4th Year Project

(1)

Plant Design for Production

of

n-Butyraldehyde

by

(2)

Session: 2005-2009

P

r o je

c t A

d v is

o r s

Prof. Dr. Muhammad Zafar Noon

Mr. Muhammad Faheem P

r o je

c t

M

e m

b ers

Hafiz Sajid Sattar 2005-Chem-62

Muhammad Waqas 2005-Chem-86

Saeed Ur Rehman 2005-Chem-98

Saad Ullah Mirza 2005-Chem-74

D

E

P

AR

T

M

E

N

T

OF C

H

E

MI

C

A

L

E

N

GI

N

EE

R

I N

G

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:

(3)

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!

(4)

Page i

ACKNOW LEDGEMENT

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.

(5)

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.

(6)

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

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

(7)

Page iv

CHAPTER 1 INTRODUCTION

CHAPTER -1

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.

(8)

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

Page 1

CHAPTER 1 INTRODUCTION

PHYSICAL PROPERTIES

Property Description Butyraldehyde Melting Point (0C) -99

Boiling Point (0C) 75.7 Density (g/cm3) 0.8048

Vapour Density (Air=1) 2.48 Refractive Index (n) 1.3843 Flash Point (0C) -9.4 Viscosity at 20 (0C) 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

(9)

Surface tension (mN/m) at 24 (0C) 29.9 Vapour Pressure (kPa) at 20 (0C) 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.

(10)

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

(11)

ethylene and synthesis gas over cobalt containing catalyst at 150 0C 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 total worldwide oxo production capacity was 6.5x106 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

None Modified Phosphines Unmodified None Modified Phosphines Process 1 2 3 4 5

Active Catalyst RCo(CO)4 Hco(CO)3(L) HRh(CO)4 HRh(CO)(L)3 HRh(CO)(L)3

Temperature deg. C 150-180 160-200 100-140 60-120 110-130

Pressure (bar) 200-300 50-150 200-300 10--50 40-60

(12)

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 200-350 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

(13)

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.

(14)

F -1 0 1 F -1 0 4 F -1 0 3 V -1 0 1 26₁5

Ruhrchemie/Rhˆone-Poulenc

(RCH/RP)

Process

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 PROCESS

SELECTION

Process Flow Diagram for RCH/RP Process for Hydroformylation of Propylene. 1 K-101 27 M-101 25 29

E-106 E-107 E-108 E-109

28 30 2

32 33 34 35

W ater E-101 _{W ater}

E-105 24 K-107 K-108 K-109 K-110 Water 31 Water Water Water 36 3 23 K-102 18 5

(15)

V -1 0 2 F -1 0 2 S -1 0 1 C -1 0 1 14 21 22 39 V-103 38 37 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

(16)

K-105

10 Steam

E-113

42

The hydroformylation plant has major four units. Propylene is compressed in compressors

K-101 and K-102 with an intercooler E-K-101 and sent to reactor R-K-101 for reaction. Synthesis gas is compressed in compressors K-103, K-104 and K-105 with intercoolers 102 and E-103 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 butanal coming from reactor in mixer M-102. After this the mixture of n-butanal and iso-butanal is heated in heat exchanger 112. After passing through heat exchanger E-112 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

(17)

Material In Material Out Stream 1 = 8827.9 kg/hr Stream 5 = 8712 kg/hr Total = 17539 kg/hr =

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

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 CAPACITY

SELECTION

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 2002-2003 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.

(18)

Stream 42 = 13888.6 kg/hr Stream 44 = 413.05 kg/hr Stream 45 = 3237.57 kg/hr Total = 17539 kg/hr Material In Material Out Stream 16 = 14364.9 kg/hr Total = 14364.9 kg/hr = Stream 17 = 14329.44 kg/hr Stream 18 = 35.46 kg/hr Total = 14364.9 kg/hr

Year Amount of n-Butanal

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

(19)

Stream 42 = 13888.6 kg/hr Stream 44 = 413.05 kg/hr Stream 45 = 3237.57 kg/hr Total = 17539 kg/hr Material In Material Out Stream 16 = 14364.9 kg/hr Total = 14364.9 kg/hr = Stream 17 = 14329.44 kg/hr Stream 18 = 35.46 kg/hr Total = 14364.9 kg/hr

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

(20)

Stream 42 = 13888.6 kg/hr Stream 44 = 413.05 kg/hr Stream 45 = 3237.57 kg/hr Total = 17539 kg/hr Material In Material Out Stream 16 = 14364.9 kg/hr Total = 14364.9 kg/hr = Stream 17 = 14329.44 kg/hr Stream 18 = 35.46 kg/hr Total = 14364.9 kg/hr 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) 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 (kg/hr) 14364.90 14329.44 35.46

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 Page 12

CHAPTER 4 MATERIAL BALANCE

(21)

Material In Material Out Stream 16 = 14364.9 kg/hr Total = 14364.9 kg/hr = Stream 17 = 14329.44 kg/hr Stream 18 = 35.46 kg/hr Total = 14364.9 kg/hr 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) 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 (kg/hr) 14364.90 14329.44 35.46 Material In Material Out Stream 21 = 13839.5 kg/hr Total = 13839.5 kg/hr = Stream 22 = 13753.6 kg/hr Stream 23 = 85.91 kg/hr Total = 13839.5 kg/hr

MATERIAL BALANCE AROUND REACTOR

Stream number 4 13 14 15 0.00 547.35 3.55 135.02 CO (kg/hr) 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 (kg/hr) 8827.91 9201.39 14364.90 3664.32

Material In Material Out

Stream 4 = 8827.9 kg/hr Stream 14 = 14364.9 kg/hr Stream 13 = 9201.4 kg/hr Stream 15 = 3664.32 kg/hr Total = 18029 kg/hr = Total = 18029 kg/hr

Page 13

(22)

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) 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 (kg/hr) 14364.90 14329.44 35.46 Material In Material Out Stream 21 = 13839.5 kg/hr Total = 13839.5 kg/hr = Stream 22 = 13753.6 kg/hr Stream 23 = 85.91 kg/hr Total = 13839.5 kg/hr Stream number 21 22 23 Hydrogen (kg/hr) 3.58 0.08 3.49 CO (kg/hr) 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 Material In Material Out Stream 36 = 3283.17 kg/hr Total = 3283.17 kg/hr = Stream 37 = 3178.46 kg/hr Stream 38 = 104.70 kg/hr Total = 3283.17 kg/hr

MATERIAL BALANCE

(23)

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) 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 (kg/hr) 14364.90 14329.44 35.46 Material In Material Out Stream 21 = 13839.5 kg/hr Total = 13839.5 kg/hr = Stream 22 = 13753.6 kg/hr Stream 23 = 85.91 kg/hr Total = 13839.5 kg/hr Stream number 21 22 23 Hydrogen (kg/hr) 3.58 0.08 3.49 CO (kg/hr) 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 Material In Material Out Stream 36 = 3283.17 kg/hr Total = 3283.17 kg/hr = Stream 37 = 3178.46 kg/hr Stream 38 = 104.70 kg/hr Total = 3283.17 kg/hr Stream number 36 37 38 Hydrogen (kg/hr) 140.18 140.16 0.02 CO (kg/hr) 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 (kg/hr) 3283.17 3178.46 104.70 Page 14

CHAPTER 4 MATERIAL BALANCE

MATERIAL BALANCE AROUND FLASH SEPARATOR

Stream number 25 26 27 Hydrogen (kg/hr) 135.02 0.00 135.02 CO (kg/hr) 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 (kg/hr) 3664.32 502.52 3161.80

Material In Material Out Stream 25 = 3664.32 kg/hr Stream 26 = 502.52 kg/hr

Stream 27 = 3161.8 kg/hr Total = 3664.32 kg/hr = Total = 3664.32 kg/hr

(24)

Material In Material Out Stream 21 = 13839.5 kg/hr Total = 13839.5 kg/hr = Stream 22 = 13753.6 kg/hr Stream 23 = 85.91 kg/hr Total = 13839.5 kg/hr Stream number 21 22 23 Hydrogen (kg/hr) 3.58 0.08 3.49 CO (kg/hr) 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 Material In Material Out Stream 36 = 3283.17 kg/hr Total = 3283.17 kg/hr = Stream 37 = 3178.46 kg/hr Stream 38 = 104.70 kg/hr Total = 3283.17 kg/hr Stream number 36 37 38 Hydrogen (kg/hr) 140.18 140.16 0.02 CO (kg/hr) 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 (kg/hr) 3283.17 3178.46 104.70 Page 15

CHAPTER 4 MATERIAL BALANCE

(25)

Stream number 21 22 23 Hydrogen (kg/hr) 3.58 0.08 3.49 CO (kg/hr) 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 Material In Material Out Stream 36 = 3283.17 kg/hr Total = 3283.17 kg/hr = Stream 37 = 3178.46 kg/hr Stream 38 = 104.70 kg/hr Total = 3283.17 kg/hr Stream number 36 37 38 Hydrogen (kg/hr) 140.18 140.16 0.02 CO (kg/hr) 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 (kg/hr) 3283.17 3178.46 104.70 Page 16

CHAPTER 4 MATERIAL BALANCE

(26)

Stream number 36 37 38 Hydrogen (kg/hr) 140.18 140.16 0.02 CO (kg/hr) 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 (kg/hr) 3283.17 3178.46 104.70 Material In Material Out Stream 17 = 14339.44 kg/hr Stream 10 = 8712 kg/hr Total = 23041.44 kg/hr = Stream 11 = 9201.94 kg/hr Stream 19 = 13839.5 kg/hr Total = 23041.44 kg/hr Page 17

CHAPTER 4 MATERIAL BALANCE

MATERIAL BALANCE AROUND MIXER

1 8

2 3

M -1 0 1

2 8

2 7

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

(27)

Material In Material Out Stream 17 = 14339.44 kg/hr Stream 10 = 8712 kg/hr Total = 23041.44 kg/hr = Stream 11 = 9201.94 kg/hr Stream 19 = 13839.5 kg/hr Total = 23041.44 kg/hr 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

Material In Material Out

Stream 18 = 35.46 kg/hr Stream 28 = 3283.17 kg/hr Stream 23 = 85.91 kg/hr Stream 27 = 3161.8 kg/hr Total = 3283.17 kg/hr = Total = 3283.17 kg/hr Page 18

CHAPTER 4 MATERIAL BALANCE

MATERIAL BALANCE AROUND MIXER

2 2

2 6

M -1 0 2

4 0

3 9

(28)

Material In Material Out Stream 17 = 14339.44 kg/hr Stream 10 = 8712 kg/hr Total = 23041.44 kg/hr = Stream 11 = 9201.94 kg/hr Stream 19 = 13839.5 kg/hr Total = 23041.44 kg/hr 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 Material In Material Out Stream 41 = 14360.82 kg/hr Total = 14360.82 kg/hr =

Stream 42 = 13888.66 kg/hr Stream 43 = 59.10 kg/hr Stream 44 = 413.06 kg/hr Total = 14360.82 kg/hr Material In Material Out

Stream 22 = 13753.6 kg/hr Stream 40 = 14360.82 kg/hr Stream 26 = 502.52 kg/hr Stream 39 = 104.70 kg/hr Total = 14360.82 kg/hr = Total = 14360.82 kg/hr Page 19

CHAPTER 4 MATERIAL BALANCE

MATERIAL BALANCE AROUND MIXER

3 7

M - 1 0 3

4 5

4 3

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

(29)

Material In Material Out Stream 17 = 14339.44 kg/hr Stream 10 = 8712 kg/hr Total = 23041.44 kg/hr = Stream 11 = 9201.94 kg/hr Stream 19 = 13839.5 kg/hr Total = 23041.44 kg/hr 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 Material In Material Out Stream 41 = 14360.82 kg/hr Total = 14360.82 kg/hr =

Stream 42 = 13888.66 kg/hr Stream 43 = 59.10 kg/hr Stream 44 = 413.06 kg/hr Total = 14360.82 kg/hr

I-butanal (kg/hr) 10.55 19.86 30.42

Total (kg/hr) 3178.46 59.10 3237.57

Material In Material Out

Stream 37 = 3178.46 kg/hr Stream 45 = 3237.57 kg/hr Stream 43 = 59.10 kg/hr Total = 3237.57 kg/hr = Total = 3237.57 kg/hr Page 20

CHAPTER 4 MATERIAL BALANCE

(30)

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 Material In Material Out Stream 41 = 14360.82 kg/hr Total = 14360.82 kg/hr =

? P

Page 21

CHAPTER 4 MATERIAL BALANCE

MATERIAL BALANCE AROUND DISTILLATION COLUMN

Stream number 41 42 43 44

(31)

Material In Material Out

Stream 41 = 14360.82 kg/hr

Total = 14360.82 kg/hr

=

? P 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 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

(32)

? P

streams we used reference temperature equal to 25 0C.

ENERGY BALANCE AROUND THE COMPRESSOR K-101

Propylene Gas P1= 101.325Kpa T1=25oC

Propylene Gas P2= 2945Kpa T2=?

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

m





T

₌

T



P

2



2 1



P



Where n=0.0583 kmol/s



1



R=0.0821 m3atm/kmol K P= 1 atm

T=298.15 K

V=1.356 m3/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





T

₌

T



P

2

_

2 1



_P



Where T1=25 oC P1=101.325Kpa P2=2945 Kpa



1





(33)

? ? P???

m

₌



γ -1



= 0.137 γ =1.12 T2 = 200.5 oC

Work per kmol



γ E

P



 n −1 n Z T R n     2 ₁ W = 1 1    − n -1  P      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 _E P 1 3600 = 793 KW = 0.793MW

(34)

Similarly by putting the values in Excel Data Sheet we can calculate the power of all compressors which is given as:

Compressor Power Compressor Power Compressor Power

K-102 0.129 MW K-105 0.694 MW K-108 0.289 MW K-103 1.024 MW K-106 0.004 MW K-109 0.119 MW K-104 0.794 MW K-107 0.114 MW K-110 0.022 MW Page 24

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND REACTOR

(35)

Temperature 0C 105 45 120 120

Pressure kPa 5010 5000 5000 5000

Heat Flow kJ/hr 5.35E+06 -3.26E+07 -4.42E+07 -1.12E+07 Heat Flow In Heat Flow Out

Stream 4 = 5.35E+06 kJ/hr Stream 14 = -4.42E+07 kJ/hr Stream13= -3.26E+07 kJ/hr Stream 15 = -1.12E+07 kJ/hr

Total = -2.72E+07 kJ/hr Total = -5.55E+07 kJ/hr Cooling Duty Qp = -2.82E+07 kJ/hr

Page 25

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND HEAT EXCHANGER E-101

Stream number 2 3

Hydrogen (kg/hr) 0.00 0.00

CO (kg/hr) 0.00 0.00

Propylene (kg/hr) 8781.69 8781.69 Propane (kg/hr) 46.24 46.24

(36)

n-butanal (kg/hr) 0.00 0.00 I-butanal (kg/hr) 0.00 0.00 Total (kg/hr) 8827.91 8827.91

Temperature 0C 200 77

Pressure kPa 2945 2925

Heat flow kJ/hr 7.02E+06 4.90E+06 Heat Flow In Heat Flow out

Stream 2 = 7.02E+06 kJ/hr Stream 3 = 4.90E+06 kJ/hr Cooling Duty Qp = -2.12E+06 kJ/hr

Page 26

CHAPTER 5 ENERGY BALANCE

Temperature of Cooling water in = 25 0C, Temperature of Cooling water out = 30 0C 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 oC Temperature of cooling water out = 30 oC Temperature of Steam in () = 120 oC Temperature of Steam out = 120 oC

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

(37)

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

(38)

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 (kg/hr) 14364.90 14329.44 35.46 Temperature oC 40 40 40 Pressure kPa 4968 4968 4968

Heat Flow kJ/hr -4.70E+07 -4.69E+07 -1.29E+05 Heat Flow In Heat Flow Out

Stream 16 = -4.70E+07 kJ/hr Stream 17 = -4.69E+07 kJ/hr Stream 18 = -1.29E+05 kJ/hr Total = -4.70E+07 kg/hr = Total = -4.70E+07 kg/hr

Page 28

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND VAPOR LIQUID SEPARATOR

Stream number 25 26 27

Hydrogen (kg/hr) 135.02 0.00 135.02 CO (kg/hr) 2382.38 0.07 2382.31

(39)

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 (kg/hr) 3664.32 502.52 3161.80 Temperature oC 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

Heat Flow In Heat Flow Out

Stream 25 = -1.21E+07 kJ/hr Stream 26 = -1.68E+06 kJ/hr Stream 27 = -1.05E+07 kJ/hr Total = -1.21E+07 kJ/hr = Total = -1.21E+07 kJ/hr

Page 29

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND VAPOR LIQUID SEPARATOR

(40)

3.58 0.08 3.49 CO (kg/hr) 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 oC 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

Heat Flow In Heat Flow Out

Stream 21 = -4.64E+07 kJ/hr Stream 22 = -4.61E+07 kJ/hr Stream 23 = -3.14E+051 kJ/hr Total = -4.64E+07 kJ/hr = Total = -4.64E+07 kJ/hr

Page 30

CHAPTER 5 ENERGY BALANCE

(41)

Stream number 36 37 38 140.18 140.16 0.02 CO (kg/hr) 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 (kg/hr) 3283.17 3178.46 104.70 Temperature oC 80 80 80 Pressure kPa 4990 4990 4990

Heat Flow kJ/hr -1.06E+07 -1.03E+07 -3.28E+05 Material In Material Out

Stream 36 = -1.06E+07 kJ/hr Stream 37 = -1.03E+07 kJ/hr Stream 38 = -3.28E+05 kJ/hr Total = -1.06E+07 kJ/hr Total = -1.06E+07 kJ/hr

Page 31

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND MIXER

1 8

2 3

M -1 0 1

2 8

2 7

(42)

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 oC 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 Heat Flow In Heat Flow Out

Stream 18 = -1.29E+05 kJ/hr Stream 28 = -1.09E+07 kJ/hr Stream 23 = -3.14E+05 kJ/hr

Stream 27 = -1.05E+07 kJ/hr

Total = -1.09E+07 kJ/hr Total = -1.09E+07 kJ/hr

Page 32

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND MIXER

22

26 M -102

40

(43)

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 oC 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 Heat Flow In Heat Flow Out

Stream 22 = -4.61E+07 kJ/hr Stream 40 = -4.81E+07 kJ/hr Stream 26 = -1.68E+06 kJ/hr

Stream 39 = -3.28E+05 kJ/hr

Total = -4.81E+07 kJ/hr = Total = -4.81E+07 kJ/hr

Page 33

CHAPTER 5 ENERGY BALANCE

ENERGY BALANCE AROUND MIXER

(44)

M - 1 0 3

4 5

4 3

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 oC 80 90 80 Pressure kPa 4990 260 260

Heat Flow kJ/hr -1.03E+07 -1.50E+05 -1.04E+07 Heat Flow In Heat Flow Out

Stream 37 = -1.03E+07 kJ/hr Stream 45 = 3237.57 kJ/hr Stream 43 = -1.50E+05 kJ/hr

Total = 3237.57 kJ/hr = Total = 3237.57 kJ/hr

Page 34

CHAPTER 5 ENERGY BALANCE

(45)

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 Temperature oC 40 211 44 115 4968 5000 4965 4990

-4.69E+07 -2.92E+07 -3.26E+07 -4.35E+07 Heat Flow In Heat Flow Out

Stream 17 = -4.69E+07 kJ/hr Stream 11 = -3.26E+07 kJ/hr Stream 10 = -2.92E+07 kJ/hr Stream 19 = -4.35E+07 kJ/hr Total = -7.61E+07 kJ/hr Total = -7.61E+07 kJ/hr

Page 35

CHAPTER 5

(46)

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 Heat Flow Out

Stream 41 = -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

(47)

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

(48)

Page 37

CHAPTER 6 DESIGN OF

EQUIPMENTS

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

(49)

R e si d e n c e T im e s Page 38

CHAPTER 6 DESIGN OF

EQUIPMENTS

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.

106

105 104

Batch Reactor

Backmix Reactor Cascade _Backmix

103 102 10 Tubular Reactor 1 10-1 10-4 10-3 10-2 10-1 1 10 102 103 Production Rate kg/s

(50)

Page 39

CHAPTER 6 DESIGN OF

EQUIPMENTS

CSTR (Continuous Stirred-Tank Reactor)

In a CSTR, one or more fluid reagents are introduced into a tank reactor equipped with an impeller w hile 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 t hrough 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

(51)

• Distribution of catalyst

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

Agitation

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

• Gases with gases

Page 40

CHAPTER 6 DESIGN OF

EQUIPMENTS

• 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.

(52)

Liquid Product

Feed

The Vessel

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

Page 41

CHAPTER 6 DESIGN OF

EQUIPMENTS

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

(53)

Partial pressure of Propylene Concentration of Rhodium PE CRh = 13.5 bar = 0.92 mol/m3 Concentration of Ligands CLig = 22.08 mol/m3 as CRh:CLig = 1:24 Conversion of Reaction XA = 95%

Initial Flow rate of Propylene FA0 = 58 mol/sec Temperature of Reaction T = 393.15 K Reaction Pressure P = 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

Page 45

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

Impeller Types

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

Page 42

CHAPTER 6 DESIGN OF

EQUIPMENTS

directions. Above a D/T ratio of 0.55, pitched blade turbines become radial flow impellers.

Flow Pattern Impeller

Axial Flow Propeller, Pitched Blade Turbine, Hydrofoils

Radial Flow Flat-blade Impeller, Disc Turbine (Rushton), Hollow-blade Turbine High Shear Cowles, Disc, Bar, Pointed blade Impeller

Specialty Retreat Curve Impeller, Sweptback Impeller, Spring Impeller Up/Down 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

(54)

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 Page 45 gas dispersion.

Impeller Location

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

Impeller Selection

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

Modeling of mass transfer and chemical reaction

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

Page 43

CHAPTER 6 DESIGN OF

EQUIPMENTS

(55)

= 16.38 m3

Page 45

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.

Page 44

CHAPTER 6 DESIGN OF

EQUIPMENTS

Kinetics

The kinetics of the hydroformylation reaction in the presence of a RhCl(CO) (TPPTS)2/TPPTS complex catalyst were experimentally determined by Yang et al.

(56)

= 16.38 m3

Page 45

(2002). These authors varied the propylene concentration, the initial pressure, the H2/CO ratio, the temperature, the rhodium concentration and the ligand to rhodium ratio in an orthogonal experimental design to obtain the following rate expression:

The constants are defined in Table 6.2:

SIZING OF CSTR

In sizing of CSTR first of all we should have rate expression 6.7 which, we have already developed.

VOLUME OF REACTOR

Partial pressure of Hydrogen PA = 17.1 bar

Partial pressure of Carbon monoxide PB = 19.4 bar

CHAPTER 6 DESIGN OF

EQUIPMENTS

Head Volume VH = 12.70 m3

Volume of Reactor V = 139.7 m3

L

ENGTH AND DIAMETER

For CSTR Length to diameter ratio is 1. So L/D = 1

Since

V = (π / 4) ×L × D2 =127 m3 Where

L = Length of the reactor D = Diameter of the reactor

(57)

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 + C_c Page 46

CHAPTER 6 DESIGN OF

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

(58)

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

Outside Diameter D0 = Di + 2t = 5.45 + 2(0.135.9) = 5.72 m

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 = P _{D ∗ 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 DESIGN OF

EQUIPMENTS

AGITATOR DESIGN

(59)

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

Impeller Height above Vessel floor

D E = T/3 = T/3 = 1.82 m = 1.82 m Length of Impeller Blade

Width of Impeller Blade

L W = 0.25D = D/5 = 0.45 m = 0.36 m Page 48

CHAPTER 6 DESIGN OF

EQUIPMENTS

(60)

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 = N_Px 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

(61)

2

EQUIPMENTS

N

₌

ND

a

ρ

Re

µ

N _Re = 6.04 × 106

For such a high Reynolds number, which is greater than 105 we use the relation for power requirement as:

Power P = KT x N3 x D5 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/m3 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

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

(62)

m 2 Page 50

CHAPTER 6 DESIGN OF

EQUIPMENTS

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 fs – shear stress = 550 kg/cm2 Zp = (1.5 x 776 x 100) /550 = 211.5 cm3 πd3/16 = 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

M_e= 1 2 [M + M 1 + T2 _] 2 2 _M_e₌ 2 [925 + 925 + (776 ∗ 1.5) ] Me = 1206 kg. m

(63)

The stress due to equivalent blending F = Me/Z

Z = π d3 / 32 = π x 103 / 32 = 98.13

F = (1206 x 100)/98.13 = 1229 Kg/cm2 This is within the allowable limits of stress.

Overhang of agitator shaft between bearing and agitator I = 130 cm

Page 51

CHAPTER 6 DESIGN OF

EQUIPMENTS

Modulus of elasticity E = 19.5 x 105 kg/cm2

Shaft deflection ɗ = (Fm x I3)/(3E x π x D4/64) ɗ = 0.54 cm

HUB AND KEY DESIGN

Hub diameter of agitator = 2 x shaft diameter = 20 cm

Length of the hub = 2.5 x 36.3 = 90.82 cm Length of key = 1.5 x shaft dia = 15 cm

HEAT TRANSFER IN REACTOR

Cooling Jacket area available A =

π DH + πD2/4

= (π x 5.42 x 5.42) + (π x 5.422 /4) = 153.29 m2 CW inlet temp = 28 oC CW outlet temp = 33 oC Approaches; ΔT1= 120 – 28 = 92 ΔT2= 120 – 33 = 87 LMTD = 89.47 0C Heat, removable by jacket

(64)

= 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 107

m = Q/( CpΔTM) = 77892 kg/hr Page 52

CHAPTER 6 DESIGN OF

EQUIPMENTS

SPECIFICATION SHEET

Identification

Item Reactor Item Number R-101 Number of Item 1 Operation Continuous

Type Continuous Stirred Tank Reactor

Design Data

Volume 139.71 m3 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