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




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



Material In Total = 3664.32 Material Out

Stream 27 = 3161.8 kg/hr 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 23 = 85.91 kg/hr Stream 27 = 3161.8 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 40 = 14360.82 kg/hr

Stream 26 = 502.52 kg/hr Stream 39 = 104.70 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


=

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 43 = 59.10 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


Total = -2.72E+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



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 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


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



Total = -1.21E+07

kJ/hr

=

Total

= -1.21E+07 kJ/hr

Page 29


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




=

Total

= -4.64E+07 kJ/hr

Page 30


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



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 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


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

=


Page 33


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 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






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


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


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


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


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


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


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


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


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


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


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


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


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


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


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


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


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


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


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<(UDcal – UDas)/UDas < 30% (UDcal – UDas)/UDas x 100 = 1% So, Calculated UD is reasonable.

PRESSURE DROP CALCULATIONS

Tubes side -0.2

fi

= 0.046(Re)

= 0.0054

Gi

= 2209 kg/m sec

2

Page 64


DESIGN OF Di

= 0.014834 m

L

= 4.88 m

np

=2

ρi

= 1015 kg/m

Bi

= 1.0001

Ǿi

= 1.099

3

2

∆P=2(1.0001 x 0.0045 x 2209 x 4.88 x 2)/(1015 x 0.0148 x 1.099)

s

Shell side

Kern method ∆P

= 34 kPa ∆Ps

=(4f x G

f

= 0.05

2

Ds(NB+1))/[2ρh x De x (µ/ µw) 0.14

]

∆Ps

2

Gs

= Mass Velocity = 532 kg/m s

Ds

= 0.387 m

NB

= Number of Baffles = 44

ρh

= Density of Hot Fluid = 866 kg/m

De

= Equivalent Diameter = 0.018 m

µ

= Viscosity of Hot Fluid = 0.00034 kg/m.s

µw

= Viscosity at wall temp. = 0.00042 kg/m.s

3

2

= 4(0.05 x 532 x 0.387 x 44)/(2 x 866 x 0.018 x (0.81)

∆Ps

0.14

)

= 13.5 kPa

Page 65


DESIGN OF

SPECIFICATION SHEET Name of equipment

Shell and Tube heat exchanger

Type

1-2 pass U-tube heat exchanger

No. of equipment

1

Heat transfer area

40 m

No. of tubes

138

Type of tube

BWG 14, SS-405

Tube dimensions

ID-0.0148m, OD-0.01905m, 0.0021m thick, 4.88m long

Tube pitch & clearance

0.0254m, 0.00635m respectively

Shell dimensions

ID- 0.387 m

No. of baffles

44

ΔP tube side

34 kPa

ΔP shell side

13 kPa

2

Page 66


DESIGN OF

DESIGN OF STRIPPING COLUMN Before going in details of stripping column design first we see what is stripping and what its industrial uses are.

STRIPPING Unit operation where one or more components of a liquid stream are removed by being placed in contact with a gas stream that is insoluble in the liquid stream. OR Stripping is a physical separation process where one or more components are removed from a liquid stream by a vapor stream. In industrial applications the liquid and vapor streams can have co-current or countercurrent flows. Stripping is usually carried out in either a packed or tray column.

THEORY Stripping works on the basis of mass transfer. The idea is to make the conditions favorable for the more volatile component in the liquid phase to transfer to the vapor phase. This involves a gas-liquid interface that the more volatile component must cross.

EQUIPMENT USED FOR STRIPPING Stripping is mainly conducted in trayed towers (plate columns) and packed columns, and less often in spray towers, bubble columns and centrifugal contactors.

PLATE COLUMN Packed columns consist of a vertical column with liquid flowing in from the top and flowing out the bottom. The vapor phase enters from the bottom of the column and exits out of the top. Inside of the column are trays or plates. These trays force the liquid to flow back and forth horizontally while forcing the vapor bubbles up through holes in the trays. The purpose of these trays is to increase the amount of contact area between the liquid and vapor phases.

Page 67


DESIGN OF

PACKED COLUMN Packed columns are similar to plate columns in that the liquid and vapor flows

enter and exit in the same manner. The difference is that in packed towers there are no trays. Instead, packing is used to increase the contact area between the liquid and vapor phases. There are many different types of packing used and each one its advantages and disadvantages. The gas liquid contact in a packed bed column is continuous, not stage-wise, as in a plate column. The liquid flows down the column over the packing surface and the gas or vapor, counter-currently, up the column. In some gas-absorption columns co-current flow is used. The performance of a packed column is very much dependent on the maintenance of good liquid and gas distribution throughout the packed bed, and this is an important consideration in packed-column design.

CHOICE OF PLATE OR PACKED COLUMN The choice between a plate and packed column for a particular application can only be made with complete assurance by costing each design. However, this will not always be worthwhile or necessary, and the choice can usually be made on the basis of experience by considering main advantages and disadvantages of each type; which are listed below: 1. Plate columns can be designed to handle a wider range of liquid and gas flowrates than packed columns. 2. Packed columns are not suitable for very low liquid rates. 3. The efficiency of a plate can be predicted with more certainty than the equivalent term for packing (HETP or HTU). 4. Plate columns 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 all operating conditions, particularly in large columns. 5. It is easier to make provision for cooling in a plate column; coils can be installed on the plates.

Page 68


DESIGN OF

6. It is easier to make provision for the withdrawal of side-streams from plate

columns. 7. If the liquid causes fouling, or contains solids, it is easier to make provision for cleaning in a plate column; manways can be installed on the plates. With small diameter columns it may be cheaper to use packing and replace the packing when it becomes fouled. 8. For corrosive liquids a packed column will usually be cheaper than the equivalent plate column. 9. The liquid hold-up is appreciably lower in a packed column than a plate column. This can be important when the inventory of toxic or flammable liquids needs t be kept as small as possible for safety reasons. 10. Packed columns are more suitable for handling foaming systems. 11. The pressure drop per equilibrium stage (HETP) can be lower for packing than plates; and packing should be considered for vacuum columns. 12. Packing should always be considered for small diameter columns, say less than 0.6 m, where plates would be difficult to install, and expensive. Packed column is selected for our operation.

TYPES OF PACKING The principal requirements of a packing are that it should: • Provide a large surface area: a high interfacial area between the gas and liquid. • Have an open structure: low resistance to gas flow. • Promote uniform liquid distribution on the packing surface. • Promote uniform vapor gas flow across the column cross-section. Many diverse types and shapes of packing have been developed to satisfy these requirements. They can be divided into two broad classes: 1. Packings with a regular geometry: such as stacked rings, grids and proprietary structured packings. 2. Random packings: rings, saddles and proprietary shapes, which are dumped into the column and take up a random arrangement. Page 69


DESIGN OF

Grids have an open structure and are used for high gas rates, where low pressure drop is essential; for example, in cooling towers. Random packings and structured packing elements are more commonly used in the process industries.

RANDOM PACKING The principal types of random packings are shown

Rasching Rings

Berl Saddles

Super Intalox Saddles

Pall Rings

Intalox Saddles

Metal Hypac Page 70

CHAPTER 6

DESIGN OF

EQUIPMENTS Raschig rings are one of the oldest specially manufactured types of random packing, and are still in general use. Pall rings are essentially Raschig rings in which openings have been made by folding strips of the surface into the ring. This increases the free area and improves the liquid distribution characteristics. Berl saddles were developed to give improved liquid distribution compared to Raschig rings. Intalox saddles can be considered to be an improved type of Berl saddle; their shape makes them easier to manufacture than Berl saddles. The Hypac and Super Intalox packings shown in can be considered improved types of Pall ring and Intalox saddle respectively. Ring and saddle packings are available in a variety of materials: ceramics, metals, plastics and carbon. Metal and plastics (polypropylene) rings are more efficient than ceramic rings, as it is possible to make the walls thinner. Raschig rings are cheaper per unit volume than Pall rings or saddles but are less efficient, and the total cost of the column will usually be higher if Raschig r ings are specified. For new columns, the choice will normally be between Pall rings and Berl or Intalox saddles. The choice of material will depend on the nature of the fluids and the operating temperature. Ceramic packing will be the first choice for corrosive liquids; but ceramics are unsuitable for use with strong alkalies. Plastic packings are attacked by some organic solvents, and can only be used up to moderate temperatures. So are unsuitable for distillation columns. Where the column operation is likely to be unstable, metal rings should be used, as ceramic packing is easily broken.

PACKING SIZE In general, the largest size of packing that is suitable for the size of column should be used, up to 50 mm. Small sizes are appreciably more expensive than t he larger sizes. Above 50 mm the lower cost per cubic meter does not normally compensate for the lower mass transfer efficiency. Use of too large a size in a small column can cause poor liquid distribution.

Page 71


DESIGN OF

Recommended size ranges are: Column diameter

Use packing size

<0.3 m

<25 mm

0.3 to 0.9 m

25 to 38 mm

>0.9 m

50 to 75 mm

STRUCTURED PACKING The term structured packing refers to packing elements made up from wire mesh or perforated metal sheets. The material is folded and arranged with a regular geometry, to give a high surface area with a high void fraction. A typical example is shown below.

Structured Packing Structured packings are produced by a number of manufacturers. The basic construction and performance of the various proprietary types available are similar. The advantage of structured packings over random packing is their low HETP (typically less than 0.5 m) and low pressure drop (around 100 Pa/m). They are being increasingly used in the following applications: 1. For difficult separations, requiring many stages: such as the separation of isotopes. 2. High vacuum distillation. 3. For column revamps: to increase capacity and reduce reflux ratio requirements. The applications have mainly been in distillation, but structured packings can

also be used in absorption; in applications where high efficiency and low pressure Page 72


DESIGN OF

drop are needed. The cost of structured packings per cubic meter will be significantly higher than that of random packings, but this is offset by their higher efficiency. Selected packing is random because its cheaper and there are no difficult or vacuum separation requirements.

CHOICE OF RANDOM PACKING Factors to be considered 1. Void fraction 2. Effective surface 3. Packing size 4. Maximum operating temperature 5. Mechanical strength 6. Material selection Packing used here is 0.038m ceramic intalox saddle because 1. One of the most efficient packings 2. Little tendency to nest and block areas of bed 3. Gives a fairly uniform bed 4. Higher flooding point 5. Lower pressure drop

PACKING PROPERTIES 1.5" Nominal size

0.038mm

Packing factor F

170

Specific gravity (g/cm )

2.3

580

Water absorption (%)

<0.3

3

Package density (kg/m )

3

Free volume (%) 3

Surface area (m2/m )

80

Acid resistance (%)

>99.6

180

Max operating temp.

1100℃

Page 73


DESIGN OF

MATERIAL BALANCE Component

10

17

11

19

202

201.23

0.80

1.88

547.27

3.64

n-Butanal

13726.88

266.37

13460.5

Iso-Butanal

315.94

8.83

307.11

48.64

8141.61

69.98

44.40

43.76

0.633

14339.44

9201.94

13839.50

Propylene Hydrogen

549

CO

8163

propane 8712.00

Total

Material In

= Material Out

Stream 10 + Stream 17 = Stream 11 + Stream 19 Total = 23041.44 kg/hr = Total = 23041.44 kg/hr STRIPPER FEED (17) Mass flow rate= 14339.44kg/hr Molar flow rate= 203.53kgmol/hr Mole Fraction: Propylene:

0.023 n-Butanal:

0.936

iso-Butanal: 0.021 STRIPPED GAS (11) Mass flow rate=9201.94kg/hr Molar flow rate= 574.1kgmol/hr Mole Fraction: Propylene: 0.0083 Hydrogen: 0.476 CO:

0.506

STRIPPING GAS (10) Mass flow rate= 8712kg/hr

Molar flowrate= 566.04kgmol/hr Mole Fraction: Hydrogen: 0.484 CO:

0.516

Product (19) Mass flow rate= 13839.5 kg/hr Molar flowrate= 195.57kgmol/hr Mole Fraction: N-Butanal: 0.956 Iso-Butanal: 0.0218 Propylene:

0.000097

Page 74


DESIGN OF

PROCESS CONDITIONS Stream

Temperature (K)

Mass Flowrate (kg/hr)

Liquid Inlet

313

14339.79

Liquid Outlet

388

13842

Gas Inlet

483

8712

Gas Outlet

317

9209

Components\Mole

10

17

11

19

0.0236

0.00834

0.00009

0.00463

0.4767

0.00932

n-Butanal

0.9366

0.00644

0.9559

Iso-Butanal

0.02155

0.00021

0.02180

0.00853

0.50655

0.01278

0.00495

0.00173

0.00007

fraction Propylene Hydrogen

CO Propane

0.4849

0.5150

DESIGN APPROACH 1. Determining the diameter of column. 2. Determining the HETP of packing 3. Determining Number of transfer units for the required separation. 4. Determining the height of overall transfer units. 5. Determining the total height of column. 6. Determining the flooding velocity. 7. Verifying the pressure drop across the column. 8. Mechanical Design Page 75

CHAPTER 6 EQUIPMENTS DIAMETER OF COLUMN The column diameter is calculated by following formula

G= Mass flowrate of gas G’= Mass flux of gas � = �. � � �

� ⁡′

�..⁡

To find G’ first find the flow parameter X as followed L= Mass flow rate of liquid stream ρg = Density of gas ρl = Density of liquid x = 0.236

DESIGN OF

Pressure drop range for strippers and absorbers is 147Pa to 490Pa. Pressure drop of 294 Pa/m of a packed bed is selected. Value of gas mass flux G’ from figure 12 Chapter 1 Rule of thumbs for chemical engineers 3ed.

2

G’=0.7 kg/m s

Diameter of packed column is 0.603m. Page 76


DESIGN OF

HEIGHT EQUIVALENT OF THEORETICAL PLATE (HETP) HETP is calculated as

HETP = Where A= Size of packing

= 38mm

σ= Surface tension of liquid

= 29.2 mN/m

µ= Overall viscosity of feed stream = 0.000414 Pa s HETP = 0.0357m

NUMBER OF TRANSFER UNITS (NTU) Number of transfer units is calculated as followed.

� � � − ��

�⁡�⁡�⁡ =

+⁡ ⁡ ⁡ − ⁡⁡

⁡−⁡

⁡−⁡

��

Where β=L/HG

= 0.0045

L=Molar liquid flow rate

= 203 kmol/hr G=Molar

gas flow rate Constant

= 566 kmol/hr H=Henry’s Law = 79.52 Pa/mol fraction

x2=Solute contents in liquid inlet stream mol fraction

= 0.0083

x1=Solute contents in liquid exit stream mol fraction

= 0.00009

y1=Solute contents in gas at bottom mol fraction

=0

Ntotal= 4.5 ~ 5

Page 77


DESIGN OF

HEIGHT OF OVERALL GAS TRANSFER UNIT (HOG) Height of overall gas transfer unit is calculated as followed.

⁡

⁡

⁡ ��

−⁡ �� ⁡ ⁡

= ⁡⁡ ⁡⁡

COLUMN HEIGHT Hog = 1.45m Packing height is calculated as followed Htotal =

Hog x Ntotal

Htotal = 7.28m Giving 0.457m allowance for disengagement of vapors at top and at bottom for liquid.

Htotal = 8.194 m

FLOODING VELOCITY Flooding velocity requires the calculation of the superficial velocity that is given as Vog

= G/Aρg

Vog

= 5.88m/s

As general rule superficial velocity is 40% to 60% of the flooding velocity. Taking superficial velocity as 60% of the flooding velocity, then the flooding velocity is given as VF = 9.8m/s

CHECK FOR PRESSURE DROP For pressure drop calculation we required flow factor and gas mass velocity.

� �

Flow factor X is calculated as � � ⁡� X = 2.66 Page 78

CHAPTER 6 EQUIPMENTS Gas mass velocity is calculated with following formula.

DESIGN OF

m

Where mv = Mass flow rate of gas stream A = Area of column

G=

v A 2

G = 0.703 kg/m s Now the Y ordinate of figure 12 Chapter 1 Rule of thumbs for chemical engineers 3ed is calculated by the given formula.

⁡= ⁡′ ⁡��.� ⁡ � ⁡⁡− ⁡� ⁡⁡ Y = 0.723 Value of pressure drop for this value of Y is 294Pa/m of packing height.

MECHANICAL DESIGN THICKNESS OF SHELL Material selection: Stainless Steel 304 Shell thickness is calculated as given below

ts =Thickness of shell 2

p=Design pressure

= O.P. × 1.1 = 55.265 N/mm

D=Inside diameter

= 0.602 m

f=Design stress

= 145 N/mm

J=Joint efficiency

= 85%

c= Corrosion allowance

= 2mm

2

ts = 82mm

Page 79


DESIGN OF

SHELL WEIGHT Shell weight is calculated as Shell Weight = Volume of shell × Density of shell material Shell weight = 12670 kg

HEAD SELECTION AND THICKNESS 2:1 Elliptical head has been selected because it is used for high pressure requirements and its manufacturing is easy as compared to other types. Material of construction is low alloy steel. Thickness of elliptical head is calculated with following formula ⁡ �� ⁡⁡⁡+ ⁡. ⁡⁡

⁡� = Where th =Thickness of head

2

p =Design pressure

= O.P. × 1.1 = 55.25N/mm

Cs=Stress concentration factor

= 1.77

Rc=Crown Radius

= 0.602m

F =Design stress

= 240N/mm

J =Joint efficiency

= 85% C =

Corrosion allowance

2

=2

HEAD WEIGHT th = 83 mm Weight of elliptical head is calculated as � �⁡⁡− ⁡�

⁡ ⁡ − ⁡ � ⁡

� =⁡ ⁡

� W = 58kg

Page 80


DESIGN OF

SUPPORT DESIGN Type of support selected is skirt type support for vertical vessels. Material of construction is construction stainless steel SS-301. First we find maximum dead weight of vessel when full of water. Max. Dead weight

= 25.5 kN

Weight of column

= 202 kN

Weight of Packing

= 2.364 kN

Wind Loading

Where �� = ⁡ ⁡⁡ ⁡

2

w= Dynamic wind pressure = 2790N/m x= Length of column

= 9.11m Ms = 69813 N

Take test thickness of support say 220mm. Tensile strength of support

?

Where

�� =

�

�

�� + ⁡⁡

⁡⁡ ⁡⁡

Ms = Wind loading Ds = Inside diameter of shell ts = Thickness of support 2

σbs= 0.81 N/mm

Test compressive strength of support ⁡⁡⁡(⁡⁡ � �) ) = ⁡

⁡ ⁡

+ ⁡⁡ ⁡ ⁡ Page 81


DESIGN OF

Where W= Dead weight of column when full of water 2

σws (test) = 0.044 N/mm

Operational compressive strength of support

?

Where �

� � (�

⁡ + ⁡�

�⁡� ⁡� ��)))

=

⁡ ⁡

⁡⁡

W= Total weight of column 2

σws (operational) = 0.359 N/mm Maximum tensile strength of support

� ⁡⁡ ⁡⁡ � ⁡⁡⁡⁡��

= ⁡⁡⁡ −

⁡� � ( ��⁡⁡⁡⁡⁡⁡⁡⁡⁡) Max σs (Tensile) = 770 kPa Maximum compressive strength of support ⁡⁡⁡ ⁡⁡ ⁡⁡⁡⁡⁡⁡⁡⁡⁡�� = ⁡⁡⁡ − ⁡�� (⁡⁡⁡�) Max σs (Compressive) = 455 kPa Check for taken thickness of support Following two conditions must be satisfied. 1. ⁡� (⁡⁡⁡⁡⁡��) < ⁡⁡ � �⁡⁡⁡ ⁡⁡ Where 2

fs= Design stress

= 240N/mm

J= Joint efficiency

= 85%

θs=Base angle (normally taken as 90°)

0.0226 < 0.770 Condition 1 is satisfied.

Page 82


DESIGN OF

2.

2

E= Young Modulus of elasticity = 11.35 N/mm

0.455 < 0.518 Condition 2 is satisfied. So thickness of support = 220mm

PACKING SUPPORT The best design of packing support is one in which gas inlets are provided above the level where the liquid flows from the bed; such as the gas-injection type. These designs have a low pressure drop and no tendency to flooding. They are available in a wide range of sizes and materials: metals, ceramics and plastics.

Gas-injection type packing support

Page 83


DESIGN OF

LIQUID DISTRIBUTER The pan-type construction provides liquid level balance. Vapor passage is provided by circular gas risers as well as around the periphery of the pan.

Pan-type distributer with bottom holes

SPECIFICATION SHEET Name of equipment

Stripper

Type

Packed column

No. of equipment

1

Type of packing

0.038m ceramic Intalox saddles

Material of construction

Low alloy steel 950X

Diameter of column

0.602m

Area of column

1.138m

NTU

5

2

Hog

1.45m

Height of column

9.11m

Weight of shell

12671kg

Pressure drop

294Pa/m of packing

Page 84


DESIGN OF

DESIGN OF DISTILLATION COLUMN DISTILLATION Distillation is a separation technique in which the components of a fluid mixture are separated by heating the mixture.

BASIC PRINCIPLE Basic principle of separation is the difference in the boiling points of mixture components or their relative volatilities. Greater the difference in boiling point or greater the relative volatility, greater will be the ease of separation. It is easier to separate the components when the relative volatility is greater than one.

SIGNIFICANCE OF EQUIPMENT Distillation column is the final equipment in our process flow-sheet. The feed to distillation column is a mixture of iso-butanal, n-butanal and propylene. Whereas our required product is 98.8% n-butanal. So the required separation is being brought about by distillation column.

CHOICE B/W TRAY & PACKED COLUMN • Plate columns are designed to handle wide range of liquid flow rates without flooding. • Packed columns are not suitable for very low liquid rates. • For large column heights, weight of the packed column is more than plate

column. • Packing should always be considered for small diameter columns <0.6m where plates are difficult to install.

Page 85


DESIGN OF

• It is easier to make provisions for cooling in plate column. •

For corrosive liquids a packed column will be cheaper than the equivalent plate column.

• Packed column are suitable for foaming liquids. •

The liquid hold up is appreciably lower in a packed column than a plate column. This can be important when the inventory of toxic or flammable liquids needs to be as small as possible for safety reasons.

•

Man-holes are provided for cleaning in tray columns. In packed columns packing must be removed before cleaning.

• Plate columns can be designed with more assured rating. So on the basis of above-mentioned points, tray column was preferred.

Page 86


DESIGN OF

SELECTION OF TRAY TYPE

Types of Trays

Valve tray

Bubble Cap

tray

Sieve tray

Valve Tray

Bubble-cap Tray

Sieve Tray

Sieve Tray has been selected because of the following reasons: • High capacity • Easier and cheaper to install • Light in weight • Low pressure drop • Maintenance and cleaning is easy • Cheapest • They can easily handle wide variations in flow-rate • Simple design • Their fundamentals are well-established

Page 87


DESIGN OF

MATERIAL BALANCE Feed = Top Product + Bottom Product + Vent F = D + B + V = 192.9 + 1.00 + 5.72 = 199.62 kmol/hr Feed Components

Mass Flow Rate (kg/hr)

Molar Flow Rate (kmol/hr)

n-butanal

14028

194.8

Iso-butanl

321.8

4.471

Propylene

7.22

0.172

Hydrogen

0.1085

0.054

CO

2.71

0.096

Propane

2.02

0.045

Total

14359.8385

199.6

Top Product

Components

Mass Flow Rate(kg/hr)

Molar folw-rate

n-butanal

29.75

0.413

Iso-butanl

20.1

0.279

Propylene

5.36

0.127

Hydrogen

0.1078

0.053

CO

2.7

0.096

Propane

1.44

0.032

Total

59.4578

1.003

Page 88


DESIGN OF

Bottom Product Components

Mass Flow Rate (kg/hr)

Molar Flow Rate (kmole/hr)

n-butanal

13722

190.6

Iso-butanl

166.7

2.3

Total

13888.7

192.9

So Feed

= 199.62 kmole/hr

Top Product

= 1.00 kmole/hr

Bottom Product = 192.9 kmole/hr Vent

= 5.72 kmole/hr

Hence Feed = Top Product + Bottom Product + Vent

F = D + B + V = 192.9 + 1.00 + 5.72 = 199.62

MOLE FRACTIONS Components

Xf

Xd

XB

Hydrogen

2.72E-04

5.37E-02

0.00E+00

CO

4.85E-04

9.61E-02

0.00E+00

Propylene

8.61E-04

1.27E-01

0.00E+00

Propane

2.30E-04

3.26E-02

0.00E+00

n-butanal

0.975993149

0.411943782

0.987997

iso-butnal

2.24E-02

0.278321681

1.20E-02

Total

1.00E+00

1.00E+00

1.00E+00

Page 89


DESIGN OF

PROCESS CONDITIONS Stream

Temperature (K)

Pressure (kPa)

Feed

378

280

Top

308.2

260

Bottom

385.2

300

PHYSICAL PROPERTIES Top Conditions Avg. Mol. Weight

72 3

Vapor Density(kg/m )

4.081

3

Liquid Density (kg/m )

787.9

Surface Tension (dynes/cm)

25.04

Bottom Conditions Avg. Mol. Weight

72 3

Vapor Density (kg/m )

2.075

Liquid Density (kg/m )

3

824.8

Surface Tension (dynes/cm)

29.76

Average Viscosity of Feed = 0.3206 kg/m.s

Page 90


DESIGN OF

CALCULATION OF RELATIVE VOLATILITY Antoine’s Constants: Components

A

B

C

Hydrogen

13.6333

-164.90

-3.19

CO

14.3686

-530.22

-13.15

Propylene

15.92

-1807.53

-26.15

Propane

16.0384

-2490.48

-43.15

n-Butanal

16.1668

-2839.09

-50.15

Iso-Butanal

15.9888

-2699.107

-51.15

Antoine’s Equation (Temperature in Kelvin & Pressure in kPa) � �� = � + � ⁡+�

VAPOUR PRESSURES FROM ANTOINE’S EQUATION (MM HG) Components' Vapour Pressures

Feed

Top

Bottom

Hydrogen

536780.0802

526917.1894

541249.6

CO

406502.4511

381681.8715

418097.1

Propylene

48184.00679

38233.64194

53412.42

Propane

5436.012502

3820.372271

6357.301

n-Butanal

1820.884345

1196.5915

2193.315

Iso-Butanal

2277.564862

1524.180165

2721.262

Page 91


DESIGN OF

α- VALUES We know that

⁡=

⁡ ⁡ ⁡ ⁡ ⁡⁡ ⁡ ⁡ ⁡ ⁡⁡ ⁡ ⁡�

⁡⁡

⁡⁡⁡⁡⁡⁡ ⁡⁡⁡⁡⁡⁡⁡� ⁡⁡

Components

Feed

Top

Bottom

n-Butanal

1

1

1

Iso-Butanal

1.250801495 1.27376817

1.240707

Propylene

26.46187108 31.95212565

…

Hydrogen

294.79087

440.348431

…

CO

223.2445197 440.348431

…

Propane

2.985369453 3.192712193

…

Light key

= iso-butanal

Heavy key

= n-butanal

αtop

= 1.27

αB

= 1.24

αAvg

= 1.255

So

NATURE OF FEED Components

Xf

K-values

K Xf

N-butanal

0.9543

0.9686

0.9243

Iso-butanal

0.0453

1.2992

0.0588

Propylene

0.00021

36.9278

0.00775

Total

1.00

0.9986

Page 92


DESIGN OF

As ∑ (K Xf )

=1

q

=1

So

STANDARD DESIGN STEPS FOR DISTILLATION COLUMN • Calculation of Minimum Reflux Ratio Rm.

• Calculation of Actual reflux ratio. • Calculation of Theoretical number of stages. • Calculation of Actual number of stages. • Calculation of Diameter of the column. • Calculation of Weeping point. • Calculation of Pressure drop. • Calculation of Entrainment. • Calculation of the Height of the column

CALCULATION OF MINIMUM REFLUX RATIO Using underwood’s equation for calculation of minimum reflux ratio

As q

=1

θ

= 1.20

So by iteration:

Rmin

= 8.1

Page 93


DESIGN OF

Now by rule of thumb: ROpt

= (1.2-1.5) Rmin

ROptimum

= 12.15

So

CALCULATING MINIMUM NUMBER OF STAGES

?

Using Fenske’s equation

⁡ + �� = �

��

��

⁡�⁡

⁡

⁡⁡ � ⁡ ⁡� ⁡ �� ⁡�

Nmin

= 16.55

THEORETICAL NUMBER OF PLATES Gilliland related the number of equilibrium stages and the minimum reflux ratio and the no. of equilibrium stages with a plot that was transformed by Eduljee into the relation;

N − N min N+1

= 

0.75 1 −



 R − Rmin  R+1

0.566

   So ideal or theoretical number of stages = 17.42

PLATE EFFICIENCY �� = � � − ��...

� ⁡⁡⁡ µ�� ⁡ �⁡�



�

= 63.82%

°

ACTUAL NUMBER OF PLATES ��

�� ⁡

Nactual

�� =

⁡

= 27 plates

Page 94


DESIGN OF

LOCATION OF FEED PLATE

So feed plate is 9th from bottom.

ESTIMATION OF COLUMN HYDRAULICS We know that Feed

= 200 kmole/hr

Top Product

= 1.00 kmole/hr

Bottom Product

= 192.9 kmole/hr

R

= 12.15 So Ln

RxD

12.21

Vn

Ln+D

13.21

Lm

Ln+F

211.83

Vm

Lm-W

18.94

VAPOUR LOAD ⁡ ⁡ ⁡ .�� ⁡⁡ ⁡

⁡ ⁡⁡ �� ⁡�

=

⁡⁡⁡⁡⁡ 3

(Qv)top = 0.0647 m /sec

⁡ ⁡⁡⁡. ⁡� ⁡⁡ ⁡�

⁡⁡⁡=

� �� 3

(Qv)B = 0.1825m /sec

Page 95


DESIGN OF

LIQUID LOAD ⁡ � ⁡ ⁡⁡⁡. ⁡⁡

⁡⁡�� =

⁡⁡ ⁡ ⁡⁡⁡⁡ 3

(QL)top = 0.00031 m /sec ⁡� ⁡ ⁡⁡⁡. ⁡⁡

⁡⁡⁡ =

⁡⁡ ⁡ ⁡⁡⁡⁡ 3

(QL)B =0.00513 m /sec

LIQUID VAPOUR FACTOR

⁡�

⁡⁡�

⁡.⁡

⁡ �

⁡

⁡��

=

⁡⁡ ⁡

(FLV)top

⁡⁡⁡ = 0.0665

For this Csb = 0.0078 m/s

�

� �

�

=

� ⁡

⁡.⁡

⁡� ⁡

⁡

� ⁡

(FLV)B

= 0.5609

For this Csb = 0.0094 m/s

FLOODING VELOCITY ⁡�

⁡��

=

⁡⁡⁡

⁡⁡⁡

⁡ ⁡⁡⁡ � �

�. �

⁡⁡ −⁡⁡ ⁡

⁡.⁡

�

(Vf)top= 0.113 m/s For 80% flooding

(Vf)top= 0.0904 m/s ⁡�

⁡⁡

⁡

=

⁡⁡⁡

⁡

⁡⁡⁡

⁡

� � �. �

�

�

−⁡�

⁡ � .� �

(Vf)B= 0.202 m/s For 80% flooding

⁡

(Vf)B= 0.162 m/s

Page 96


DESIGN OF

NET VAPOUR VELOCITY ⁡�

= ⁡. ⁡⁡

⁡��

⁡�

⁡⁡⁡

NET AREA (Vn)top = 0.0678 m/s ⁡�

⁡

= ⁡. ⁡⁡ ⁡�

⁡

(Vn)B = 0.121 m/s

⁡⁡ ⁡�

⁡�

⁡⁡⁡

(An)top

2

= 0.954 m ⁡

⁡��

=

�

�

�

�

�

=

⁡� �

=

⁡

=

�

�

(An)B

2

= 1.50 m

CROSS-SECTIONAL AREA ⁡� ⁡� ⁡. � � � ��

(Ac)top

2

= 1.12 m

⁡⁡ ⁡� ⁡. ��

⁡

COLUMN DIAMETER (Ac)B

2

= 1.76 m

� = � ⁡� ⁡ Diameter at Top

= 1.19 m

Diameter at Bottom

= 1.49 m

So diameter of column is taken as 1.49 meters q

Page 97


DESIGN OF

Fractional Entrainment

FLOODING CHECK

�

�

=

�

�

�

� �

Vn = 0.135 m/s � = ⁡� ⁡ � �

F = 79.86 %

ESTIMATION OF FRACTIONAL ENTRAINMENT FACTOR

Entrainment correlation

So from graph Ψ = 0.04 As

Ψ < 0.1

for sieve plate

So process is satisfactory Page 98


DESIGN OF

ESTIMATION OF WEEP POINT

(Ūh)min = 7.85m/s Actual Minimum Vapour Velocity = (Min. Vapour Rate) / Ah Actual Minimum Vapour Velocity = 10.5 m/s As Actual Minimum Vapour Velocity > (Ūh)min

So no weeping

TOTAL ACTIVE & DOWNCOMER AREA

TOTAL PRESSURE DROP

??

�� = ⁡. ��

⁡� 2

Aa = 1.34 m

⁡� = ⁡. ��

⁡⁡

2

Ad = 0.264 m

⁡�� = ⁡. � � � �⁡�⁡ ⁡⁡ ΔPt = 663.96 Pa = 0.66396 kPa Page 99


DESIGN OF ⁡� = ⁡� + ⁡� + ⁡��

+ ⁡�

Ht = 82.05mm ⁡� = ⁡⁡ ⁡�

⁡

⁡�

�� ⁡⁡⁡

Hd = 0.3360 mm �� = ⁡

Uh =

�

1.35 m/s ⁡��

�

��

�

��⁡ ��

How = 16.563 mm

ESTIMATION OF WEIR LENGTH

= ⁡⁡⁡

Weir Length = lw = 1.56m

Page 100

CHAPTER 6 EQUIPMENTS ESTIMATION OF ORIFICE CO-EFFICIENT

DESIGN OF

Orifice Co-efficient = C⁡ = 0.84

DOWNCOMER CALCULATIONS Back-up in Downcomer ⁡⁡⁡ = ⁡⁡⁡ + ⁡⁡+ ⁡��+ ⁡⁡ Hbc = 185.46mm = 0.1854m

Page 101


DESIGN OF

Down-comer Height ⁡ ��

⁡

⁡⁡� =

⁡⁡

⁡⁡ ⁡⁡

RESIDENCE TIME Hdc = 36.83 mm = 0.03683m ⁡⁡ = (⁡⁡ ⁡⁡⁡⁡ ⁡⁡⁡ )/⁡⁡ Tr = 9.56 sec

HOLES CALCULATIONS Total Number of Holes

�� =

⁡� ⁡

⁡

nT = 6837holes Area of Single Hole

⁡⁡ =

��⁡ ⁡ 2

αh = 0.000019635m So Number of Holes/Plate = 250 holes

COLUMN HEIGHT ⁡ � =

⁡ ⁡ ⁡� − ⁡⁡⁡ +

⁡⁡ Hc = 14.840 m

HEIGHT TO DIA RATIO H / D =14.84 /1.49 =9.9

Page 102


DESIGN OF

SPECIFICATION SHEET Identification: Item

Distillation column

No. required 1 Tray type

Sieve tray

Function:

To separate n-butanal from feed mixture

Operation:

Continuous

Design Data:

Parameters

Values

Parameters

Type of Column

Tray Column

Material

Values of Carbon Steel

Construction Tray Type

Sieve Tray

Hole Diameter

5 mm

No. of Trays

28

Weir Length

1.56 m

Height of Column

14.84 m

Pressure Drop

0.663 kPa

Diameter of

1.5 m

Tray Thickness

5 mm

Tray spacing

0.45 m

Active Area

1.35 m

Flooding

80%

Reflux Ratio

12.2

Column 2

Page 103

CHAPTER 7

INSTRUMENTATION AND

CONTROL CHAPTER -7

INSTRUMENTATION AND CONTROL INSTRUMENTS Instruments are provided to monitor the key process variables during plant operation. They may be incorporated in automatic control loops or used for the manual monitoring of the process operation. They may also be part of an automatic computer data logging system. Instruments monitoring critical process variables will be fitted with automatic alarms to alert the operators to critical and hazardous situations. It is desirable that the process variable to be monitored be measured directly; often, however, this is impractical and some dependent variable that is easier to measure, is monitored in its place. For example, in the control of distillation columns the continuous on-line, analysis of the over-head product is desirable but it is difficult and expensive to achieve reliably, so temperature is often monitored as an indication of composition. The temperature instrument may form part of a control loop controlling, say, reflux flow; with the composition of the overheads checked frequently by sampling and laboratory analysis.

INSTRUMENTATION AND CONTROL OBJECTIVE The primary objective of the designer when specifying instrumentation and control schemes are: 1)

Safer Plant Operation a) To keep the process variables within known safe operating limits. b) To detect dangerous situations as they develop and to provide alarms and automatic shut-down systems. c) To provide inter locks and alarms to prevent dangerous operating procedures.

Page 104

CHAPTER 7 CONTROL 2)

INSTRUMENTATION AND

Production Rate To achieve the design product output.

3)

Product Quality To maintain the product composition within the specified quality standards

4)

Cost To operate at the lowest production cost, commensurate with the other

objectives. These are not separate objectives and must be considered together. The order in which they are listed is not meant to imply the precedence of any objective over another, other than that of putting safet y first. Product quality, production rate and the cost of production will be dependent on sales requirements. For example, it may be a better strategy to produce a better quality product at a higher cost. In a typical chemical processing plant these object ives are achieved by a combination of automatic control, manual monitoring and laboratory analysis.

COMPONENTS OF CONTROL SYSTEM Process Any operation or series of operations that produces a desired final result is a process. In this discussion the process is the n-butanal production.

Measuring Means Of all the parts of the control system the measuring element is perhaps the most important. If measurements are not made properly the remainder of the system cannot operate satisfactorily. The measured available is dozen to represent the desired condition in the process.

ANALYSIS OF MEASUREMENT VARIABLE TO BE MEASURED Measured a) Pressure measurements

b) Temperature measurements Page 105

CHAPTER 7 CONTROL

INSTRUMENTATION AND

c) Flow Rate measurements d) Level measurements

Variables to be Recorded Indicated temperature, composition, pressure, etc.

Controller The controller is the mechanism that responds to any error indicated by the error detecting mechanism. The output of the controller is some predetermined function of the error. In the controller there is also and error-detecting mechanism which compares the measured variables with the desired value of the measured variable, the difference being the error.

Final Control Element The final control element. receives the signal from the controller and by some predetermined relationships changes the energy input to the process.

CLASSIFICATION OF CONTROLLER In general the process controllers can be classified as: a) Pneumatic controllers b) Electronic controllers c) Hydraulic controllers In the n-butanal manufacturing from propylene the controller and the final control element may be pneumatically operated due to thse following reasons:

i)

The pneumatic controller is vary rugged and almost free of maintenance. The maintenance men have not had sufficient training and background in

electronics, so basically pneumatic equipment is simple.

Page 106

CHAPTER 7 CONTROL ii)

INSTRUMENTATION AND

The pneumatic controller appears to be safer in a potentially explosive atmosphere which is often present in the petro-chemical industry.

iii)

Transmission distances are short.

Pneumatic and electronic transmission

systems are generally equal upto about 250 to 300 feet. Above this distance, electronic systems begin to offer savings.

MODES OF CONTROL The various type of control are called "modes" and they determine the type of response obtained. In other words these describe the action of the controller that is the relationship of output signal to the input or error signal. It must be noted that it is error that actuates the controller. The four basic modes of control are: i)

On-off Control

ii)

Integral Control

iii)

Proportional Control

iv)

Rate or Derivative Control In industry purely integral, proportional or derivative modes seldom occur

alone in the control system. The On-off controller is the controller with very high gain. In this case the error signal at once off the valve or any other parameter upon which it sits or completely sets the system.

ALARMS AND SAFETY TRIPS AND INTERLOCKS Alarms are used to alert operators of serious and potentially hazardous, deviations in process conditions. Key instruments are fitted with switches and relays to operate audible and visual alarms on the control panels. The basic components of an automatic trip systems are:

i)

A sensor to monitor the control variable and provide an output signal when a preset valve is exceeded (the instrument).

Page 107

CHAPTER 7 CONTROL ii)

INSTRUMENTATION AND

A link to transfer the signal to the actuator usually consisting of a system of pneumatic or electric relays.

iii)

An actuator to carry out the required action; close or open a valve, switch off a motor. A safety trip can be incorporated in control loop. In this system the high-

temperature alarm operates a solenoid valve, releasing the air on the pneumatic activator closing the valve on high temperature.

Interlocks Where it is necessary to follow the fixed sequence of operations for example, during a plant start-up and shut-down, or in batch operations-inter-locks are included to prevent operators departed from the required sequence. They may be incorporated in the control system design, as pneumatic and electric relays or may be mechanical interlocks.

DIFFERENT TYPES OF CONTROLLER Flow Controllers These are used to control feed rate into a process unit. Orifice plates are by far the most common type of flow rate sensor. Normally, orifice plates are designed to give pressure drops in the range of 20 to 200 inch of water. Venturi tubes and turbine meters are also used.

Temperature Controller Thermocouples are the most commonly used temperature sensing devices. The two dissimilar wires produce a millivolt emf that varies with the "hot-junction" temperature. Iron constrictant thermocouples are commonly used over the 0 to 1300°F

temperature range.

Pressure Controller Bourdon tubes, bellows and diaphragms are used to sense pressure and differential pressure. For example, in a mechanical system the process pressure force

Page 108

CHAPTER 7 CONTROL

INSTRUMENTATION AND

is balanced by the movement of a spring. The spring position can be related to process pressure.

Level Controller Liquid levels are detected in a variety of ways. The three most common are: 1. Following the position of a float, that is lighter them the fluid. 2. Measuring the apparent weight of a heavy cylinder as it buoyed up more or less by the liquid (these are called displacement meters). 3. Measuring the difference between static pressure of two fixed elevation, one on the vapor which is above the liquid and the other under the liquid surface. The differential pressure between the two level taps is directly related to the liquid level in the vessel.

Transmitter The transmitter is the interface between the process and its control system. The job of the transmitter, is to convert the sensor signal (millivolts, mechanical movement, pressure differential, etc.) into a control signal 3 to 15 psig air-pressure signal, 1 to 5 or 10 to 50 milliampere electrical signal, etc.

Control Valves The interface with the process at the other end of the control loop is made by the final control element is an automatic control valve which throttles the flow of a stem that opens or closes an orifice opening as the stem is raised or lowered. The stem is attached to a diaphragm that is driven by changing air-pressure above the diaphragm. The force of the air pressure is opposed by a spring.

Page 109

CHAPTER 7 CONTROL

INSTRUMENTATION AND

CONTROL SCHEME ON DISTILLATION COLUMN GENERAL CONSIDERATIONS Objectives In distillation column control any of following may be the goals to achieve 1. Over head composition. 2. Bottom composition 3. Constant over head product rate.

.

4. Constant bottom product rate.

Manipulated Variables Any one or any combination of following may be the manipulated variables 1. Steam flow rate to reboiler. 2. Reflux rate. 3. Overhead product withdrawn rate. 4. Bottom product withdrawn rate 5. Water flow rate to condenser.

Loads or Disturbances Following are typical disturbances

1.

Flow rate of feed

2.

Composition of feed.

3.

Temperature of feed.

4.

Pressure drop of steam across reboiler

5.

Inlet temperature of water for condenser.

Page 110

CHAPTER 7 CONTROL

INSTRUMENTATION AND

Control Scheme Overhead product rate is fixed and any change in feed rate must be absorbed by changing bottom product rate. The change in product rate is accomplished by direct level control of the reboiler if the stream rate is fixed feed rate increases then vapor rate is approximately constant & the internal reflux flows must increase. Advantage Since an increase in feed rate increases reflux rate with vapor rate being approximately constant, purity of top product increases. Disadvantage The overhead reflux change depends on the dynamics of level control system that adjusts it.

Figure: Control scheme Page 111

CHAPTER 7 CONTROL

INSTRUMENTATION AND

CONTROL SCHEME OF CSTR GENERAL CONSIDERATION Objective In CSTR control any of following may be the goals to achieve 1. Constant Pressure inside the reactor 2. Constant Temperature inside the reactor 3. Constant Level 4. High quality of Product Reactor Variable

The independent variable for the dryer may be divided into two categories 1. Uncontrolled variables 2. Manipulated variables 3. Controlled Variables Uncontrolled Variables The variables, which cannot be controlled by controller, are called uncontrolled variables. The Uncontrolled variables include 1.Vent gases rate 2.Temperature of feed, etc. Manipulated Variables The independent manipulated inputs are variables, which are adjusted to control the chemical reaction. Any one or any combination of following may be the manipulated variables 1.Flow rate of cooling water 2.Flow rate of Feed 3.Flow rate of Product stream

Page 112

CHAPTER 7 CONTROL

INSTRUMENTATION AND

Controlled Variables Any process variable that is selected to be maintained by a control system is called a controlled variable. Following are the controlled variables 1.Inside reactor Temperature 2.Inside reactor Pressure 3.Level of reacting mixture in reactor

CONTROL SCHEME Temperature Control

The simplest method of cooling a CSTR is shown in diagram. Here we measure the reactor temperature and manipulated the flow of cooling water to the jacket. Using a jacket for cooling has two advantages. First, it minimizes the risk of leaks and thereby cross contamination between the cooling system and the process. Second, there are no internals to obstruct an agitator from providing effective mixing. A temperature sensor measure the inside reactor temperature and transfer signal to temperature transducer, transducer convert these signals in other form and the output of transducer is accepted by controller and controller transfer its signal to final control element. Final control element takes step to overcome these disturbances. Pressure Measurement Similarly as temperature controller, there is a pressure control loop, which controls the pressure inside the reactor. This controller takes action on two valves at a same time. One at the valve of feed stream and other at the valve of product stream. If pressure is high in the reactor then product stream valve will open and feed valve will close and vice versa. Level Measurement A sensor measures level of reacting materials inside the reactor and these signals are transferred to transducer and controller takes action on solvent valve. If inside level is below the required level then valve will open and vice versa.

Page 113

CHAPTER 7 CONTROL

INSTRUMENTATION AND

CSTR CONTROL CONFIGURATION

Figure: Control Scheme

Control Loop around Heat Exchanger Single Loop

Page 114

CHAPTER 7 CONTROL Double Loop

Cascade Control

INSTRUMENTATION AND

Page 115

CHAPTER 8 HAZOP STUDY CHAPTER -8

HAZOP STUDY GENERAL DESCRIPTION Background A HAZOP study identifies hazards and operability problems. The concept involves investigating how the plant might deviate from the design intent. If in the process of identifying problems during a HAZOP study, a solution becomes apparent, it is recorded as part of the HAZOP result; however, care must be taken to avoid trying to find solutions which are not apparent, beause the prime objective for HAZOP is problem identification.

The ― Guide-Word‖ HAZOP is the most well known of the HAZOP; however, several specializations of this basic method have been developed.

INDUSTRIES IN WHICH THIS TECHNIQUE IS EMPLOYED HAZOP were initially invented by ECI in the United Kindom, but the technique only started to be more widely used within the chemical process industry after the Flixborogh disaster in which a chemical plant explosion killed twenty eight people, many of which were ordinary householders living nearby. Through the general exchange of ideas and personnel, the system was then adopted by the petroleum industry, which has a similar potential for major disasters. This was then followed by the food and water industries, where the hazard potential is a great, but a different nature, the concerns being more to do with contamination rather than explosions or chemical releases.

THE REASON FOR SUCH WIDE SE OF HAZOP Safety and reliability in the design of plant initially relies upon the application of various codes of practice, or design codes and standards. Such application is usually backed up by the experience of the engineers involved, who might well have been previously concerned with the design, commissioning or operation of similar plant. Page 116

CHAPTER 8 HAZOP STUDY However, it is considered that although codes of practice are extremely valuable, it is important to supplement then with an imaginative anticipation of deviations which might occur because of, for example, equipment malfunction or operator error. In addition, most companies will admit to the fact that for a new plant, design personnel are under pressure to keep the project on schedule. This pressure always results in errors and oversights. The HAZOP Study is an opportunity to correct these before such changes become too expensive, or impossible to accomplish. Although no statistics are available to verify the claim, it is believed that the Hazop methodology is perhaps the most widely used aid to loss prevention. The reason for this can most probably be summarized as follows: • It is easy to learn • It can be easily adapted to almost all the operations that are carried out within

process industries. • No special level of academic qualification is required. One does not need to be a university graduate to participate in a study.

CONCEPT The best time to conduct a HAZOP is when the design is fairly firm. At this point, the design is well enough defined to allow meaningful answers to the questions raised in the Hazop process. Also, at this point it is still possible to change the design without a major cost. However, HAZOP can be done at any stage after design is nearly firm. The success or failure of the HAZOP depends on several factors: • The completeness and accuracy of drawing and other data used as a basis for the study. • The ability of the team to used the approach as an aid to their imagination in visualizing deviations, causes, and consequences. • The technical skills and insights of the team. • The ability of the team to concentrate on the more serious hazards which are identified. The process is systematic and it is helpful to define the terms that are used: Page 117

CHAPTER 8 HAZOP STUDY • • • • • •

Study Nodes Intention Deviations Causes Consequences Guide Words

The guide words shown in table are the ones most often used in a HAZOP Study. Guide Words

Parameter

Deviation

NO

FLOW

NO FLOW

MORE

PRESSURE

HIGH PRESSURE

AS WELL AS

ONE PHASE

TWO PHASE

OTHER THAN

OPERATION

MAINTENANCE

Guide Words

Meaning

No

Negation of the design intent

Less

Quantitative Decrease

More

Quantitative Increase

Part of

Qualitative Decrease

As Well As

Qualitative Increase

Reverse

Logical Opposite of the intent

Other Than

Complete Substitution

GUIDELINES FOR USING PROCEDURE The concepts presented above are put into practice in the following steps: • • • • •

Define the purpose, objectives, and scope of the study Select the team Prepare for the study Carry out the team review Record the results Page 118

CHAPTER 8 HAZOP STUDY Define the Purpose, Objectives, and Scope of the Study Reasons for a study might be to: • • • • • •

Check the safety of a design Decide whether and where to build Develp a list of questions to ask a supplier Check operating/ safety procedures Improve the safety of an existing facility Verify that safety instrumentation is reacting to best parameter It is also important to define what specific consequences are to be considered:

• • • • • • •

Employee safety Loss of plant or equipment Loss of production (lose competitive edge in market) Liability Insurability Public Safety Environmental Impacts

Select the Team A team might include: • • • • • • •

Design Engineer Process Engineer Operations Supervisor Instrument design Engineer Chemist Maintenance Supervisor Safety Engineer

Prepare for the Study • Obtain the necessary data. • Convert the data into a suitable form and plan the study sequence. • Arrange the necessary meetings

Page 119

CHAPTER 8 HAZOP STUDY Carry out the Team Review

Page 120

CHAPTER 8 HAZOP STUDY Record the Results The recording process is an important part of the HAZOP. It is impossible to record manually all that is said, yet it is very important that all ideas are kept. The success of this process demands a good recording scheme.

HAZOP STUDY ON THE BUTANAL PLANT

The study is carried out according to the HAZOP procedure shown in figure,

where all causes and consequences are considered. Bases on the HAZOP Study, relevant safety measurements are taken to prevent hazards, some safety precautions are discussed in the following section.

SAFETY MEASUREMENTS Safety Valves The objective of this part of the project is to select and size the necessary safety valves for the design of Butanal production plant. Page 121

CHAPTER 8 HAZOP STUDY Safety valves are pressure relieving devices, which automatically relieve a pressure system of excess pressure when abnormal operating conditions cause the pressure to exceed a set limit and re-close when the abnormal pressure recedes again below the set limit.

Types of Safety Valves The following types of safety valve are defined in the DIN 3320 standard: • • • • • • • •

Standard safety valve Full lift safety valve Direct loaded safety valve Proportional safety valve Diaphragm safety valve Bellows safety valve Controlled safety valve Foil safety valve

Safety valve Selection Factors governing the selection of specific type of safety are: •

Cost: when making cost comparisons, it is imperative to consider that capacity and nominal size of the valve. • Type of Disposal System: Valves with open bonnet for steam or non-toxic gas and valves with closed bonnet for gas/liquid applications, where escape to the atmosphere is not permitted. • Valve construction: Semi-nozzle and material of construction • Operating Characteristics: performance requirements such as overpressure and blowdown values Sizing of Safety Valves The sizing of the safety valves is done in accordance to the standard rules. The minimum flow area, Ao to discharge the required mass flow is calculated using the equations as follows. The flow coefficient Ψ is calculated using the following formulas. For subcritical pressure condition, �⁡⁡ ⁡ ⁡

>

⁡ ⁡+ ⁡

⁡ � −⁡

2

and

Ψ=

⁡ ⁡− ⁡

x

? k

−

⁡⁡ ⁡ ⁡⁡ ⁡⁡ ⁡ ⁡⁡

?? ?

o d

? ???

? k +1 k

2 k ?1

Page 122

CHAPTER 8 HAZOP STUDY For supercritical pressure conditions,

�+⁡

��

>

Ψ=

and

x

⁡ −⁡ k +1

Where Po Pa o

= absolute pressure in the pressure chamber in bar = absolute back pressure in bar

Gases and vapors Applications For gases and vapors, the smallest flow area is calculated as follows, Ao = 0.1791

qm Ψα p

T. Z M 2

Where Ao is minimum flow area in mm do

= minimum flow diameter in mm

qm

= mass flow to be discharged in kg/hr

Z

= Compressibility factor of the medium in the pressure chamber in bar

αd

= certified coefficient of discharge given by the manufacturer

Steam Applications With the aid of pressure medium coefficient x, which takes into account the properties of the out flowing steam and the conversion of the non-coherent units, the following is used for calculating the minimum flow area. x. qm

Ao = αd po For subcritical flow and pressure less than 2 bar the pressure medium coefficient x has to be calculated with the following equations. x = 0.6211 p0 . v Ψ Non Flashing Liquids For non-flashing liquids the following is applicable. A

= 0.6211 qm αd

Δp − ρ Page 123

CHAPTER 8 HAZOP STUDY 3

Where ρ is the density in kg/m

Δp is pressure difference in bars HAZOP Report Study Title : Butanal Manufacturing Plant Study No.

Guide Parameter Word

Deviation

Possible Causes

Section: Feed inlet to Reactor 1 Flow

None

No flow from Storage

Pump fails

2

Less

Less flow from Storage

Less output from Storage

3

More

More flow from Storage

Increased feed from storage

4 Temp.

More

High Recycle rate

Compressor Failure

5

Less

Side reactions

Fresh Feed compressor fails

6 Pressure

More

Reactor Failure

High flow rate of inlet gases

7

Less

No flow from reactor

Leakage from reactor

Section: Reactor to Stripper 8 Flow

None

No output from reactor

High pressure in the stripper

9 Temp.

More

Phase Change

High synthesis gas flow rate

10

less

No stripping

Compressor Failure

11 Pressure

More

No Stripping

Throttle valve failure

12

Less

No flow

High Pressure difference

Section: Stripper to Distillation Column 13 Flow

None

High Pressure in D.C

Blockage

14

More

Leakage in the D.C

High Pressure control failure

15

Less

Less pressure gradient

Stripper Leakage

16 Temp.

More

High condenser duty

High Stripping gas Temp.

17

Less

High energy cost

High syn gas Flow rate

18 Pressure

More

Feed subcooling

Blockage

19

Less

No Flow

Leakage in the pipeline

Page 124

CHAPTER 9 ASSESSMENT

ENVIRONMENTAL IMPACT

CHAPTER -9

ENVIRONMENTAL IMPACT ASSESSMENT ENVIRONMENTAL IMPACT ASSESSMENT Environmental Impact Assessment can be defined as: ―The process of identifying, predicting, evaluating and mitigating the biophysical, social, and other relevant effects of development proposals prior to major decisions be ing taken and commitments made.‖ ―An important procedure for ensuring that the likely effects of new development on the environment are fully understood and taken into account before the development is allowed to go ahead.‖

―Environmental Impact Assessment is a process, set down as a repeatable series of steps to be taken, to allow the environmental consequences of a proposed development to be assessed.‖ The environmental consequences have to be those incremental effects which are due to the proposed development and not those which are due to the passage of time or other developments not included in the proposal.

PURPOSE OF EIA An Environmental Impact Assessment (EIA) is an assessment of the likely positive and/or negative influence a project may have on the environment. The purpose of the assessment is to ensure that decision-makers consider environmental impacts before deciding whether to proceed with new projects. EIA is intended to identify the Environmental, Social and Economic impacts of a proposed development prior to decision making. This means that it is easy to identify: 1. The most environmentally suitable option at an early stage.

Page 125



2. The Best Practicable Environmental Option. 3. Alternative processes.

ORIGINS AND HISTORY OF EIA First formal system of EIA established in the US following the National Environmental Policy Act (NEPA) of 1969 NEPA sought to ensure that environmental concerns were considered in the decisionmaking of Federal Government agencies It is required for agencies to prepare a detailed statement on the environmental impact of ―proposals for legislation and other major Federal actions significantly affecting the quality of the human environment‖

STEPS INVOLVED IN EIA

Page 126



POTENTIAL HEALTH EFFECTS Eye contact: Immediately flush eyes or skin with copious amounts of water for at least 15 minutes. Assure adequate flushing of the eyes by separating the eyelids with fingers, and seek medical advice. Skin contact: Immediately flush skin with copious amounts of water for at least 15 minutes while removing contaminated clothing and shoes. Inhalation:

If inhaled, remove to fresh air. If not breathing give artificial respiration. If breathing is difficult, give oxygen. If swallowed: Wash out mouth with water provided the person is conscious. Call a physician. Wash contaminated clothing before reuse. Chronic Exposure: Marked impairment of vision has been reported. Repeated or prolonged exposure may cause skin irritation. Aggravation of Pre-existing Conditions: Persons with pre-existing skin disorders or eye problems or impaired liver or kidney function may be more susceptible to the effects of the substance.

FIRST AID MEASURES Inhalation: Remove to fresh air. If not breathing, give artificial respiration. If breathing is difficult, give oxygen. Get medical attention immediately. Ingestion: Induce vomiting immediately as directed by medical personnel. Never give anything by mouth to an unconscious person. Get medical attention immediately.

Page 127



Skin Contact: Immediately flush skin with plenty of water for at least 15 minutes while removing contaminated clothing and shoes. Get medical attention. Wash clothing before reuse. Thoroughly clean shoes before reuse. Eye Contact: Immediately flush eyes with plenty of water for at least 15 minutes, lifting lower and upper eyelids occasionally. Get medical attention immediately.

FIRE FIGHTING MEASURES

Fire: Flash point: 262 K Auto ignition temperature: 503 K Flammable limits in air 1.4% - 12.5% by volume. Flammable Liquid and Vapor Flammable vapor and air can form explosive mixtures. Case fire, high-heat combustion caused the explosion. With the oxidizer reaction will be fierce. In case of high fever, can occur polymerization reaction and emit heat and large containers ruptured and caused the explosion. Its vapor is heavier than air, will spread to the lower Department of considerable local and met the fire will generate a return to burning. Combustion (decomposition) products: carbon monoxide, carbon dioxide. Extinguishing media • Carbon dioxide, dry chemical powder or appropriate foam • Water may be effective for cooling, but may not effect extinguishment.. Special fire-fighting procedures • Wear self-contained breathing apparatus and protective clothing to prevent contact with skin and eyes. • Extremely flammable. • Use water spray to cool fire-exposed containers.

Page 128



Unusual fire and explosion hazards • Vapor may travel considerable distance to source of ignition and flash back, • Container explosion may occur under fire conditions. • Forms explosive mixtures in air.

ACCIDENTAL RELEASE MEASURES • Evacuate personnel to safe areas.

• Keep people away from and upwind of spill/leak. • Shut off all sources of ignition. • Wear self-contained breathing apparatus, rubber boots and heavy rubber gloves. • Cover with an activated carbon adsorbent take up and place in closed containers. Transport outdoors. • Ventilated and wash spill site after material pickup is complete.

HANDLING AND STORAGE Handling: Use only in well-ventilated areas. Do not breathe vapors or spray mist. Avoid contact with skin, eyes and clothing. Take necessary action to avoid static electricity discharge (which might cause ignition of organic vapors.) Storage: Keep tightly closed in a dry, cool and well-ventilated place. Keep away from heat and sources of ignition. Store in original container. Electrical equipment should be protected to the appropriate standard.

EXPOSURE CONTROLS/PERSONAL PROTECTION • Wear approved respirator EN141, chemical resistant gloves, safety goggles, other protective clothing. • Use only in a chemical fume cupboard. • Safety shower and eye bath. • Use non-sparking tools. • Do not breathe vapor. • Do not get in eyes, on skin, on clothing. • Avoid prolonged or repeated exposure. Page 129



• Wash thoroughly after handling. • Keep tightly closed. • Keep away form heat, sparks and open flame. •

Store in a cool dry place.

STABILITY AND REACTIVITY Stability: • Stable. Conditions to avoid: • Heat • Protect from light. Incompatibilities: • Oxidizing agents. • Strong bases. • Strong reducing agents. • Strong acids. Hazardous combustion or decomposition products: • Toxic fumes of carbon monoxide, carbon dioxide. Hazardous polymerization: • May undergo autopolymerisation.

TOXICOLOGICAL INFORMATION Acute effects: • Harmful if swallowed, inhaled or absorbed through skin. Page 130



• Butanal is extremely destructive to tissue of the mucous membranes and upper respiratory tract, eyes and skin. • Inhalation may result in spasm, inflammation and oedema of the larynx and bronchi, chemical pneumonitis and pulmonary oedema.

• Symptoms of exposure may include burning sensation, coughing, wheezing, and laryngitis, shortness of breath, headache, nausea and vomiting. • Prolonged or repeated exposure may cause allergic reactions in certain sensitive individuals. Toxicity data: • orl-rat LD50:2490 mg/kg • ipr-rat LD50:800 mg/kg • scu-rat LD50:10 gm/kg

ECOLOGICAL INFORMATION Data not available yet.

DISPOSAL CONSIDERATIONS • Burn in a chemical incinerator equipped with an afterburner and scrubber but exert extra care in igniting as this material is highly flammable. •

Observe all EU, Irish and Local Environmental Regulations.

TRANSPORT INFORMATION Contact chemical supplier for transportation information.

REGULATORY INFORMATION Classification according to European directive on classification of hazardous preparations 90/492/EEC Contains: Butyraldehyde Symbol(s): 1. Highly flammable 2. Harmful 3. Corrosive Page 131

CHAPTER 9 ASSESSMENT Risk – phrase(s) • R11 – Highly flammable.


• R20 – Toxic by inhalation and if swallowed. • R21 – Harmful in contact with skin • R22 – Harmful if swallowed • R34 – Causes burns Safety – phrase(s) • S2 – Keep locked-up and out of reach of children. • S9 – Keep container in a well-ventilated place

• S29 – Do not empty into drains • S33 – Take precautionary measures against static discharges Recommended restrictions: Take notice of labels and material safety data sheets for the working chemicals. Take necessary action to avoid static electricity discharge (which might cause ignition of organic vapors.)

PERSONAL PROTECTION EQUIPMENT • Respiratory protection: Incase of insufficient ventilation wear suitable respiratory equipment. • Hand protection: Neoprene gloves / butylrubber gloves. • Eye protection: Goggles giving complete protection to eyes. • Skin and body protection: Rubber or plastic boots, Chemical resistant apron / complete suit protecting against chemicals.

Page 132

CHAPTER 10 ESTIMATION

COST

CHAPTER -10

COST ESTIMATION PURCHASED EQUIPMENT COST Marshall and Swift Equipment Cost Index Cost index in 2002

=

1116.9

Cost index in 2003

=

1184.6

Cost index in 2004

=

1265.2

Cost index in 2005

=

1350.9

Cost index in 2006

=

1405.4

Cost index in 2007

=

1452.3

Cost index in 2008

=

1491.7

Cost index in 2009

=

1532.4

Present Cost = Orignal Cost Index value at present Index value at time orignal cost was obtained

REACTOR 3

Capacity

= 140 m

Material of construction

= Stainless Steel

Cost of reactor in 2002

= $ 1250000

Cost of reactor in 2009

= 1250000 x (1532.4/1116.9) = $ 171415

Page 133


COST

HEAT EXCHANGER Type

= U-Tube

Area

= 40 m

Cost of Heat Exchanger in 2002

= $ 18864

Cost of Heat Exchanger in 2009

= 18864 x (1532.4/1116.9)

2

STRIPPER = $ 25873 Type

= Packed Column

Total Height

= 9.11 m

Inside Diameter

= 0.604 m

Cost of Stripper in 2002

= $ 82300

Cost of Stripper in 2009

= 82300 x (1532.4/1116.9)

DISTILLATION COLUMN = $ 122500 Type

= Tray Column

Diameter of column

= 1.5 m

Type of Tray

= Sieve Tray

No. of Trays

= 28

Cost of Distillation column in 2002 = $ 110000 Cost of Distillation column in 2009 = 110000 x (1532.4/1116.9) = $ 150870 Cost of Sieve Trays in 2002

= $ 3100

Cost of Sieve Trays in 2009

= 3100 x (1532.4/1116.9) = $ 4115

Total Cost of Distillation Column

=$ 154984 Page 134


COST

COMPRESSOR hp of compressor

= 1005hp

Purchased Cost of compressor in 2007

= $ 450200

Purchased cost in 2009

= 450200 x (1532.4/1452.3) = $ 475030

TOTAL PURCHASED COST IN 2009 Equipment

Reactor 171414 Heat Exchanger E-101 19716 Heat Exchanger E-102

Purchased Cost ($)

32829 Heat Exchanger E-103 24645 Heat Exchanger E-104 25873 Heat Exchanger E-105 8510 Heat Exchanger E-106 2510 Heat Exchanger E-107 7580 Heat Exchanger E-108 2910 Heat Exchanger E-109 2241 Heat Exchanger E-110 27530 Heat Exchanger E-111 84816 Heat Exchanger E-112 25575 Heat Exchanger E-113 86862

CHAPTER 10 Stripper

COST 122500

ESTIMATION

Distillation Column Compressor154984 k-101 578100 Compressor k-102 475030 Compressor k-103 465900 Compressor k-104 471900

Page 135

107600 Compressor k-106 99900 Compressor k-107 209800 Compressor k-108

Page 136

103000

CHAPTER 10

COST ESTIMATION

Compressor k-109

Indirect Costs 26400 Engineering and Supervision Compressor k-110

$ 1103727

Contraction Expenses 6500

$ 1371298

Legal Expanses

$ 133785

Total Contractor’s PurchasedFee Equipment Cost Contingency = $ 3344630

$ 735818 $ 1471637

CAPITAL INVESTMENT ESTIMATION Total Indirect Plant Cost $ 4816265 Direct Costs Fixed Capital Investment Working Capital Purchased Equipment Delivered Cost $ 3344630 Total Capital Investment Purchased Equipment Installation $ 1571976 Instrumentation and Control $ 1204066 Piping $ 2274348 Electrical System $ 367909

$ 16856931 $ 2528540 $ 19385471

$ 602033 Yard Improvement $ 334463 Service Facilities $ 2341241 Total Direct Plant Cost Page 137

$ 12040666

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4th Year Project

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