4th Year Project

Plant Design for Production of

n-Butyraldehyde by Hydroformylation of Propylene

Session: 2005-2009 Project Advisors Prof. Dr. Muhammad Zafar Noon Mr. Muhammad Faheem Project Members Hafiz Sajid Sattar Muhammad Waqas Saeed Ur Rehman Saad Ullah Mirza

2005-Chem-62 2005-Chem-86 2005-Chem-98 2005-Chem-74

DEPARTMENT OF CHEMICAL ENGINEERING UNIVERISITY OF ENGINEERING & TECHNOLOGY PLANT DESIGN FOR Production

of

n-Butyraldehyde by Hydroformylation of Propylene This report is submitted to department of Chemical Engineering, University of Engineering & Technology Lahore- Pakistan for the partial fulfillment of the requirements for the Bachelor’s Degree In

CHEMICAL ENGINEERING Internal Examiner:

Sign : Name:

External Examiner

Sign : Name:

DEPARTMENT OF CHEMICAL ENGINEERING UNIVERISITY OF ENGINEERING AND TECHNOLOGY LAHORE-PAKISTAN

DEDICATED TO Our Beloved Parents, Respected Teachers, And Sincere Friends!

Page i

ACKNOWLEDGEMENT All praises to ALMIGHTY ALLAH, who provided us with the strength to accomplish the final year project. All respects are for His HOLY PROPHET (PBUH), whose teachings are true source of knowledge & guidance for whole mankind. Before anybody else we thank our Parents who have always been a source of moral support and driving force behind whatever we do. We are indebted to our project advisor Professor Dr. Muhammad Zafar Noon for

his

worthy discussions, encouragement, inspiring guidance,

remarkable suggestions, keen interest, constructive criticism & friendly discussions which enabled us to complete this report. He spared a lot of his precious time in advising & helping us in writing this report. Without his painstaking tuition, kind patronization, sincere coaching and continuous consultation, we would not have been able to complete this arduous task successfully.

We are also grateful to Prof. Dr. A.R. Saleemi , Dr. Ing. Naveed Ramzan, Mr. Muhammad Faheem and Hafiz Zaheer Aslam for their profound gratitude and superb guidance in connection with the project. We are also thankful to librarians of National Library of Engineering Sciences and Departmental Library.

Authors

Page ii

PREFACE n-Butyraldehyde, also known as n butanal, is a colourless, flammable liquid with a characteristic aldehydic ordourm. It was discovered shortly after 1860 and was prepared by the reduction of crotonaldehyde as early as 1880. Butyraldehyde became a commercial chemical in the decade following World War II. It is used chiefly as an intermediate in the production of synthetic resins, rubbers accelerators, solvents and plasticizers. Because of large number of condensation and addition reactions it can undergo, it is useful starting material in the production of wide variety of compounds containing at least six to eight carbon atoms. N-butanal also finds its application in Pakistan for vriety of purposes. Keeping these points in mind we urged to work & we are feeling great to present our work on ―Production of n-Butanal by catalytic hydroformylation of propylene‖. This report is divided in different sections. First of all the introduction of n-butanal is given, which highlights its importance. Next are different

manufacturing

processes

for

n-butanal

production.

Detailed

description of ―Production of n-Butanal by catalytic hydroformylation of propylene‖ is presented in preceding chapter. Afterwards material and energy balance is presented. In preceding chapters introduction to different equipments of plant along with their designing procedure and specification sheets is presented. Instrumentation & Control, HAZOP Study, EIA and Cost Estimation for

this plant are also included in this report. A compact disc is also provided with report which includes soft copy of this report and HYSYS simulation of this plant and other softwares. Page iii

Table of Contents CHAPTER -1

INTRODUCTION

1

CHAPTER -2

PROCESS SELECTION

4

CHAPTER -3

CAPACITY SELECTION

9

CHAPTER -4

MATERIAL BALANCE

11

CHAPTER -5

ENERGY BALANCE

23

CHAPTER -6

DESIGN OF EQUIPMENTS

37

CHAPTER -7

INSTRUMENTATION AND CONTROL

104

CHAPTER -8

HAZOP STUDY

116

CHAPTER -9

ENVIRONMENTAL IMPACT ASSESSMENT

125

CHAPTER -10

COST ESTIMATION

133

References

138

Page iv

CHAPTER 1 INTRODUCTION CHAPTER -1

INTRODUCTION INTRODUCTION Normal-Butyraldehyde, also known as Aldehyde butyrique (French), Aldeide butirrica (Italian), Butal, Butaldehyde, Butalyde, Butanal, n-Butanal (Czech), Butanaldehyde, Butyl aldehyde, n - Butyl aldehyde, Butyral, Butyraldehyd (German) occurs naturally in small quantities. It is isolated in small quantities in the essential oils of several plants. It is also detected in oil of Lavender and Eucalyptus globules of california, in tobacco smoke, in tea leaves and in other leaves. Normal-Butyraldehyde is a colourless, flammable liquid with a characteristic aldehydic ordourm. It is used chiefly as an intermediate in the production of synthetic resins, rubbers accelerators, solvents and plasticizers. Because of large number of condensation and addition reactions it can undergo, it is useful starting material in the production of wide variety of compounds containing at least six to eight carbon atoms. Butyraldehyde became a commercial chemical in the decade following World War II. It was discovered shortly after 1860 and was prepared by the reduction of crotonaldehyde as early as 1880. Normal butyraldehyde is miscible with all common organic solvents, e.g., alcohols, ketones, aldehydes, ethers, glycols, and aromatic and aliphatic hydrocarbons, but is only sparingly soluble in water. It is an extremely flammable liquid and vapor. The vapor may cause a flash fire. N-butyraldehyde may irritate the skin and burn the eyes. Upon degradation,

peroxides are formed. Inhalation of vapors and mists may cause a narcotic effect.

Page 1

CHAPTER 1 INTRODUCTION PHYSICAL PROPERTIES Property Description

Butyraldehyde

0

-99

0

75.7

Melting Point ( C) Boiling Point ( C) 3

Density (g/cm )

0.8048

Vapour Density (Air=1)

2.48

Refractive Index (n)

1.3843

0

Flash Point ( C) 0

-9.4

Viscosity at 20 ( C)

0.433

Heat of Formation (KJ/mol)

240.3

Specific Heat (J/kg.K)

2121

Heat of Vaporization at boiling poinjt (J/g)

436

Heat of combustion (KJ/mol)

2478.7

Dipole Moment (vap.) C.m

9.07 x 10

-30

0

Surface tension (mN/m) at 24 ( C) 0

Vapour Pressure (kPa) at 20 ( C)

29.9 12.2

Page 2

CHAPTER 1 INTRODUCTION APPLICATIONS OF N-BUTANAL n-Butanal is a widely used organic compound and its consumption is approxemately 65% of whole oxo chemicals consumption.

i.

The primary use for n-butyraldehyde is as a chemical intermediate in producing other chemical commodities such as 2-Ethylhexanol (2-EH) and n-butanol.

ii.

Other products requiring n-butyraldehyde include trimethylolpropane (TMP),

n- butyric acid, polyvinyl butyral (PVB) and methyl amyl ketone. iii.

Smaller applications include intermediates for producing pharmaceuticals, crop protection agents, pesticides, synthetic resins, antioxidants, vulcanization accelerators, tanning auxiliaries, perfumery synthetics and flavors.

Page 3

CHAPTER 2 PROCESS SELECTION CHAPTER -2

PROCESS SELECTION DIFFERENT PRODUCTION ROUTS 1. Fermentation N- butyraldehyde was exclusively produced by bacterial fermentation of carbohydrate contating materials until the early 1930s. ―Pullicker industries‖ were using this process. However this technology is very old and selectivity of process is also very low. 2. Aldol Condensation The aldol route from acetaldehyde was formerly the dominant synthetic route to n- butyraldehyde.It has been shut down in favour of the more economical oxo route in 1950s. ―Celanese‖ in United States has been using this process. 3. Hydroformylation Hydroformylation which is also known as oxo synthesis was discovered in 1938 by Otto Roelen. He detected this new chemical reaction when he aimed at increasing the chain length of Fisher-Tropsch hydrocarbons by passing a mixture of

0

ethylene and synthesis gas over cobalt containing catalyst at 150 C and 100 bar in the laboratories of Ruhrchemie AG at Oberhausen, Germany. In hydroformylation olefinic double bond reacts with synthesis gas (carbon monoxide and hydrogen) in the presence of transition metal catalyst to form linear (n) and branched (b) aldehydes containing an additional carbon atom as primary products shown below. RCH2 = CH2

+ CO +

H2



RCH2CH2CHO +

RCH(CH3)CH

Starting from mid 1950s hydroformylation gained an importance. In 1997 the 6

total worldwide oxo production capacity was 6.5x10 t/year for aldehydes and Page 4

CHAPTER 2 PROCESS SELECTION alcohols. Today hydroformylation is the largest scale application of homogeneous organo-metallic catalysis.

DIFFERENT TECHNIQUES OF HYDROFORMYLATION The basic classification of Hydroformylation techniques in based on the selection of catalyst. 1. Cobalt based catalyst 2. Rhodium based catalyst The comparison of these two techniques is given in the table below. Catalyst Metal

Cobalt

Rhodium

Variant Ligand

Unmodified

Modified

Unmodified Modified Phosphines

None

Phosphines

None

Process

1

2

3

Active Catalyst

RCo(CO)4

Temperature deg. C

4

5

Hco(CO)3(L) HRh(CO)4

HRh(CO)(L)3

HRh(CO)(L)3

150-180

160-200

100-140

60-120

110-130

Pressure (bar)

200-300

50-150

200-300

10--50

40-60

Catalyst to Olefin %

0.1-1

0.6

0.0001-0.01 0.01-0.1

0.001-1

Products

Aldehydes

Alcohols

Aldehydes

Aldehydes

Aldehydes

By Products

High

High

Low

Low

Negligible

n/b ratio

80/20

88/12

50/50

92/8

43/1 – 45/1

Selectivity to Poison

No

No

No

yes

No

Process 1: BASF Process Process 2: Shell Process Process 3: Ruhrchemie Process Process 4: Union Carbide Process Process 5: RCH/RP Process

Page 5

CHAPTER 2 PROCESS SELECTION The most important of rhodium based processes on an industrial scale uses the so called phosphine modified catalyst system. The unmodified rhodium carbonyl complex is used for the reaction of special olefins. As the reaction products consist of roughly equal amount of branched and linear aldehydes, this catalyst is only applicable if both aldehyde are valuable products or if the formation of the branched aldehyde is impossible (e.g., hydroformylation of ethylene to give propanal). Up until the mid 1970s cobalt was used as catalyst metal in commercial processes (e.g., by BASF, Ruhrchemie, Kuhlmann). Because of instability of cobalt carbonyl, the reaction conditions were harsh with the pressure range of 200350 bar to avoid decomposition of the catalyst and deposition of the metallic cobalt. The ligand

modification introduced by ―Shell Researchers‖

was

significant progress in hydroformylation. The replacement of carbon monoxide with phosphines (or arsines) enhances the selectivity towards linear aldehyde (n/b) and the stability of cobalt carbonyl, leading to reduced carbon monoxide pressure. In 1974-1976 Union Carbide Corporation (UCC) and Celanese Corporation, independently of one another, introduced rhodium based catalysts on an industrial

scale. These processes combined the advantages of ligand modification with the use of rhodium as a catalyst metal. As the reaction conditions were much milder, the process was named as low-pressure oxo (LPO). Then low-pressure oxo (LPO) processes took the leading role and despite the higher price of rhodium, cobalt catalysts for the hydroformylation of propene was replaced in nearly all major plants by rhodium catalysts. Higher price of rhodium was offset by mild reaction conditions, simpler and therefore cheaper equipment, high efficiency and high yield of linear products and easy recovery of the catalyst. In addition, with respect to raw material utilization and energy conversation, the LPO processes were more advantageous than the cobalt technology, thus leading to their rapid growth. In 1980s elegant solution with respect to catalyst recovery was offered by the Ruhrchemie / Rhˆone-Poulenc (RCH/RP) process. Idea of two phase catalysis was applied to hydroformylation by using water soluble rhodium as a catalyst. Page 6

CHAPTER 2 PROCESS SELECTION Trisulfonated triphenylphosphine (TPPTS, as sodium salt) as the ligand yields the water soluble catalyst HRh(CO)(TPPTS)3. The biphasic but homogeneous reaction system exhibits distinct advantages over the conventional one phase processes. Because of mutual insolubility, the separation of the aqueous catalyst phase and reaction products, including high-boiling by-products, is achieved most simply and efficiently. However, the application of this process is limited to low molecular mass olefins which have adequate water solubility. The commercial hydroformylation of higher olefins

(C6 or larger) is performed exclusively with cobalt carbonyl catalyst.

Several approaches have been developed for the hydroformylation of high olefins: 1. Anchoring of rhodium catalyst to resins, polymeric or mineral support. 2. Homogeneous catalyst with amphiphilic complexes which can be extracted in another phase at the end of the reaction. 3. Aqueous organic biphasic catalyst involving use of particular ligands, co-solvent 4. Supported hydrophilic liquid phase or aqueous phase catalysis.

1 V-10 F-10

26

Ruhrchemie/Rhˆone-Poulenc Process

(RCH/RP)

RCH/RP process is based on a water soluble rhodium catalyst, namely HRh(CO)(TPPTS)3 complex. The use of a water soluble catalyst system brings substantial advantages for industrial practice, because the catalyst can be considered to be heterogeneous. The separation of catalyst solution and reaction products, including high-boiling by-products, is achieved most simply and efficiently. Losses of the rhodium in the crude aldehyde stream are negligible. High-boiling by-products are also negligible by using this aqueous catalyst. Purification of synthesis gas and propene is not necessary, because the catalyst is not sensitive to oxo poisons that may enter with the feed. The following figure shows the flow sheet of RCH/RP process.

Page 7

CHAPTER 2 SELECTION

PROCESS

Process Flow Diagram for RCH/RP Process for Hydroformylation of Propylene. 1 K-101

27

M -101 25 29

E-106

E-107

E-109

E-108

28 32

33 W ater

30

2 34

35

E-101 W ater

E-105 24 K-107

K-108 Water

K-109 31

K-110

Water

Water

Water 36 23

3 K-102 5

18

V-10 C-10 F-10 S-10

14 21

39

22

V-103

37

38 K-103

4 6

R-101

16 E-104 M -102

20

40 43

E-111 M-103 45 E-102 Water 7

13 R K-104 8

W ater

12 11 17 K-106 W ater

19 E-110

E-112 Steam 41

Water D-101

44

W ater E-103

9

K-105

10

Steam E-113

42

The hydroformylation plant has major four units. Propylene is compressed in compressors K101 and K-102 with an intercooler E-101 and sent to reactor R-101 for reaction. Synthesis gas is compressed in compressors K-103, K-104 and K-105 with intercoolers E-102 and E103 and sent to the stripper S-101, where it strips out the unreacted Propylene from aldehyde products coming from reactor R-101. Unreacted propylene and synthesis gas is compressed in K-106 and recycled back to reactor R-101. From reactor R-101 gases leaving contain n-butanal and iso-butanal, which are separated by several flashing after compression and cooling in compressor K-107, K-108, K-109, K-110 and in cooler E-106, E-107, E-108, E-109 respectively and mixed with n-butanal and isobutanal coming from reactor in mixer M-102. After this the mixture of n-butanal and isobutanal is heated in heat exchanger E-112. After passing through heat exchanger E112 it is sent to distillation column C-101 where n-butanal is obtained as bottom product and iso-butanal and some impurities are obtained from top of the distillation column. The condenser in distillation column is partial condenser because some gases are present in top product stream. Page 8

CHAPTER 3 CAPACITY SELECTION CHAPTER -3

CAPACITY SELECTION CAPACITY SELECTION In order to select the capacity of plant, we needed to have the knowledge of following

1. Consumption of n-Butanal in different industrial sectors of Pakistan. 2. Current production of n-Butanal in Pakistan. 3. Import of n-Butanal from different countries to Pakistan. Consumption of n-Butanal Main uses of n-Butanal in Pakistan are listed below. 1.

Production of n-Butanol by catalytic hydrogenation of n-Butanal. It is widely used as a solvent and as an esterifying agent. For example its ester with acrylic acid is used in paint, adhesive and plastic industries.

2. It is used in production of 2-Ethylexanol which is a colorless liquid and it is one of the chemical used for producing PVC plasticizers, trimethylolpropane (TMP), n-butyric acid, polyvinyl butyral (PVB), and methyl amyl ketone. 3.

Smaller applications include intermediates for producing pharmaceuticals, crop protection agents, pesticides, synthetic resins, antioxidants, vulcanization accelerators, tanning auxiliaries, perfumery synthetics, and flavors. The overall use of n-Butanal in different industries in Pakistan is estimated.

1. Paint industries 40% 2. Plastic industries

60%

Production n-Butanal in Pakistan Currently there is no plant for production of n-Butanal in Pakistan.

Page 9

CHAPTER 3 SELECTION

CAPACITY

Import of n-Butanal to Pakistan Data obtained from Lahore chamber of commerce shows that in year 2001-2002 import of n-Butanal was about 52468MTPY from countries China, . And in year 20022003 it was about 57954MTPY. Amount of n-Butanal imported in recent years according to the data obtained from Lahore chamber of commerce is listed below.

Amount of n-Butanal

Year

imported (MTPY)

Year

Amount of n-Butanal imported (MTPY)

1997-1998

32235

2000-2001

46589

1998-1999

36524

2001-2002

52468

1999-2000

41524

2002-2003

57954

A graph is potted and is extrapolated up to year 2010 as shown blow.

According to graph the amount of n-Butanal required up to 2010 is more than 100000MTPY so we selected the capacity of our plant 100000MTPY. Page 10

CHAPTER 4 MATERIAL BALANCE CHAPTER -4

MATERIAL BALANCE

Capacity of plant

= 100,000 MT/Year of 98.8% n-Butanal

Selectivity of n/iso

= 43.4/1

So total production of Butanal= 105284.4 MT/Year Production of butanal

= 14622.84 kg/hr = 202.79 kmol/hr

Production of n-butanal

= 198.2 kmol/hr = 14290.5 kg/hr

Production of i-butanal

= 4.59 kmol/hr = 331 kg/hr

Conversion is 95% 2C3H6 + 2H2 + 2CO

nC4H8O + iso C4H8O

By calculating the recycled propylene and butanal the propylene needed Propylene (99.5%) needed

= 209.7 kmol/hr = 8927 kg/hr

Syn. Gas and Propylene ratio = 2.66 Syn. Gas needed

= 536.8 kmol/hr = 8712 kg/hr

Butanal to purification plant = 733.9 kg/hr 98.8% butanal achieved

= 13889 kg/hr = 100,000 MT/year

Page 11

CHAPTER 4 MATERIAL BALANCE OVERALL Material MATERIAL BALANCE OF PLANT In

al Out Stream 1 = 8827.9 kg/hr

Total

= 17539 kg/hr

= 42 = 13888.6 kg/hr Stream 44 = 413.05 kg/hr Stream 45 = 3237.57 kg/hr 17539 kg/hr

Basis : 1 hour Process Stream number

1

5

42

44

45

Hydrogen (kg/hr)

0.00

549.04

0.00

0.00

140.27

CO (kg/hr)

0.00

8162.96

0.00

0.02

2483.30

Propylene (kg/hr)

8781.69

0.00

0.00

1.96

246.47

Propane (kg/hr)

46.24

0.00

0.00

0.66

45.23

n-butanal (kg/hr)

0.00

0.00

13722.02 275.16

291.88

I-butanal (kg/hr)

0.00

0.00

166.64

135.25

30.42

Total kg/hr

8827.91

8712

1388.6

413.05

3237.57

Material In

CHAPTER 4 Material OutBALANCE MATERIAL Stream 16 = 14364.9 kg/hr

Page 12

Total = 14364.9

kg/hr

MATERIAL BALANCE AROUND REACTOR = Stream 17 = 14329.44 kg/hr Stream 18 = 35.46 kg/hr Total

= 14364.9 kg/hr

Stream number

4

13

14

15

0.00

547.35

3.55

135.02

0.00

8142.78

80.79

2382.38

Propylene (kg/hr)

8781.69

199.84

202.49

246.59

Propane (kg/hr)

46.24

42.35

43.73

44.87

n-butanal (kg/hr)

0.00

260.44

13718.63

830.25

I-butanal (kg/hr)

0.00

8.63

315.70

25.22

Total

8827.91

9201.39

14364.90

3664.32

CO

(kg/hr)

(kg/hr) Material In

Material Out

Stream 4 = 8827.9 kg/hr

Stream 14 = 14364.9 kg/hr

Stream 13 = 9201.4 kg/hr

Stream 15 = 3664.32 kg/hr

Total

= 18029 kg/hr

=

Total

= 18029 kg/hr

Page 13

CHAPTER 4

MATERIAL BALANCE MATERIAL BALANCE AROUND FLASH SEPARATOR

Stream number 16 17 18 Hydrogen (kg/hr)

3.55 1.88 1.67

CO

(kg/hr)

80.79 48.60 32.20 Propylene (kg/hr) Page 14

202.49 CHAPTER 4 201.83 MATERIAL BALANCE 0.66 Propane (kg/hr)

MATERIAL BALANCE AROUND FLASH SEPARATOR 43.73 43.61 0.12 n-butanal (kg/hr) 13718.63 13717.85 0.79 I-butanal (kg/hr) 315.70 315.67 0.03 Total (kg/hr) Stream number 14364.90 14329.44 Hydrogen (kg/hr) 35.46

25

26

27

135.02

0.00

135.02

2382.38

0.07

2382.31

Propylene (kg/hr)

246.59

3.59

243.00

Propane (kg/hr)

44.87

0.79

44.08

n-butanal (kg/hr)

830.25

485.72

344.53

I-butanal (kg/hr)

25.22

12.35

12.87

Total

3664.32

502.52

3161.80

CO

(kg/hr)

(kg/hr)

Material In

Material Out

Stream 25 = 3664.32 kg/hr

Stream 26 = 502.52 kg/hr

Material In Total = 3664.32 Material Out

Stream 27 = 3161.8 kg/hr kg/hr

=

Stream 21 = 13839.5 kg/hr

Total

= 3664.32 kg/hr

Total = 13839.5 kg/hr

= Stream 22 = 13753.6 kg/hr

CHAPTER Stream 23 = 4 85.91 kg/hr MATERIAL BALANCE Total

= 13839.5 kg/hr

MATERIAL BALANCE AROUND FLASH SEPARATOR

Stream number 21 22 23 Hydrogen (kg/hr)

Material In Material Out 3.58 0.08 Stream 36 = 3283.17 kg/hr 3.49

Page 15

CO

(kg/hr)

Total = 3283.17 kg/hr 68.76 2.16 66.60 Propylene (kg/hr) =

1.18 Stream 1.09 37 = 3178.46 kg/hr 0.09 Stream 38 = 104.70 kg/hr Propane (kg/hr)

Page 16 Total

= 3283.17 kg/hr 0.90 CHAPTER 4 0.84 MATERIAL BALANCE 0.06 n-butanal (kg/hr)

MATERIAL BALANCE AROUND FLASH SEPARATOR 13458.03 13442.88 15.16 I-butanal (kg/hr) 307.06 306.55 0.51 Total

(kg/hr)

13839.50 13753.60 85.91

Stream number 36 37 38

Hydrogen (kg/hr) 140.18 140.16 0.02

2481.10 2480.61 0.49 Propylene (kg/hr) 243.75 240.86 2.89 Propane (kg/hr) 44.26 43.63 0.62 Page 17 n-butanal (kg/hr)

CHAPTER 4 MATERIAL BALANCE 360.47

262.65 97.82 MATERIAL BALANCE AROUND MIXER I-butanal (kg/hr) 13.41 10.55 2.85 Total

18 23 (kg/hr)

M -1 0 1

27

28

3283.17 3178.46 104.70 Stream number

18

23

27

28

Hydrogen (kg/hr)

1.67

3.49

135.02

140.18

CO (kg/hr)

32.20

66.60

2382.31

2481.10

Propylene (kg/hr)

0.66

0.09

243.00

243.75

Propane (kg/hr)

0.12

0.06

44.08

44.26

n-butanal (kg/hr)

0.79

15.16

344.53

360.47

I-butanal (kg/hr)

0.03

0.51

12.87

13.41

Total (kg/hr)

35.46

85.91

3161.80

3283.17

Material In

Material Out

Stream 18 = 35.46 kg/hr

Stream 28 = 3283.17 kg/hr

Stream 23 = 85.91 kg/hr Stream 27 = 3161.8 kg/hr Total = 3283.17 kg/hr

=

Total = 3283.17 kg/hr

Page 18

CHAPTER 4 MATERIAL BALANCE MATERIAL BALANCE AROUND MIXER

22 26 39

M -1 0 2

40

Stream number

22

26

39

40

Hydrogen (kg/hr)

0.08

0.00

0.02

0.11

CO (kg/hr)

2.16

0.07

0.49

2.72

Propylene (kg/hr)

1.09

3.59

2.89

7.57

Propane (kg/hr)

0.84

0.79

0.62

2.25

n-butanal (kg/hr)

13442.88

485.72

97.82

14026.42

I-butanal (kg/hr)

306.55

12.35

2.85

321.76

Total (kg/hr)

13753.60

502.52

104.70

14360.82

Material In

Material Out

Stream 22 = 13753.6 kg/hr

Stream 40 = 14360.82 kg/hr

Stream 26 = 502.52 kg/hr Stream 39 = 104.70 kg/hr Total = 14360.82 kg/hr

=

Total = 14360.82 kg/hr

Page 19

CHAPTER 4 MATERIAL BALANCE MATERIAL BALANCE AROUND MIXER

37 43

M -1 0 3

45

Stream number

37

43

45

Hydrogen (kg/hr)

140.16

0.11

140.27

CO (kg/hr)

2480.61

2.69

2483.30

Propylene (kg/hr)

240.86

5.62

246.47

Propane (kg/hr)

43.63

1.59

45.23

29.23

291.88

Material In n-butanal (kg/hr) 262.65 Material Out

Stream 17 = 14339.44 kg/hr

Total = 23041.44 kg/hr

=

I-butanal (kg/hr)

10.55

Stream 11(kg/hr) = 9201.94 3178.46 kg/hr Total

19.86

30.42

59.10

3237.57

Stream 19 = 13839.5 kg/hr Total = 23041.44 kg/hr Material In

Material Out

Stream 37 = 3178.46 kg/hr

Stream 45 = 3237.57 kg/hr

Stream 43 = 59.10 kg/hr Total = 3237.57 kg/hr

=

Total = 3237.57 kg/hr

Page 20

CHAPTER 4 MATERIAL BALANCE MATERIAL BALANCE AROUND STRIPPER

Stream number 17 10 11 19

Hydrogen (kg/hr) 1.88 549.04 547.35 3.58 CO (kg/hr)

48.60 8162.96 8142.80 68.76 Propylene (kg/hr) 201.83 0.00 200.65 1.18 Propane (kg/hr) 43.61 0.00 42.71 0.90 n-butanal (kg/hr) 13717.85 0.00 259.82 13458.03 I-butanal (kg/hr) Page 21

315.67

CHAPTER 0.00 4 MATERIAL 8.61 BALANCE 307.06 Total (kg/hr) MATERIAL BALANCE AROUND DISTILLATION COLUMN

41 = 14360.82 kg/hr

14339.44 8712.00 9201.94 13839.50

Stream number

41

Material0.11 In

42

43

44

0.00

0.11

0.00

14360.82 kg/hr CO (kg/hr)

2.72

0.00

2.69

0.02

Propylene (kg/hr)

7.57

0.00

5.62

1.96

Propane (kg/hr)

2.25

0.00

1.59

0.66

n-butanal (kg/hr)

14026.42

13722.02

29.23

275.16

(kg/hr) 321.76 44 = 413.06 166.64 kg/hr 19.86 135.25 kg/hr 88.66 kg/hr Stream 43I-butanal = 59.10 kg/hr Stream Total = 14360.82 Total (kg/hr)

14360.82

13888.66

59.10

413.06

Page 22

CHAPTER 5 ENERGY BALANCE CHAPTER -5

ENERGY BALANCE According to law of conservation of energy [Rate of Accumulation of Energy within system =Rate of Energy entering the system – Rate of energy leaving the system + Rate of Energy generation] For steady state system there is no accumulation of mass or energy within system. So by modifying above equation, the energy balance around all equipments is as under. For case of energy balance across each equipment to determine the enthalpy of

0

streams we used reference temperature equal to 25 C.

ENERGY BALANCE AROUND THE COMPRESSOR K-101 o

Propylene Gas P1= 101.325Kpa T1=25 C Propylene Gas P2= 2945Kpa T2=?

Inlet flow rate = 209.7 kmol/hr = 0.0583 kmol/s Inlet volumetric flowrate m

  P T 2 = T1  2  P 

Where n=0.0583 kmol/s



1



3

R=0.0821 m atm/kmol K P= 1 atm

T=298.15 K 3

V=1.356 m /s From fig 3.6 Coulson Vol. 6 for this flow rate centrifugal compressor would be used with efficiency EP=78% Page 23

CHAPTER 5 ENERGY BALANCE Outlet temperature



m

P  T2 =2 T1  P  o

Where T1=25 C P1=101.325Kpa



1



P2=2945 Kpa



??

P

m=  γ -1   = 0.137 γ =1.12 o

T2 = 200.5 C Work per kmol

 γ EP 



n −1

n  





W= 

 n -1 

P

n

Z T R

1

2

− 









 

1

1

1

Where n=



1 

= 1.16 Z1=1

 1- m  R=8.314 kJ/kmolK By putting values

W=10622 kJ/kmol

Power requirement Power =

W × kmol/h EP 1 3600

= 793 KW = 0.793MW

Similarly by putting the values in Excel Data Sheet we can calculate the power of all compressors which is given as: Compressor K-102 K-103 K-104

Power 0.129 MW 1.024 MW 0.794 MW

Compressor K-105 K-106 K-107

Power 0.694 MW 0.004 MW 0.114 MW

Compressor K-108 K-109 K-110

Power 0.289 MW 0.119 MW 0.022 MW

Page 24

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND REACTOR

Stream number

4

13

14

15

Hydrogen (kg/hr)

0.00

547.35

3.55

135.02

CO

0.00

8142.78

80.79

2382.38

Propylene (kg/hr)

8781.69

199.84

202.49

246.59

Propane (kg/hr)

46.24

42.35

43.73

44.87

n-butanal (kg/hr)

0.00

260.44

13718.63

830.25

I-butanal (kg/hr)

0.00

8.63

315.70

25.22

Total

8827.91

9201.39

14364.90

3664.32

(kg/hr)

(kg/hr)

0

Temperature C

105

45

120

120

Pressure kPa

5010

5000

5000

5000

Heat Flow kJ/hr

5.35E+06

-3.26E+07

-4.42E+07

-1.12E+07

Heat Flow In

Heat Flow Out

Stream 4 = 5.35E+06 kJ/hr

Stream 14 = -4.42E+07 kJ/hr

Stream13= -3.26E+07 kJ/hr

Stream 15 = -1.12E+07 kJ/hr

Total = -2.72E+07 kJ/hr

Total = -5.55E+07 kJ/hr

Cooling Duty Qp = -2.82E+07 kJ/hr Page 25

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND HEAT EXCHANGER E-101

Stream number

2

3

Hydrogen (kg/hr)

0.00

0.00

CO (kg/hr)

0.00

0.00

Propylene (kg/hr)

8781.69

8781.69

Propane (kg/hr)

46.24

46.24

n-butanal (kg/hr)

0.00

0.00

I-butanal (kg/hr)

0.00

0.00

Total (kg/hr)

8827.91

8827.91

Temperature C

200

77

Pressure kPa

2945

2925

Heat flow kJ/hr

7.02E+06

4.90E+06

0

Heat Flow In

Heat Flow out

Stream 2 = 7.02E+06 kJ/hr

Stream 3 = 4.90E+06 kJ/hr

Cooling Duty Qp = -2.12E+06 kJ/hr Page 26

CHAPTER 5 ENERGY BALANCE 0

0

Temperature of Cooling water in = 25 C, Temperature of Cooling water out = 30 C Mass Flow rate of cooling water

= m = Q/(∆T.Cp) = 101313.7 kg/hr

Mass Flow rate of Steam

= m = Q/λ

λ = 3957 kJ/kg. K

Similarly for the other heat exchanger in flow sheet we can calculate the heat duty and mass flow rate of water or steam need to cool or heat the process fluid with the help of spread sheet. For all these calculations we have used: Temperature of cooling water in

= 25

o

Temperature of cooling water out

= 30

o

Temperature of Steam in ()

= 120 C

Temperature of Steam out

= 120 C

C C

o

o

Heat Exchanger

Heating/Cooling Duty kJ/hr

CW/Steam Flow Rate kg/hr

E-102

-3.53E+06

168899.54

E-103

-2.65E+06

126794.26

E-104

-2.76E+06

132256.42

E-105

-9.15E+05

43786.35

E-106

-2.72E+05

13022.23

E-107

-8.15E+05

38983.59

E-108

-3.13E+05

14979.65

E-109

-2.41E+05

11553.60

E-110

-2.96E+06

141596.05

E-111

-9.12E+06

436456.74

E-112

2.75E+06

694.97

E-113

9.34E+06

2360.37

Page 27

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND VAPOR LIQUID SEPARATOR

Stream number

16

17

18

Hydrogen (kg/hr)

3.55

1.88

1.67

CO

80.79

48.60

32.20

(kg/hr)

Propylene (kg/hr)

202.49

201.83

0.66

Propane (kg/hr)

43.73

43.61

0.12

n-butanal (kg/hr)

13718.63

13717.85

0.79

I-butanal (kg/hr)

315.70

315.67

0.03

Total

14364.90

14329.44

35.46

Temperature C

40

40

40

Pressure kPa

4968

4968

4968

Heat Flow kJ/hr

-4.70E+07

-4.69E+07

-1.29E+05

(kg/hr) o

Heat Flow In

Heat Flow Out

Stream 16 = -4.70E+07 kJ/hr

Stream 17 = -4.69E+07 kJ/hr Stream 18 = -1.29E+05 kJ/hr

Total = -4.70E+07

kg/hr

=

Total

= -4.70E+07 kg/hr Page 28

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND VAPOR LIQUID SEPARATOR

Stream number

25

26

27

Hydrogen (kg/hr)

135.02

0.00

135.02

CO

2382.38

0.07

2382.31

(kg/hr)

Propylene (kg/hr)

246.59

3.59

243.00

Propane (kg/hr)

44.87

0.79

44.08

n-butanal (kg/hr)

830.25

485.72

344.53

I-butanal (kg/hr)

25.22

12.35

12.87

Total

3664.32

502.52

3161.80

Temperature C

1.43E+01

1.43E+01

1.43E+01

Pressure kPa

3.00E+02

3.00E+02

3.00E+02

Heat Flow kJ/hr

-1.21E+07

-1.68E+06

-1.05E+07

(kg/hr) o

Heat Flow In

Heat Flow Out

Stream 25 = -1.21E+07 kJ/hr

Stream 26 = -1.68E+06 kJ/hr Stream 27 = -1.05E+07 kJ/hr

Total = -1.21E+07

kJ/hr

=

Total

= -1.21E+07 kJ/hr

Page 29

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND VAPOR LIQUID SEPARATOR

Stream number

21

22

23

3.58

0.08

3.49

68.76

2.16

66.60

Propylene (kg/hr)

1.18

1.09

0.09

Propane (kg/hr)

0.90

0.84

0.06

n-butanal (kg/hr)

13458.03

13442.88

15.16

I-butanal (kg/hr)

307.06

306.55

0.51

Total

(kg/hr)

13839.50

13753.60

85.91

Temperature C

o

2.47E+01

2.47E+01

2.47E+01

Pressure kPa

3.00E+02

3.00E+02

3.00E+02

Heat Flow kJ/hr

-4.64E+07

-4.61E+07

-3.14E+05

CO

(kg/hr)

Heat Flow In

Heat Flow Out

Stream 21 = -4.64E+07 kJ/hr

Stream 22 = -4.61E+07 kJ/hr Stream 23 = -3.14E+051 kJ/hr

Total = -4.64E+07 kJ/hr

=

Total

= -4.64E+07 kJ/hr

Page 30

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND VAPOR LIQUID SEPARATOR

Stream number

36

37

38

140.18

140.16

0.02

2481.10

2480.61

0.49

Propylene (kg/hr)

243.75

240.86

2.89

Propane (kg/hr)

44.26

43.63

0.62

n-butanal (kg/hr)

360.47

262.65

97.82

I-butanal (kg/hr)

13.41

10.55

2.85

Total

3283.17

3178.46

104.70

Temperature C

80

80

80

Pressure kPa

4990

4990

4990

Heat Flow kJ/hr

-1.06E+07

-1.03E+07

-3.28E+05

CO

(kg/hr)

(kg/hr) o

Material In

Material Out

Stream 36 = -1.06E+07 kJ/hr

Stream 37 = -1.03E+07 kJ/hr Stream 38 = -3.28E+05 kJ/hr

Total

= -1.06E+07 kJ/hr

Total

= -1.06E+07 kJ/hr

Page 31

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND MIXER

18 23 27

M -1 0 1

28

Stream number

18

23

27

28

Hydrogen (kg/hr)

1.67

3.49

135.02

140.18

CO (kg/hr)

32.20

66.60

2382.31

2481.10

Propylene (kg/hr)

0.66

0.09

243.00

243.75

Propane (kg/hr)

0.12

0.06

44.08

44.26

n-butanal (kg/hr)

0.79

15.16

344.53

360.47

I-butanal (kg/hr)

0.03

0.51

12.87

13.41

Total (kg/hr)

35.46

85.91

3161.80

3283.17

Temperature C

40

25

14

15

Pressure kPa

4968

300

300

300

Heat Flow kJ/hr

-1.29E+05

-3.14E+05

-1.05E+07

-1.09E+07

o

Heat Flow In

Heat Flow Out

Stream 18 = -1.29E+05 kJ/hr

Stream 28 = -1.09E+07 kJ/hr

Stream 23 = -3.14E+05 kJ/hr Stream 27 = -1.05E+07 kJ/hr Total

= -1.09E+07 kJ/hr

Total

= -1.09E+07 kJ/hr

Page 32

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND MIXER

22 26 39

M -1 02

40

Stream number

22

26

39

40

Hydrogen (kg/hr)

0.08

0.00

0.02

0.11

CO (kg/hr)

2.16

0.07

0.49

2.72

Propylene (kg/hr)

1.09

3.59

2.89

7.57

Propane (kg/hr)

0.84

0.79

0.62

2.25

n-butanal (kg/hr)

13442.88

485.72

97.82

14026.42

I-butanal (kg/hr)

306.55

12.35

2.85

321.76

Total (kg/hr)

13753.60

502.52

104.70

14360.82

Temperature C

25

14

74

25

Pressure kPa

300

300

300

300

Heat Flow kJ/hr

-4.61E+07

-1.68E+06

-3.28E+05

-4.81E+07

o

Heat Flow In

Heat Flow Out

Stream 22 = -4.61E+07 kJ/hr

Stream 40 = -4.81E+07 kJ/hr

Stream 26 = -1.68E+06 kJ/hr Stream 39 = -3.28E+05 kJ/hr Total = -4.81E+07 kJ/hr

=

Total = -4.81E+07 kJ/hr

Page 33

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND MIXER

37

43

M -1 0 3

45

Stream number

37

43

45

Hydrogen (kg/hr)

140.16

0.11

140.27

CO (kg/hr)

2480.61

2.69

2483.30

Propylene (kg/hr)

240.86

5.62

246.47

Propane (kg/hr)

43.63

1.59

45.23

n-butanal (kg/hr)

262.65

29.23

291.88

I-butanal (kg/hr)

10.55

19.86

30.42

Total (kg/hr)

3178.46

59.10

3237.57

Temperature C

80

90

80

Pressure kPa

4990

260

260

Heat Flow kJ/hr

-1.03E+07

-1.50E+05

-1.04E+07

o

Heat Flow In

Heat Flow Out

Stream 37 = -1.03E+07 kJ/hr

Stream 45 = 3237.57 kJ/hr

Stream 43 = -1.50E+05 kJ/hr Total = 3237.57 kJ/hr

=

Total = 3237.57 kJ/hr

Page 34

CHAPTER 5 ENERGY BALANCE ENERGY BALANCE AROUND STRIPPER

Stream number

17

10

11

19

Hydrogen (kg/hr)

1.88

549.04

547.35

3.58

CO (kg/hr)

48.60

8162.96

8142.80

68.76

Propylene (kg/hr)

201.83

0.00

200.65

1.18

Propane (kg/hr)

43.61

0.00

42.71

0.90

n-butanal (kg/hr)

13717.85

0.00

259.82

13458.03

I-butanal (kg/hr)

315.67

0.00

8.61

307.06

Total (kg/hr)

14339.44

8712.00

9201.94

13839.50

40

211

44

115

4968

5000

4965

4990

-4.69E+07

-2.92E+07

-3.26E+07

-4.35E+07

o

Temperature C

Heat Flow In

Heat Flow Out

Stream 17 = -4.69E+07 kJ/hr

Stream 11 = -3.26E+07 kJ/hr

Stream 10 = -2.92E+07 kJ/hr

Stream 19 = -4.35E+07 kJ/hr

Total = -7.61E+07 kJ/hr

Total = -7.61E+07 kJ/hr Page 35

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND DISTILLATION COLUMN

Stream number

41

42

43

44

0.11

0.00

0.11

0.00

CO (kg/hr)

2.72

0.00

2.69

0.02

Propylene (kg/hr)

7.57

0.00

5.62

1.96

Propane (kg/hr)

2.25

0.00

1.59

0.66

n-butanal (kg/hr)

14026.42

13722.02

29.23

275.16

I-butanal (kg/hr)

321.76

166.64

19.86

135.25

Total (kg/hr)

14360.82

13888.66

59.10

413.06

105

112

90

90

280

300

260

260

-4.54E+07

-4.37E+07

-1.50E+05

-1.33E+06

Heat Flow In Stream 41

Heat Flow Out

= -4.54E+07 kJ/hr

Stream 42

= -4.37E+07 kJ/hr

Reboiler Duty = 9.34E+06 kJ/hr

Stream 43

= -1.50E+05 kJ/hr

Stream 44

= -1.33E+06 kJ/hr

Condenser Duty = 9.12E+06 kJ/hr Total = -3.61E+07 kJ/hr

Total = -3.61E+07 kJ/hr Page 36

CHAPTER 6

DESIGN OF

EQUIPMENTS CHAPTER -6

DESIGN OF EQUIPMENTS CHEMICAL REACTOR Reactor is the heart of a chemical plant. Chemical reactors are the vessels that are designed for a chemical reaction to occur inside them. The design of a chemical reactor deals with multiple aspects of chemical engineering. It is the job of a chemical engineer to ensure that the reaction proceeds with the highest efficiency towards the desired output product, producing the highest yield of product while requiring the least amount of money to purchase and operate. Normal operating expenses include energy input, energy removal, raw material costs, etc. energy changes can come in the form of heating or cooling, pumping to increase pressure, frictional pressure loss. However, in searching for the optimum it is not just the cost of the reactor that must be minimized. Rather, the economics of the overall process must be considered.

Reactor Selection With the variety of reactors available, some engineers believe that reactor classification is not possible. No matter how incomplete a classification may be, however, the designer needs some guidance, even though there may be some reactor types that do not fit into any classification. Accordingly, we will classify reactors using the following criteria: 1. Form of energy supplied 2. Phases in contact 3. Catalytic or noncatalytic 4. Batch or continuous

Page 37

CHAPTER 6 EQUIPMENTS

DESIGN OF

Form of Energy Supplied In hydroformylation of propylene we use thermal energy for reaction completion.

Phases in Contact The next consideration is classifying the reactors according to the phases in contact. These are: 1. gas-liquid 2. liquid-liquid 3. gas-solid 4. liquid-solid 5. gas-liquid-solid After specifying the energy form, the catalyst and the phases in contact, the next task is to decide whether to conduct the reaction in a batch or continuous mode. In the batch mode, the reactants are charged to a stirred-tank reactor (STR) and allowed to react for a specified time. After completing the reaction, the reactor is emptied to obtain the products. This operating mode is unsteady state. Other unsteady-state reactors are: (1) Continuous addition of one or more of the reactants with no product withdrawal, and (2) All the reactants added at the beginning with continuous withdrawal of product. At steady-state, reactants flow into and products flow out continuously without a change in concentration and temperature in the reactor.

Our system is gas-liquid. So for gas liquid continuous flow we can use tank reactor or tubular counter current reactor. Now we have to select either CSTR or PFR. There are two ideal models for developing reactor-sizing relationships: the plug flow and the perfectly stirred-tank models. In the plug-flow model, the reactants flowing through the reactor are continuously converted into products. During reaction there is

Residence Time

Page 38

CHAPTER 6 EQUIPMENTS

DESIGN OF

no radial variation of concentration, backmixing or forward mixing. In a perfect STR, the reactants are thoroughly mixed so that the concentration of all species and temperature are uniform throughout the reactor and equal to that leaving the reactor. 10

6

10

5

10

4

10

2

Batch Reactor

Cascade

Backmix Reactor

10

Backmix

3

10 Tubular Reactor

1

10

-1

10

-4

10

-3

10

-2

10

-1

1

10

10

2

10

3

Production Rate kg/s

From Reaction time and Production we have selected the CSTR for our process.

Page 39

CHAPTER 6 EQUIPMENTS

DESIGN OF

CSTR (Continuous Stirred-Tank Reactor) In a CSTR, one or more fluid reagents are introduced into a tank reactor equipped with an impeller while the reactor effluent is removed continuously. The impeller stirs the reagents to ensure proper mixing. The contents of the reactors are completely mixed so that the complete contents of the reactors are at the same concentration and temperature as the product stream. Since the reactor is designed for steady state, the flow rates of the inlet and outlet streams, as well as the reactors conditions, remain unchanged with time. Simply dividing the volume of the tank by the average volumetric flow rate through the tank gives the residence time, or the average amount of time a discrete quantity of reagent spends inside the tank. In short CSTR has following properties. • Mixing of reactants • Good temperature control • High heat and mass transfer efficiencies • Useful for slow reactions requiring large hold up time • Uniform composition though out the reactor

• Distribution of catalyst In our process carbon monoxide, hydrogen and propylene are converted to n-butyraldehyde in an aqueous phase containing a water soluble rhodium catalyst. The reaction, therefore, system consists of three different phases: the aqueous phase, the organic phase and the gas phase. It has been shown that mass transfer plays an important role in this reaction system. In order to transfer the gas to the reaction site and to make the separate organic phase as dispersed phase we need agitation. Keeping these points in view CSTR has been selected.

Agitation Agitation is a mean whereby mixing of phases can be accomplished and by which mass and heat transfer can be enhanced between phases or with external surfaces. In its most general sense, the process of mixing is concerned with all combinations of phases of which the most frequently occurring ones are. • Gases with gases Page 40

CHAPTER 6 EQUIPMENTS

DESIGN OF

• Gases with liquids • Gases with granular solids • Liquids into gases • Liquids with granular solids • Pastes with each other • Solids with solids The dimensions of the liquid content of a vessel and the dimensions and arrangement of impellers, baffles and other internals are factors that influence the amount of energy required for achieving the required amount of agitation or quality of mixing. The internal arrangements depend on the objectives of the operation: whether it is to maintain the homogeneity of reacting mixture or to keep a solid suspended or a gas dispersed or to enhance heat or mass transfer. A basic range of design factors, however can be defined to cover the majority of cases, for example as in figure. Gaseous Product

Liquid Product

Feed

The Vessel A dished bottom requires less power than a flat one. When a single impeller is to be used, a liquid level equal to the diameter is optimum, with the impeller located at the center for all liquid systems. Economic and manufacturing considerations, however often dictate higher ratios of depth to diameter.

Page 41

CHAPTER 6 EQUIPMENTS

DESIGN OF

Baffles Except at very high Reynolds numbers, baffles are needed to prevent vortexing and rotation of the liquid mass as a whole. A baffle width one-twelfth the tank diameter, W=D/12; a length extending from one half the impeller diameter, d/2, from the tangent line at the bottom to the liquid level, but sometimes terminated just above the level of the eye of the uppermost impeller. When solids are present or when heat transfer jacket is used, the baffles are offset from the wall a distance equal to one sixth, W/6 the baffles width. Four radial baffles at equal spacing are standard; six are only slightly more effective, and three appreciably less so. When the mixer shaft is located off center, the resulting flow pattern has fewer swirls, and baffles may not be needed, particularly at low viscosities.

Draft Tubes A draft tube is a cylindrical housing around a slightly larger in diameter than

the impeller. Its height may be little more than the diameter of the impeller or it may extend the full depth of the liquid, depending on the flow pattern that is required. Usually draft tubes are used with axial impellers to direct suction are discharge streams. An impeller draft tube system behaves as an axial flow pump of somewhat low efficiency. Its top to bottom circulation behavior is of particular value in deep tanks for suspension of solids and for dispersion of gases.

Impeller Types The typical impellers used in transitional and turbulent mixing are listed in Table 6-1. These have been divided into different general classes, based on flow patterns, applications, and special geometries. The classifications also define application types for which these impellers are used. For example, axial flow impellers are efficient for liquid blending and solids suspension, while radial flow impellers are best used for gas dispersion. Up/down impellers can be disks and plates, are considered low-shear impellers, and are commonly used in extraction columns. The pitched blade turbine, although classified as an axial flow impeller, is sometimes referred to as a mixed flow impeller, due to the flow generated in both axial and radial

Page 42

CHAPTER 6 EQUIPMENTS

DESIGN OF

directions. Above a D/T ratio of 0.55, pitched blade turbines become radial flow impellers. Flow Pattern Axial Flow Radial Flow High Shear Specialty Up/Down

Impeller Propeller, Pitched Blade Turbine, Hydrofoils Flat-blade Impeller, Disc Turbine (Rushton), Hollow-blade Turbine Cowles, Disc, Bar, Pointed blade Impeller Retreat Curve Impeller, Sweptback Impeller, Spring Impeller Disks, Plate, Circles

Impeller Size This depends on the kind of impeller and operating conditions described by the Reynolds, Froude, and Power numbers as well as individual characteristics whose effects have been correlated. For the popular turbine impeller, the ratio of diameters of impeller and vessel falls in the range, d/D = 0.3 – 0.6, the lower vales at high rpm, in

gas dispersion.

Impeller Location Expert opinions differ somewhat on this factor. As first approximation, the impeller can be placed at 1/6 the liquid level off the bottom. In some cases there is provision for changing the position of the impeller on the shaft. For off-bottom suspension of solids, am impeller location of 1/3 diameter off the bottom may be satisfactory. A rule is that a second impeller is needed when the liquid must travel more than 4 ft before deflection.

Impeller Selection For gas dispersion radial flow impellers are commonly used so from table we have selected flat blade impeller.

Modeling of mass transfer and chemical reaction The model that is used in this section takes both the mass transfer and the chemical reaction into account. The governing equations that determine the flux of the three gasses (A = H2, B = CO and E = propylene) into the aqueous liquid phase are:

Page 43

CHAPTER 6 EQUIPMENTS

DESIGN OF

From these equations the flux of the different gasses into the liquid can be calculated according to:

The average flux in time can be determined using the penetration model:

In the reactor model a constant partial pressure of the gaseous reactants was assumed and the overall loss of CO, H2 and propylene from the liquid phase is neglected. In the steady state the fluxes of all components are then equal to the total reaction rate in the solution:

The bulk concentrations of the three different reactants can be determined from this equation. Partial pressure of Propylene Concentration of Rhodium

CHAPTER 6 PE EQUIPMENTSCRh

Page 44

DESIGN OF

= 13.5 bar = 0.92 mol/m3Kinetics The kinetics of the hydroformylation reaction in the presence of a RhCl(CO) Concentration of Ligands (TPPTS)2/TPPTS complex catalyst were experimentally determined by Yang et al. CLig = 22.08 mol/m3

Conversion of Reaction XA = 95% Initial Flow rate of Propylene FA0 (2002). These authors varied the propylene concentration, the initial pressure, the = 58 mol/sec H2/CO ratio, the temperature, the rhodium concentration and the ligand to rhodium Temperature ofratio Reaction in an orthogonal experimental design to obtain the following rate expression: T = 393.15 K Reaction Pressure P in Table 6.2: The constants are defined = 50 bar Rate of Reaction Volume of Reaction Volume of Catalyst rA Vr Vcat = 0.485 mol/m3.sec = FA0 x XA/rA = 110.63 m3 = 16.38 m3

SIZING OF CSTR In sizing of CSTR first of all we should have rate expression 6.7 which, we have Page 45

already developed.

VOLUME OF REACTOR Partial pressure of Hydrogen PA Partial pressure of Carbon monoxide PB

= 17.1 bar = 19.4 bar

CHAPTER 6 EQUIPMENTS

DESIGN OF

Head Volume Volume of Reactor

VH V

LENGTH AND DIAMETER For CSTR Length to diameter ratio is 1. So L/D = 1 Since 3

2 V = (π / 4) ×L × D =127 m

Where L = Length of the reactor D = Diameter of the reactor After putting L/D = 1 calculated that

3

= 12.70 m 3 = 139.7 m

Length

L

= 5.45 m

Diameter

D

= 5.45 m

WALL THICKNESS For the calculation of wall thickness we have to calculate the total pressure which is the sum of static pressure inside the reactor. Static Pressure can be calculated as: Static pressure = Ps = ρ× g× h Putting the values and found that Ps = 940 × 9.81 × 5.45 = 50196 Pa = 50.19 kPa Pressure in the reactor

P1 = 5000 kPa

Total pressure = Pt = Ps+ P1 = 50.19 + 5000 = 5050.19 kPa Maximum allowable internal pressure = 1.1 × P = 5555 kPa For cylindrical Shells thickness of wall can be found as: t=

P × ri SE j − 0.6P

+ Cc

Page 46

CHAPTER 6 EQUIPMENTS Where t = minimum wall thickness, m P = maximum allowable internal pressure, kPa ri = inside radius of shell before corrosion allowance is added, m S = maximum allowable working stress, kPa Cc = corrosion allowance and its value is taken 3 mm

DESIGN OF

Ej = efficiency of joints expressed as a friction and its value is 0.85 Putting the values of all variable t=

5555 × 2.72 (96105.2 × 0.85) − (0.6 × 5555)

+ 0.003 t = 132.9 + 3.0 = 135.9 mm

OUTSIDE DIAMETER D0 = Di + 2t = 5.45 + 2(0.135.9) = 5.72 m

Outside Diameter

REACTOR HEAD There are three types of head: 1. Ellipsoidal head 2. Torispherical head 3. Hemispherical Heead Ellipsoidal head is used for pressure greater than 150 psig and for less than that pressure we use Torispherical head. That’s why we have selected Ellipsoidal head. Head thickness = tH = PD ∗D i 2S E j − 0.2P D

+ Cc

= 5555 x 5.45 2 x 137895 x 0.85 –(0.2 x 5555 )

+ .003

= 132.74 mm Page 47

CHAPTER 6 EQUIPMENTS AGITATOR DESIGN

DESIGN OF

Viscosity of Mixture at 393K = µ = 0.45 cp Shape Factors are S1 = D/T = 1/3 S2 = E/T = 1/3 S3 = L/D = 1/4 S4 = W/D = 1/5 S5 = J/T = 1/10 Agitator Dimensions are:

Impeller Diameter

D

= T/3

= 1.82 m

Impeller Height above Vessel floor

E

= T/3

= 1.82 m

Length of Impeller Blade

L

= 0.25D

= 0.45 m

Width of Impeller Blade

W

= D/5

= 0.36 m

CHAPTER 6 EQUIPMENTS

Page 48

DESIGN OF

Width of Baffle

J

= T/10

= 0.54 m

Length of Sparger

Ls

= T/3

= 0.36 m

For Gas-liquid-liquid mixture and reaction with heat transfer: Tip Velocity = 10 – 20 ft/sec Tip Velocity = 5 m/sec Tip Velocity = π x Da x N Form this equation we can fine speed of Impeller as: Speed of Impeller N = 5/( π x 1.82) = 53 RPM POWER CALCULATIONS Power required by the impeller is given by following equation P = NP x ρ x N3 D5 Where P = Power, watts Np = Dimensionless power number

ρ = average density, Kg/m3 N = no. of revolutions per min of impeller, RPM D = diameter of the impeller, m Power number is related with the Reynold’s number of the impeller. REYNOLD’S NUMBER: Reynold’s no. of impeller is given by following equation

Page 49

CHAPTER 6

DESIGN OF

2

EQUIPMENTS N Re =

NDa ρ µ N Re

6

= 6.04 × 10

5

For such a high Reynolds number, which is greater than 10 we use the relation for power requirement as: Power

3

5

P = KT x N x D x ρ /gc

KT from literature for six blade disc turbine = 5.75 Putting these values in above equation we get: Power

P = 7346 Watts = 9.9886 hp

Power consumption by gas sparger Gas mass flow rate

= 8828 kg/hr

Compressor efficiency

= 0.78

Pressure difference due to sparger

= 10 kPa

Gas density

= 19.8 kg/m

Power consumption by sparger

= (mG x ή x ΔP)/ρG

Power consumption by sparger

= 0.966 watts = 0.0013hp

Total Power consumption

= (0.0013+9.9886) = 9.989 hp

3

It is assumed that gear derive requires 5% of the impeller horsepower and system variations require a minimum of 10% of this impeller horsepower Thus Actual minimum motor horsepower =impeller required hp/0.85 = 9.989/0.85 = 11.75 hp

m 2

Page 50

CHAPTER 6 EQUIPMENTS

DESIGN OF

SHAFT DESIGN Continuous average rated torque on the agitator shaft, Tc = (hp x 360 x 60)/ (2 π N) = (11.75 x 360 x 60)/ (2 π x 53) = 775.5 Kg m Polar modulus of the shaft, Zp = Tm/fs Tm = 1.5 Tc 2 fs – shear stress = 550 kg/cm Zp = (1.5 x 776 x 100) /550 3 = 211.5 cm 3 πd /16 = 211.5 d = 10 cm Diameter of shaft = 10 cm Force, Fm = Tm/3.61Rb Rb – Radius of blade Fm = (1.5 x 158 x 100) / (3.61 x 45) = 711.6 Kg Maximum bending momentum M = Fm x l.3 = 701 x 1.3 = 925 Kg-m Equivalent bending moment Me = 1 [M + 2

M

1 + T2 ] 2

2

[925 + 925 2 + (776 ∗ 1.5) ] Me = 1206 kg. m

Me =

The stress due to equivalent blending F = Me/Z 3

3

Z = π d / 32 = π x 10 / 32

= 98.13

F = (1206 x 100)/98.13

= 1229 Kg/cm

2

This is within the allowable limits of stress. Overhang of agitator shaft between bearing and agitator I = 130 cm Page 51

CHAPTER 6 EQUIPMENTS

DESIGN OF 5

2

Modulus of elasticity

E = 19.5 x 10 kg/cm

Shaft deflection

ɗ = (Fm x I )/(3E x π x D /64)

3

4

ɗ = 0.54 cm

HUB AND KEY DESIGN Hub diameter of agitator = 2 x shaft diameter Length of the hub Length of key = 1.5 x shaft dia

= 20 cm = 2.5 x 36.3 = 90.82 cm = 15 cm

HEAT TRANSFER IN REACTOR Cooling Jacket area available A = π DH + πD2/4 2

= (π x 5.42 x 5.42) + (π x 5.42 /4) 2

= 153.29 m o

CW inlet temp = 28 C o

CW outlet temp = 33 C Approaches; ΔT1= 120 – 28 = 92 ΔT2= 120 – 33 = 87 0

LMTD = 89.47 C Heat, removable by jacket Q = UAΔTM

= 590 x 153.29 x 89.47 = 2.9e+7 KJ/hr This heat is Sufficient, so we can use jacket Now Cooling water Flow rate can be calculated as: Heat to be remove from reactor = 2.82 x 10

7

m = Q/( CpΔTM) = 77892 kg/hr

Page 52

CHAPTER 6 EQUIPMENTS

DESIGN OF

SPECIFICATION SHEET Identification Item

Reactor

Item Number

R-101

Number of Item

1

Operation

Continuous

Type

Continuous Stirred Tank Reactor

Design Data 3

Volume

139.71 m

Width of baffles

0.545 m

Length

5.45 m

Impeller above bottom 0.363 m

Diameter

5.45 m

Length of sparger

1.089 m

Number of Baffles

4

Speed of impeller

52.6 RPM

Type of Impeller

Disc turbine

Diameter of shaft

10 cm

Number of blades

6

Hub diameter

20 cm

Wall thickness

135.9087 mm

Length of hub

90.82 cm

Head thickness

132.7441 mm

Length of key

15 cm

Impeller Diameter

1.82 m

Power requirements

11.75 hp

Length of blade

0.45 m

Jacket area

153.29 m

Width of Blade

0.363 m

Water requirements

77891.86 kg/hr

2

Page 53

CHAPTER 6 EQUIPMENTS

DESIGN OF

HEAT EXCHANGER DESIGN A Heat Exchanger is a heat transfer device that is used for transfer of internal thermal energy between two or more fluids available at different temperatures. In most of the exchangers the fluids are separated by a heat transfer surface and ideally don’t mix with each other.

CLASSIFICATION OF HEAT EXCHANGER In general industrial heat exchangers are classified according to their: • Construction • Transfer processes • Degrees of surface compactness • Flow arrangements • Pass arrangements • Phase of the process fluid • Heat transfer mechanism

SELECTION CRITERIA FOR HEAT EXCHANGERS The selection process includes a number of factors, depending upon heat transfer

application. These are as follows: • Space

• Operating temperature

• Efficiency

• Flow rates

• Availability

• Flow arrangements

• Ease of construction.

• Intended application

• Operating pressure

• Fouling tendencies

• Material Compatibility

• Types and phases of fluids

• Material of construction

• Fabrication technique

• Operational maintenance

• Overall economy

Page 54

CHAPTER 6 EQUIPMENTS

DESIGN OF

• Thermal requirement and repair possibilities • Maintenance, inspection, cleaning, extension, • Environmental, health, and safety considerations and regulations • Performance parameters-- thermal effectiveness and pressure drops

INDUSTRIAL APPLICATIONS OF HEAT EXCHANGERS Heat exchangers are commonly used in a wide variety of industrial, chemical, and electronics processes to transfer energy and provide required heating or cooling. Industrial types of heat exchangers are common in everyday equipment such as • Boilers

• Reaction vessels.

• Cooling towers

• Furnaces

• Chillers

• Coolers

• Refrigerators

• Evaporators

• Condensers

• Dryers

• Pre heaters

• Distillation

In fact, every air conditioning system and refrigeration system has at least two heat exchangers one for the cooling side, and one to expel the heat. In the majority of chemical processes heat is either given out or absorbed, and fluids must often be either heated or cooled in a wide range of plant.

SHELL AND TUBE HEAT EXCHANGERS In process industries, shell and tube exchangers are used in great numbers, far more than any other type of exchanger. More than 90-95% of heat exchangers used in industry are of the shell and tube type. The shell and tube heat exchangers are the ―work horses‖ of industrial process heat transfer. They are the first choice because of well-established procedures for design and manufacture from a wide variety of materials, many years of satisfactory service, and availability of codes and standards for design and fabrication. They are produced in the widest variety of sizes and styles. There is virtually no limit on the operating temperature and pressure.

Page 55

CHAPTER 6 EQUIPMENTS

DESIGN OF

We employed shell and tube heat exchanger due to following reasons: • It occupies less space. • Its maintenance is easy. • Its compactness is more. • They can tolerate dirty fluids. • It is used for high heat transfer duties. • These are mostly employed in industry. • Means of directing fluid through the tubes. • Means of controlling fluid flow through the shell. • Used where large heat transfer surfaces are required • Consideration for ease of maintenance and servicing.

• Consideration for differential thermal expansion of tubes and shell. • It can be fabricated with any type of material depend up fluid properties. • They can be operated at higher temperature difference b/w coolant and gas. • Shell and Tube heat exchangers are used on applications where the demands on high temperatures and pressures are significant. Shell and tube (or tubular) heat exchangers are used in applications where high temperature and pressure demands are significant. These heat exchangers consist of a bundle of parallel sanitary tubes with the ends expanded in tube sheets. The bundle is contained in a cylindrical shell. Connections are such that the tubes can contain either the product or the media, depending upon the application. The major limitation is that they cannot be used to regenerate, but they can transfer lots of heat due to the surface area. There are many different types or designs of shell and tube heat exchangers to meet various process requirements. Shell and Tube heat exchangers can provide steady heat transfer by utilizing multiple passes of one or both fluids. Tubular heat exchangers are also employed when fluid contains particles that would block the channels of a plate heat exchanger.

Page 56

CHAPTER 6 EQUIPMENTS

DESIGN OF

DESIGN STANDARDS FOR SHELL AND TUBE HEAT EXCHANGERS There are two major standards for designing shell and tube heat exchangers: • TEMA standards • ASME Standards TEMA STANDARDS The Standards of the Tubular Exchanger Manufacturers Association (TEMA) describe these various components of shell and tube heat exchanger in detail. An STHE is divided into three parts: the front head, the shell, and the rear head. Figure illustrates the TEMA nomenclature for the various construction possibilities. Exchangers are described by the letter codes for the three sections — for example; a

BFL exchanger has a bonnet cover, a two-pass shell with a longitudinal baffle, and a fixed tube sheet rear head.

CLASSIFICATION OF SHELL AND TUBE HEAT EXCHANGERS Three principal types of shell and tube heat exchangers are: • Fixed tube-sheet exchangers • U-tube exchangers • Floating head exchanger.

GENERAL DESIGN CONSIDERATIONS The points for designing a shell and tube heat exchanger are: • Flow rates of both streams inlet and outlet temperatures of both streams. •

Operating pressure of both streams. This is required for gases, especially if the gas density is not furnished; it is not really necessary for liquids, as their properties do not vary with pressure.

•

Allowable pressure drop for both streams. This is a very important parameter for 2

heat exchanger design. Generally, for liquids, a value of 0.5–0.7 kg/cm is permitted per shell. A higher pressure drop is usually warranted for viscous liquids, especially in the tube side. For gases, the allowed value is generally 0.05– 2

2

0.2 kg/cm , with 0.1 kg/cm being typical.

Page 57

CHAPTER 6 EQUIPMENTS •

DESIGN OF

Fouling resistance for both streams. If this is not furnished, the designer should adopt values specified in the TEMA standards or based on past experience.

•

Physical properties of both streams. These include viscosity, thermal conductivity, density, and specific heat, preferably at both inlet and outlet temperatures. Viscosity data must be supplied at inlet and outlet temperatures, especially for liquids, since the variation with temperature may be considerable and is irregular (neither linear nor log-log).

•

Heat duty. The duty specified should be consistent for both the shell side and the tube side.

•

Type of heat exchanger. If not furnished, the designer can choose this based upon the characteristics of the various types of construction described earlier. In fact, the designer is normally in a better position than the process engineer to do this.

•

Line sizes. It is desirable to match nozzle sizes with line sizes to avoid expanders or reducers. However, sizing criteria for nozzles are usually more stringent than for lines, especially for the shell side inlet.

•

Nozzle sizes must sometimes be one size (or even more in exceptional circumstances) larger than the corresponding line sizes, especially for small lines.

•

Maximum shell diameter. This is based upon tube-bundle removal requirements and is limited by crane capacities. Such limitations apply only to exchangers with removable tube bundles, namely U-tube and floating-head. For fixed-tube sheet exchangers, the only limitation is the manufacturer’s fabrication capability and the availability of components such as dished ends and flanges. Thus, floating-head heat exchangers are often limited to a shell I.D. of 1.4–1.5 m and a tube length of 6 m or 9 m, whereas fixed tube sheet heat exchangers can have shells as large as 3 m and tube length up to 12 m or more.

•

Materials of construction. If the tubes and shell are made of identical materials, all components should be of this material. Thus, only the shell and tube materials of construction need to be specified. However, if the shell and tubes are of different metallurgy, the materials of all principal components should be specified to avoid any ambiguity. The principal components are shell (and shell cover), tubes, channel (and channel cover), tube sheets, and baffles. Page 58

CHAPTER 6 EQUIPMENTS

DESIGN OF

• Tube sheets may be lined or clad.

TUBE SIDE AND SHELL SIDE FLUID ALLOCATION The criteria for fluid allocation in shell and tube heat exchangers are: • Specific pressure drop. • The most corrosive to be tube side • The higher pressure fluid to be tube side. • Shell side boiling or condensation is usually preferred. Inlet Temperature

= T1 = 120 0C Outlet Temperature = T2 = 40 0C • Severe fouling fluids shall be allocated the side which is accessible.

PRELIMINARY THERMO- HYDRAULICS DESIGN STEPS Following are the Coulson’s Design Steps for shell and Tube Heat Exchanger • Defining heat-transfer rate, fluid flow-rates and temperatures. • Collect physical properties data. • Decide the type of exchanger. • Select a trial value for the overall coefficient U. • Calculate the LMTD required. • Calculate the area required. • Calculate the individual coefficients •

Calculate the overall coefficient and compare with trial value. If the calculated value differs significantly from estimated value, substitute the calculated value for estimated value.

•

Calculate the exchanger pressure drop, if unsatisfactory, change exchanger configuration.

THERMO-HYDRAULICS CALCULATIONS SHELL SIDE DATA Raw Butanal data Process Conditions

Page 59

CHAPTER 6

DESIGN OF EQUIPMENTS o

Mean temperature

= Tm

80 C

Mass Flow Rate

= mh

= 31420 kg/hr

= Cp

= 1.923 kJ/kg C

Physical Properties

Specific Heat

o

o

Thermal Conductivity = k

= 0.125 W/m C

Density

= 866 kg/m

=ρ

3

Viscosity

-3

=µ

= 0.34 x10 kg/m.s

Inlet Temperature

= t1

= 30 C

Outlet Temperature

= t2

= 37 C

Mean temperature

= tm

= 33.5 C

= Cp

= 4.23 kJ/kg C

Thermal Conductivity = k

= 0.61 W/m C

Density

=ρ

= 1015 kg/m

Viscosity

=µ

= 0.72 x10

TUBE SIDE DATA Cooling Water data o o

o

Physical Properties Specific Heat

o

o

3

-3

kg/m.s

DESIGN CALCULATIONS Calculation of Heat Duty From Energy Balance across heat exchanger E-14 we have Heat load 6

q = 2.76 x 10 kJ/hr Mass flow rate of water needed = 93300 kg/hr Calculation of LMTD Calculate the LMTD

∆T

log mean

=

∆T 1 − ∆T 2 ∆T 1 ∆T 2

ln ⁡ (

)

Page 60

CHAPTER 6 EQUIPMENTS 83− 10 80

DESIGN OF

∆T

log mean

=

Correction Factor Calculation ln ⁡( ) 10

0

= 34.5 C R

= [Th,i - Th,o] / [Tc,o- Tc,i]

R

= 11.4

P

= [Tc,o - Tc,i] / [Th,i-Tc,i]

P

= 0.08

Correction Factor F = 0.89 Corrected LMTD

0

= 30.7 C

SELECTION OF HEAT EXCHANGER Selection Criteria according to, Plant design and Economics for Chemical Engineers by Max Peter 1. Heat Duty of Exchanger 2. Mass cooling water needed

6

q

= 2.76 x 10 J/s

m

= 93300 kg/hr o

3. Log mean temperature difference (LMTD)

= 30.72 C

4. Average Value of UD

= 1245 W/m K

[

5. Area at average overall heat transfer coefficient

2

2

= 50 m

For this area : Approx. Cost of multiple-pipe heat exchanger

= $16500

Approx. Cost of U-tube Heat exchanger

= $ 9095

Approx. Cost of fixed tube heat exchanger

= $ 18190

Approx. Cost of floating head heat exchanger

= $ 55000

The most suitable from these exchangers is U-tube heat exchanger.

Assumption of overall dirt Heat Transfer Coefficient Assume: 2

Ud = 700 W/m K

Tube Specifications Standard tube specification are taken from D.Q.Kern, Tabel 10

Page 61

CHAPTER 6

DESIGN OF

EQUIPMENTS Tube side dimensions(cold fluid) BWG=14 OD=0.019m ID=0.0148m 2

Inside Surface Area/m=0.047m /m Triangular pitch= 0.0254m No. of passes=2 Tube Wall Thickness = 0.0021m

TUBE SIDE CALCULATIONS Flow Area Flow area/tube

=AC

2

= 0.00017 m

Heat Transfer Area for assumed UD Area

=A

2

= 35.64 m

Outside surface Area of Tube Outside surface Area

= Aot

= 3.14× do× L 2

= 0.292 m For this area number of tubes = A/A0t

= 122 tubes

Nearest number of tubes from literature

= 138 tubes

Corrected Heat Transfer Coefficient Udc

= 619 W/m K

Corrected Heat transfer area = AC

= 40 m

2

Mass Velocity

Velocity Gt

= mc/ (AC x no of tubes per pass) = 93300/(0.00017 x 69 x 3600)

2

2

= 2209 kg/m -s

Vt

= Gt/(density) Page 62

CHAPTER 6 EQUIPMENTS

DESIGN OF

Heat Transfer Coefficient = 2209/1015= 2.18 m/s (within the range) 2

hi from literature for water

= 10000 W/m K

hi,o

= 7787 W/m K

2

= hi x (ID/OD)

Reynolds Number Reynolds number

= Gt ID/ µ = 2209 x 0.0148/0.00072 = 45519 (turbulent flow)

SHELL SIDE CALCULATIONS Flow Area

Mass Velocity (Gs) ACS

= (ID x Pd x Lb)/Pt = (0.54 x 0.108 x 0.00635)/0.0254 2

= 0.0145 m

GS

= mh/(flow area x 3600) = (14380)/(0.0145 x 3600) 2

= 532 kg/m s Viscosity at Wall Temperature

t

Viscosity at wall temperature = µ w

= 0.00042 kg/ms

Equivalent Diameter of Shell De

= 4(0.86 x P

2

2

- 3.14 x D0 /4)/(3.14 x D0)

Reynolds Number De = 0.01805 m

Re

= (Gs x De)/ µ

= (532 x 0.01805)/0.00034 = 28270 (turbulent region) Page 63

CHAPTER 6 EQUIPMENTS

DESIGN OF

Prandtl Number Pr

= (Cp x µ)/k =3.06 x 0.00034/0.000124 = 8.33

Heat transfer coefficient h₀

= 0.36 x (k/Di) x (Re)

0.55

0.33

x (Pr)

(µ / µw )

0.14

2

= 1399.5 W/m K

Overall Clean Heat transfer coefficient UC h i ,o h o h i ,o +ho

UC

=

2

UC

= 1186 W/m K

Overall Dirt Heat transfer coefficient UD 2

RD factor from literature = 0.0006 m K/W Using this RD value and clean coefficient: 2

UD

= 693 m K/W

Check for Assumed UD Check for UD

0