS 1

fig.6-4 C1 5,000 1 1 £5,000 £5,000 €6,800

S03 Vessel Vertical 0.65 32.1 CS 6.5 fig.6-4 C1-- 22,000 1 1.1 £24,200 Trays Sieve 0.65 SS fig. 6-6 C! 108 150 1.7 £27,500 £51,700 €70,000

S04 Vessel Vertical 0.4 8.8 CS 0.1 fig.6-4 C1 7,000 1 2.2 £15,400 Trays Sieve 0.4 SS fig. 6-6 C! 60 50 1.7 £5,100 £20,500 €27,880

Total £118,000 €160,480

Remarks: - The figure used in the table are from J. M. Coulson & J. F. Richarson, 1979, Chemical Engineering, volume 6, p 223-224 - Pressure factor of column is estimated as same as the factor when pressure is 50-60 bar - The rate of exchange (ROE) between euro € and UK £ is 1.36



Table A6.1.2: Summary of purchase costs of reactor, evaporator, compressor and pumps

Reactor& Vessels @ 1992

Name Equip. Type Capacity Constant Index Comment Equation Cost Cost total

Cost total

(UK £) (UK £) (€) S C n Ce

R01 Reactor Jacketed 1.44 m3 £8,000 0.4 CS Ce=CSn £9,150 £9,150 €12,440C01 compressor 56kw £800 0.8 £20,000 £20,000 €27,200

S05 evapotator 0.4 m3 13,000 0.52 £8,000 £8,000 €10,880

H01-H08 pumps £36,760 €50,000

Total £73,910 €100,520

Remarks: - The equation is from J. M. Coulson & J. F. Richarson, 1979, Chemical Engineering, volume 6, p225; Total cost of Pumps is estimated to € 50,000,on the basis of Dutch Association of Cost Engineers, Prijzenboekje (21st edition, Dec 2000, p35). This cost have been considered the space factor of 2.- The rate of exchange (ROE) between euro € and UK £ is 1.36

Table A6.1.3: Summary of purchase costs of heat exchangers

Heat exchangers & Filters @1992

Name M of C Curve Surface Cost Type Factor Press. Factor Cost Cost

Sh/Tubes (in fig.6-3 (m2) (UK £) (bar) (UK £) (€)

H01 CS/Brass c2 11 4000 Float Hd. 1.0 10 1 4000 5,440H02 CS/Brass c2 2 2,000 Float Hd. 1.0 10 1 2,000 2,720H03 CS/Brass c2 5 2,000 Float Hd. 1.0 10 1 2,000 2,720H04 CS/Brass c2 4 2,000 Float Hd. 1.0 10 1 2,000 2,720H11 CS/Brass c2 81 18,000 Float Hd. 1.0 10 1 18,000 24,480H12 CS/Brass c2 94 20,000 Float Hd. 1.0 10 1 20,000 27,200H13 CS/Brass c2 1 2,000 Float Hd. 1.0 10 1 2,000 2,720

H14 CS/Brass c2 1 2,000 Float Hd. 1.0 10 1 2,000 2,720

Total Heat exchangers (@1992) 52,000 70,720

Remarks: - The figure 6.3 used in the table from J. M. Coulson & J. F. Richarson, 1979, Chemical

Engineering, volume 6, p222 - The equation from J. M. Coulson & J. F. Richarson, 1979, Chemical Engineering, volume 6, p222 - Pressure factor of column is estimated as same as the factor when pressure is 50-60 bar - The rate of exchange (ROE) between euro € and UK £ is 1.36



Appendix 6.2

Capital investment cost calculation From the purchased equipment costs (refer to appendix A6.1 for summary of equipment cost) the following terms is calculated: Direct Capital, Indirect Capital Costs and Fixed Capital Costs in UK@1992 by escalating given totals to UK@2001 (approximate interest is 8% per year). Lang’s method is used for calculation.

Table A6.2.1: Capital investment cost calculation

Purchased Equipment Costs ( PEC )

Items @ 1992 Cost (UK £)

Reactors 73,910

Columns & Storage vessels 118,000

Heat Exchangers 52,000

Purchased Equipment Costs @ 1992 243,910

Total Purchased Equipment Costs @ 2002 (Index correction, UK, @ 1992-2001, 8% per year: 2.159)

526,600

Total Direct Capital Cost @ 2002 (Lang factor, process type “Fluids”: 3.4 with respect to PEC)

1,790,440

Total Indirect Capital Cost @ 1992 ( Lang factor, process type “Fluids”: 0.45 with respect to PPC)

805,700

Fixed Capital Costs @ 1992 ( Total Direct Capital Cost + Total Indirect Capital Cost )

2,596,140

Total Direct Capital Cost @ 2002 (Dfl ) (ROE=3.0)

5,371,320

Total Indirect Capital Cost @ 2002 (Dfl ) (ROE=3.0)

2,417,100

Fixed Capital Costs @ 2002 (Dfl ) (ROE=3.0)

7,788,420

Fixed Capital Costs @ 2002 (Euro) (ROE=2.2)

3,540,190

The Lang’s factors from J. M. Coulson & J. F. Richarson, 1979, Chemical Engineering, volume 6 (chapter 6, table 6.1 of typical factors for estimation of project fixed capital cost )

Remarks:

- Rate of exchange (ROE) between Dfl and UK £ is 3.0 - Rate of exchange (ROE) between Dfl and Euro is 2.2



Apppendix 6-4

Appendix 6.3


Table A6.3.1 Net present and future values for interest of 8%

NET FUTURE VALUES (1) NET PRESENT VALUES No Discount Discounted, Accumulated END CAPIT. COSTS CASH FLOW DISC. CAPIT. CASH

YEAR ANN. ACCUM. ANN. ACCUM. NFV FACT. COSTS FLOW NPVNO. Interest ACCUM. ACCUM.

(k€) (k€) (k€ ) (k€ ) (k€ ) 8.0% (k€) (k€l) (k€l)

1 3,980 3,980 1 3,980 3,980 -3,9802 2,420 2,420 -1,560 0.926 2241 2,240.92 -1,7393 2,420 4,840 860 0.857 2074 4,315 3354 2,420 7,260 3,280 0.794 1921 6,236 2,2565 2,420 9,680 5,700 0.735 1779 8,015 4,0356 2,420 12,100 8,120 0.681 1648 9,663 5,6837 2,420 14,520 10,540 0.63 1525 11,188 7,2088 2,420 16,940 12,960 0.583 1411 12,599 8,6199 2,420 19,360 15,380 0.54 1307 13,905 9,925

10 2,420 21,780 17,800 0.5 1210 15,115 11,13511 2,420 24,200 20,220 0.463 1120 16,236 12,256

ACCUM. 3,980 24,200 133,100 97,280 8 20,216 16,236 12,256RATIO: [Cash Flow / Capital] @ Disc. 4 NET PRESENT VALUE:

[Cash Flow - Capital] @ Disc. 12,356

N.B. : 1. Cash Flows "Before Tax". 2. Earning Power = Interest, for which [Cash Flow - Capital]@Disc. Disc. Factor = 1/(1 + r) n with r = interest fraction



Apppendix 6-5

Table A6.3.2 Net present and future values for interest of 60.3%

NET FUTURE VALUES (1) NET PRESENT VALUES

No Discount Discounted, Accumulated END CAPIT. COSTS CASH FLOW DISC. CAPIT. CASH

YEAR ANN. ACCUM. ANN. ACCUM. NFV FACT. COSTS FLOW NPV NO. @ ACCUM. ACCUM.

kDfl kDfl kDfl kDfl kDfl 90.2% kDfl kDfl kDfl

1 3,980 3,980 1 3,980 3,980 -3,9802 1 -3,9803 2,420 2,420 -1,560 0.6238 1510 1,509.67 -2,4704 2,420 4,840 860 0.3892 942 2,451 -1,5295 2,420 7,260 3,280 0.2428 588 3,039 -9416 2,420 9,680 5,700 0.1514 367 3,405 -5757 2,420 12,100 8,120 0.0945 229 3,634 -3468 2,420 14,520 10,540 0.0589 143 3,777 -2039 2,420 16,940 12,960 0.0368 89 3,866 -11410 2,420 19,360 15,380 0.0229 56 3,921 -5911 2,420 21,780 17,800 0.0143 35 3,956 -24

ACCUM. 24,200 133,100 97,280 4 7,957 3,956 -24RATIO : [Cash Flow / Capital] @ Disc. 1 NET PRESENT VALUE : [Cash Flow - Capital] @ Disc. 0

N.B. : 1. Cash-Flows "Before Tax". 2. Earning Power = Interest, for which [Cash Flow - Capital]@Disc. = 0 Disc. Factor = 1/(1 + r) n with r = interest fraction

When interest is 60.3%, the [Cash Flow – Capital] = 0 DCFROR = 60.3%

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