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Ethyl Benzene Plant Design

Project Report

Duration:12th Aug, 2014 - 12th Nov, 2014

Department of Chemical Engineering

IIT Kanpur

Submitted by :

Rohit Kakkar|Salman Ahmad Khan|Shivang Sharma

Rohan Bishnoi|Himanshu Bareja|Nikhil Kumar

Umang Arora

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Table of Contents 1. Acknowledgement3

2. Executive Summary..4

3. Introduction...4

3.1 General Philosophy behind design of chemical processes.4

3.2 Introduction to Ethyl Benzene6

3.3 Ethyl Benzene Reaction System.6

3.4 Reaction Kinetics7

3.5 Commercial methods for production of Ethyl Benzene..8

4. Process Simulation9

4.1 Fluid Property Package...9

4.2 Process Description11

4.3 Process flow diagram.12

4.4 Degrees of freedom13

4.5 Optimization...13

4.6 Idea for heat integration. 17

4.7 Complex Column Configuration17

4.8 Control Structure21

4.9 Controller Performance..24

5. Economics27

5.1 Size of the Heat Exchangers...27

5.2 Size, Capital Costs and Operating Cost of the equipments28

5.3 Objective function(J)..29

6. References29

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Acknowledgement

We are highly indebted to Prof. Nitin Kaistha for making us learn

simulations and then challenging us by giving an open ended design

project which truly tested our skills, patience, team work and

commitment towards a particular task. We would also like to thank Sir

and Mr. Vivek Kumar for their guidance and supervision during the

entire period of project.

We are really grateful to all the Teaching Assistants and Lab staff who

directly or indirectly helped us in completion of our project since it

would not have been possible without their strong support and co-

operation.

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2 Executive Summary

An ethyl benzene plant was designed using the liquid phase alkylation of benzene with ethylene

over zeolite acid catalyst. Fresh ethylene was assumed to have 5% ethane impurity while fresh

benzene had 0.01% toluene impurity. Alkylation reactions, which led to the formation of ethyl

benzene and a side product diethyl benzene and transalkylation reaction in which diethyl

benzene with benzene formed ethyl benzene, were all carried out in the packed bed reactors. To

limit the formation of side products, alkylator was run at high benzene to ethylene ratio.

The separation system consisted of a flash drum and 3 distillation columns. Vapor stream from

the flash drum was separated into ethane from the vent and benzene in the bottoms by first

distillation column. This benzene mixed with the liquid stream from the flash drum and the

combined stream went to second distillation column. Second column separated benzene in the

distillate and mixture of ethyl benzene and diethyl benzene at the bottoms. A part of benzene

from distillate was recycled to the alkylator and rest of it was sent to transalkylator. Third

column took bottoms of the second column as feed and gave ethyl benzene in the distillate and

diethyl benzene in the bottoms which was recycled to the transalkylator.

The dominant design variables were benzene to ethylene excess ratio, split ratio of benzene sent

to the transalkylator and temperature of transalkylator. Each of them were varied with the

objective function (J) which determined the annual profit of the plant. We mostly obtained bell

shaped curves and the values of variables at which the J got maximized were used in the final

flowsheet.

The plant capacity is 84086.64 tonnes of ethyl benzene per annum. The overall capital cost of the

plant came out to be $2.5 million and it consumed yearly raw materials worth $67.2 million and

yearly energy of $4.63 million. The revenue from the plant was calculated to be $89.74 million

per annum and the annual profit was $17.02 million.

3 Introduction

3.1 General philosophy behind design of chemical processes[1]

The chemical process industry is mainly involved in manufacturing of wide range of products

which improves the quality of life of humankind and generates employment. Chemical engineers

deal with a lot of obstacles while designing a process especially when substances involved have

high chemical reactivity, high toxicity, and high corrosivity operating at high pressures and

temperatures.

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Designing of the plant involves a thorough understanding of the reactions taking place in the

process with the major emphasis on reaction kinetics. Most of the time before starting the

simulation goes in testing the reaction kinetics and choosing a relevant thermodynamic package

which satisfies the experimental vapor liquid equilibrium data and enthalpy data of components.

Douglas[2,3] has proposed a hierarchical approach to the conceptual design in which design

process follows a series of decisions and steps. The decisions are listed as follows:

1. Decide whether the process will be batch or continuous.

2. Identify the input/output structure of the process.

3. Identify and define the recycle structure of the process.

4. Identify and design the general structure of the separation system.

5. Identify and design the heat-exchanger network or process energy recovery system.

High efficiency of usage of raw material is a requirement of the majority of chemical processes.

The extent of recycling of the unused reactants depends on the ease with which they can be

separated from the products and the single pass conversion of the reactors. We can recycle the

unreacted raw materials in three ways:

1. Separate and purify unreacted feed material from products and then recycle.

2. Recycle feed and product together and use a purge stream.

3. Recycle feed and product together and do not use a purge stream.

Specifications of streams and process conditions are influenced by physical processes as well as

economic conditions. The conditions used in the process generally represents a trade-off between

process performance and capital and operating cost of the equipment. Final selection of the

operating conditions should be made only after the economic analysis of the process.

On operating at higher pressure (>10 bar), we would be needing thicker walled more expensive

equipment while at pressure lower than ambient tends to make equipment large and may require

special construction techniques to prevent inward flow of outside impurities, thus increasing the

cost of equipment.

Several critical temperature limits apply to the chemical processes. At higher temperatures, there

is a significant drop in the physical strength of the common construction materials (primarily

carbon steel) and it must be replaced by a costly material. There should be a rational explanation

for selecting particular operating conditions which should be supported by economic analysis.

Most of the commercial reactions involve catalysts and the competence of a company is often the

result of a unique catalyst they use.

Choosing the best physical method is an extremely important part of any simulation. Wrong

property package will lead to incorrect simulation results which could not be trusted. Everything

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ranging from volumetric flow rates to energy balances to separation in the equilibrium stage

units depends on correct thermodynamic data.

In general, seven steps are involved in simulating any chemical process in a software. They are

selection of chemical components, thermodynamic models, process topology, feed stream

properties, equipment parameters, output options and convergence criteria.

For optimizing a problem, we look for decision variables, objective function, constraints, global

optimum and minimum optimum.

3.2 Introduction to Ethyl Benzene

Ethyl Benzene (EB), compound with a chemical formula of C6H5CH2CH3 is used as an

intermediate in the making of styrene[4]. Styrene is a building block in the manufacturing of

polystyrene which is used for producing disposable plastic cutlery and dinnerware, CD jewel

cases, smoke detector housings and so on. EB and styrene units are generally installed together

in order to facilitate the energy economy by integrating energy flows of the two units.

Ethyl Benzene is mainly manufactured via liquid phase alkylation or a gas phase alkylation of

benzene. The traditional method used in industries is the liquid phase alkylation of benzene with

ethylene over AlCl3 as the catalyst. It has been observed that in case of liquid phase alkylation,

temperature is lower while selectivity of ethyl benzene and pressure required is higher compared

to gas phase alkylation.[5]

A major difficulty faced during manufacturing of EB is that ethyl benzene is more reactive

compared to benzene with respect to ethylene due to lower activation energy and thus it leads to

the formation of diethyl benzenes (DEBs). To limit the formation of DEBs and other polyethyl

benzene (PEBs) we use a large excess ratio of Benzene to Ethylene in the feed to the alkylation

reactor. However, large excess ratio leads to higher equipment, separation and recycle cost.

3.3 Ethyl Benzene Reaction System

Ethyl Benzene is produced by reacting benzene with ethylene in liquid phase over AlCl3 as the

catalyst. Ethylene contains impurities of ethane while benzene contains slight impurities of

toluene. Ethane does not participate in the reaction while toluene reacts with ethylene to form

ethyl benzene and propylene but we are assuming only 0.01% toluene in the fresh benzene and

thus the amount of EB produced using toluene has been considered negligible.

Ethylene on reacting with benzene forms ethyl benzene and ethyl benzene on reacting further

with ethylene forms diethyl benzene. Diethyl benzene again on reacting with ethylene forms

triethyl benzene. The activation energy for the reaction between ethyl benzene and ethylene is

lower than the activation energy of ethylene reacting with benzene and thus if reaction is allowed

to proceed without any restriction, we will get more diethyl benzene than ethyl benzene. Diethyl

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benzene is an undesirable product because most of the ethyl benzene is used in the making of

styrene and for styrene process, fresh ethyl benzene should have less than 2ppm diethyl benzene.

So to limit the formation of diethyl benzene and other polyethyl benzenes, we use a high benzene

to ethylene ratio in the alkylation reactor. The reactions that take place in the entire process are

as follows:

C6H6 + C2H4 C6H5C2H5 .1

Benzene Ethylene EB

C6H6 + 2C2H4 C6H4 (C2H5)2 ......2

Benzene Ethylene Diethyl Benzene

C6H5C2H5 + C2H4 C6H4 (C2H5)2 ..3

EB Ethylene DEB

C6H6 + C6H5 (C2H5)2 2 C6H5C2H5 ..........................4

Benzene DEB EB

2 C6H5C2H5 C6H6 + C6H4 (C2H5)2 ......5

EB Benzene DEB

Reactions 4 and 5 is basically an equilibrium reaction and is called trans-alkylation reaction.

Reactions 1, 2 and 3 occur in alkylation reactor while 4 and 5 occur in transalkylation reactor

both of which are packed bed reactors. Catalysts used for the alkylation reactions are different

from that of the transalkylation reaction and so we can safely ensure that first three reactions

occur in alkylator while 4 and 5 occur in transalkylator.

3.4 Reaction Kinetics[6,7]

Chemical kinetics is the study of rate of chemical processes. It is one of the most fundamental

thing that is researched out before designing a process. Getting relevant kinetics is considered as

one of the major success for designing a process.

During the course of our project, we faced a lot of hardships in finding kinetics for ethyl benzene

process. Some papers considered transalkylation as an equilibrium reaction while some did not,

value of A(Arrhenius constant) in different papers were of different order of magnitudes, and the

one we chose finally also doesnt gives us practical results. The kinetics used is as follows:

r1 = 0.084*exp(-9502/RT)CE1.0CB

0.32

r2 = 0.603*exp(-15396/RT)CE1.3CB

0.33

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r3 = 0.00085*exp(-20643/RT)CE1.77CEB

0.35

For reaction 4 and 5, we regressed data for CBZ0 and CDEB

0 and initial rates and obtained rate

constant of the forward reaction. Alternatively, chemical equilibrium constant for 4th reaction is:

K = XBXDEB/XEB2

At 571.15 K, K is 0.883. At ordinary temperature T, K can be obtained the relationship

ln (

571.15) =

1

(

1

1

571.15)

Using the equilibrium constant and forward rate constant, we found out the backward rate

constant. From the above kinetics, we were able to do the reaction with 100% ethylene

conversion even at 60 C but the literature, encyclopaedias and actual plant data tells us that the

temperature should be in the range of 200 C - 300C for 100% ethylene conversion. Thus we

obtained a wide deviation from the experimental data which proved our kinetics to be incorrect.

3.5 Commercial Methods for Production of Ethyl Benzene[7]

There are primarily two sources to produce ethylbenzene. The major being alkylation of benzene

and other being superfractionation of C8 aromatic streams. The alkylation of benzene is further

carried out commercially by

a) liquid phase alkylation

b) gas phase alkyllation

LIQUID PHASE ALKYLATION:

Alkylation of benzene in liquid phase using aluminium chloride as catalyst is the most used

method for ethylbenzene production. Different companies like shell, union carbide, Dow

chemicals use this method

The reaction is exothermic in nature (H-114 kJ/mol), kinetically fast and produce good yields

of ethylbezene. Generally hydrogen chloride or ethyl chloride is used as a catalyst promoter

which decreases the amount of aluminium chloride needed. Instead of AlCl3 lewis acids

catalysts, BF3,FeCl3, AlBr3 are also used.

VAPOR-PHASE ALKYLATION:

Earlier vapor phase alkylation were not able to compete with liquid phase alkylation. The

alkylation process using boron trifluoride as catalyst had little success and suffered drawbacks

like high maintenance costs caused by corrosion from small amount of water.

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However the Mobil Badger process is the most modern and successful vapor phase process for

production of ethylbenzene. The process uses ZSM-5 synthetic zeolite catalyst.

Zeolites were used earlier for alkylation process but they deactivated quickly because of coke

formation and low catalytic activity. The catalyst used by Mobil had larger resistance to coke

formation and high catalytic activity for transalkylation and alkylation.

Liquid phase alkylation of benzene with ethylene over zeolite acid catalyst has been used in our

project.

4 Process Simulation

4.1 Fluid Property Package

Vapor liquid equilibrium data of Benzene and Ethyl Benzene for SRK, Peng Robinson, Lee

Kesler and UNIFAC was obtained in Aspen Plus. Then it was compared with the data obtained

from Dechema, Volume 7 which basically represents the experimental data. Following results

were obtained :

Table 4.1.: Root Mean Square difference between the values obtained from 2 data

Package SRK Peng Robinson UNIFAC Lee Kesler

Root Mean Square Difference .0377 .0392 .0387 .0394

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(a)

(b)

(c)

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(d)

Figure 4.1.1 : Blue line represents Y and red line reprsents X of experimental data from Dechema while

Green line represents Y and Purple line represents X of Aspen Plus data. 1(a) Plot for UNIFAC, 1(b) Plot

for Lee Kesler method, 1(c) Plot for SRK, 1(d) Plot for Peng Robinson

4.2 Process Description

Stream containing mostly ethylene with an impurity of ethane (~5%) and stream having benzene

with a slight toluene impurity (~0.01%) were used as fresh feeds. Fresh feed containing ethylene

was passed through compressor to raise the pressure to 50 bars since it was in the gas phase

while fresh benzene was passed through a pump which raised its pressure to 50 bar. Both these

fresh feeds passed through a valve which reduced their pressure to around 40 bars. Pump and

compressor were supplied with the required amount of electricity.

Fresh benzene mixed with recycle benzene and the combined stream along with fresh ethylene

with some ethane was used as the feed to the alkylation reactor. The feed passed through a heater

which was used for controlling the temperature before sending to the reactor.

Output of the alkylator reactor mixed that of the transalkylator and the combined stream passed

through a cooler followed by a valve to reduce pressure to 1 atm. This stream entered the

separator which was used to separate liquid and vapor components of the inlet stream. It was

used because ethane in the vapor phase made the separation difficult and if the inlet stream was

directly send for the distillation, then it would have significantly enhanced the reboiler duty of

the column.

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The vapor stream from the separator went through compressor to increase stream pressure to

column pressure and cooler to reduce temperature which rose due to compression and finally into

the distillation column. The column gave almost pure ethane in the distillate and almost pure

benzene in the bottoms and both these distillate and bottoms which mixed with the liquid stream

from separator.

The liquid stream from the separator passed through a valve and finally went into second

distillation column. This column gave almost pure benzene in the distillate and a mixture of ethyl

benzene and diethyl benzene in the bottoms. The benzene from the distillate was split in two

parts one of which was recycled back to mix with the fresh benzene and served as feed to the

alkylation reactor while second part was send to the transalkylation reactor in which it reacted

with diethyl benzene.

The bottoms from the second distillation column went to the third column. This column gave out

almost pure ethyl benzene with slight impurities of benzene and toluene in the distillate and

diethyl benzene in the bottoms. The diethyl benzene from the bottoms passed through a cooler

followed by a mixer where it mixed with benzene. The mixed stream went through a pump

followed by a heater and finally into a transalkylation reactor. The output from the

transalkylation reactor was recycled to mix with the output stream from alkylation reactor.

4.3 Process Flow Diagram

Figure 4.31 : Process Flow Diagram of the fully optimized flow sheet developed on Aspen Hysys

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4.4 Degrees of Freedom

Table 4.4.1 : Degrees of freedom in the flowsheet

Control Degrees of Freedom 43

Steady State Operating Degrees of Freedom 17

Steady State Design Degrees of Freedom 25

4.4 Optimization

Sizing the Reactors

Alkylator

An adjustor block was used to size the reactor for 99% conversion of ethylene. This adjustor

block was used during the optimization of other variables, so as to achieve a final value of the

volume for which the optimized flow sheet runs, with a 99% conversion. The final value arrived

at was 20m3.

Transalkylator

Using a procedure similar to that of the alkylator, an adjustor block was used on the transalkator

to achieve a conversion of 99% of the equilibrium conversion. The adjustor block was used on

the transalkylator while optimizing the flow sheet for other variables, so as to get the volume of

the transalkylator for the final optimized value of other dominant design variables. The final

volumne arrived at was 7 m3.

The final flowsheet was optimized for 4 dominant design variables.

The final dominant deign variables were chosen to be the following

-Excess Ratio of Benzene to Ethylene

-Split of Benzene sent to the trnasalkylator

-Inlet Temperature of the Transalkylator

-Benzene Leakage down the benzene recycle column

Excess Ratio Of Benzene

The excess ratio of benzene was defined as the ratio of the total benzene entering the reactor to

the amount of ethylene entering the reactor. On varying the excess ratio, the maximum profit

achieved was at a value of 1.8 excess ratio.

When the value of the excess ratio is below 1.8, concentration of ethylene increases in the

stream, this leads to the production of more diethylbenzene due to the side reaction. This causes

more diethylbenzene to enter the second and third columns. Especially, in the third column the

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boilup increases because increase in the amount of diethylbenzene causes the concentration of

ethylbenzene to decrease and this leads to a more difficult separation. Because of this reason the

running and fixed costs of the reboiler end up increasing along with the size of the column.

While on increasing the excess ratio above 1.8, though the amount of diethylbenzene decreases

but the excess benzene leads to a more difficult separation in the first colum, this causes an

increase in the operating and capital in the second column.

Figure 4.4.1 : Variation of J with Benzene to Ethylene excess ratio in alkylator

Split Ratio

This optimization was done after optimizing the excess benzene ratio and setting its value to 1.8.

On varying the Split ratio, a maximum profit was achieved at a split of 90 kmol/hr. i.e. out of the

total benzene coming out as distillate from the second column, 90 kmol/hr was diverted to the

transalkylator.

If the split ratio is kept below 90kmol/hr, the Diethyl Benzene in the transalkylator doesnt get

enough Benzene to react with, thus the amount of DiethylBenzene in the recycle increases, this

leads to an increase in the size of the equipment of the second and third distillation columns,

namely the reboiler duty and the size of the reboiler.

1.20E+07

1.30E+07

1.40E+07

1.50E+07

1.60E+07

1.70E+07

1.80E+07

1 1.5 2 2.5 3 3.5 4 4.5

J (E

con

om

ic F

un

ctio

n)(

$)

Excess Ratio

J vs Excess Ratio

J vs1.8

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And on increasing the amount of benzene split, the amount of benzene increases in the recycle

and though the amount of diethyl benzene decreases , the amount of benzene recycle increases

and this causes an increase in the size of the distillation column separating benzene.

Figure 4.4.2 : Variation of J with the ratio of benzene sent to transalkylator to total benzene produced

from recycle column

Inlet Temperature of The TransAlkylator

On varying the temperature of the transalkylator, the maximum profit was achieved at an inlet

temperature of 1600C.

Below an inlet temperature of 1600C, the profits increase because of an increase in the

conversion, even though the reaction is mildly exothermic. This happens because at low

temperatures the reaction doesnt reach equilibrium and kinetically the reaction is favored at

higher temperatures. This causes the conversion to increase because of temperature rise. This

leads to more conversion for the same amount of split thus rendering the process cheaper.

When the temperature is increased beyond 1600C, the reaction having reached equilibrium,

starts to shift backwards because of its exothermicity. This causes the process to become more

costly for the same split ratio thus decreasing profits.

1.30E+07

1.35E+07

1.40E+07

1.45E+07

1.50E+07

1.55E+07

1.60E+07

1.65E+07

1.70E+07

1.75E+07

40 65 90 115 140 165 190 215

J (E

con

om

ic F

un

ctio

n)(

$)

Split Ratio

J vs Split Ratio

J vs Split Ratio

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Figure 4.4.3: Varitaion of J with the inlet temperature of transalkylator

Benzene Leakage from The Recycle Column

There was no change in the profit as the benzene leakage was changed. So the final value of the

benzene leakage chosen was around 2e-07, so as to always remain in the optimum region for

evwn substantial changes in operating conditions.

Figure 4.4.4 : Variation of J with benzene leakage from the recycle column

4.00E+06

6.00E+06

8.00E+06

1.00E+07

1.20E+07

1.40E+07

1.60E+07

1.80E+07

100 150 200 250 300

J (E

con

om

ic F

un

ctio

n)(

$)

Temperature (oC)

J vs Transalkylator Temperature

J vs Transalkylator Temperature

1.54E+07

1.56E+07

1.58E+07

1.60E+07

1.62E+07

1.64E+07

1.66E+07

1.68E+07

1.70E+07

1.72E+07

0.00E+00 2.00E-07 4.00E-07 6.00E-07 8.00E-07 1.00E-06 1.20E-06

J (E

con

om

ic F

un

ctio

n)(

$)

Benzne Leakage

J vs Benzene Leakage

J vs Benzene

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4.5 Ideas for heat integration

Hot Streams: Compressed Ethylene at 267 oC

Liquid flow to reboiler of Product column at 203 oC

Cold Stream: Liquid flow to reboiler of Benzene recycle column at 182.3 oC

The cold stream is first heated using liquid flow to reboiler of product column by passing

through a heat exchanger. This increases the quality of old stream from 0 to 0.05. This is then

heated using compressed ethylene stream with the help of another heat exchanger to increase the

vapour quality to 0.085.

To get a rough idea of the energy savings, the cold stream is copied from the stream entering the

reboiler of Benzene recycle column which is the passed through 2 heat exchangers, one heated

by liquid flow to reboiler of Product column and another by compressed ethylene. The rest of

enegry was supplied using an auxillary reboiler. This decreased energy consumption from 6181

kW to 1283 kW, thus saving $2.16 million.

4.6 Complex Column Configurations

Four complex column configurations were made to show different separation methods possible

and the reduction in the total reboiler duty as a consequence of using the complex distillation

columns was analyzed.

In the first complex configuration of the columns, the Benzene Recycle Column and the Product

Column were replaced. The configuration assembled was, as shown in figure 4.5.1. The feed was

introduced in a regular column, but with a vapor side draw. Benzene was separated at the top,

EthylBenzene was removed at the middle and this stream was subjected to further distillation in

a side stripper and DiethylBenzene was removed at the bottom.

Reboiler duty of this configuration was 5164 kW while that of conventional configuration is

5738 kW. Thus reboiler duty of this configuration is about 10% lower than that of the

conventional column

Energy Consumption Savings

Without Heat Integration 6181kW 0

Heat Integration 1283kW $2.16 million

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Figure 4.5.1: First complex column configuration showing main column with a side rectifier. Mostly

Benzene in distillate, Ethyl Benzene from side rectifier and Diethyl Benzene from the bottoms of main

column

For the second complex configuration, a simple Petlyuk column was assembled as shown in

figure 4.5.2. This configuration was designed to replace the Ehtane separation and Benzene

Recycle Columns. The feed is first partially distilled in the pre-fractionator. This tray section

separates benzene from diethylbenzene. The exit streams form the pre-fractionator are fed to the

main separation section. Benzene is removed at the top, Ethylbenzene is removed at middle and

a mixture of diethyl benzene and ethyl benzene was removed at the bottoms.

The pre-fractionator gives out mostly benzene, ethyl benzene and very little diethyl benzene as

vapor from the top and very little benzene, ethyl benzene and diethyl benzene as liquid from the

bottom to the main column. It also receives the liquid stream at the top and vapor stream at the

bottom from the main column. The main column gives out benzene from the top, ethyl benzene

from the side draw and diethyl benzene from the bottoms.

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Figure 4.5.2: Petlyuk Column

Reboiler duty of petlyuk column came out to be around 5738 kW, which is only a 1.7 % decrease

from the initial reboiler duty. This column doesnt prove to be much efficient in terms of

reduction of total reboiler duty.

The purpose of the third assembled column was to replace the Ethane Separation column and the

Benzene recycle column. The column structure is fairly simple, it has a conventional separation

section with a regular reboiler and condenser, with a side draw in the middle. Ethane is vented

out at the top, benzene is removed at middle and a mixture of ethyl benzene and diethyl benzene

was removed at the bottom.

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Figure 4.5.3: Column with a side draw. Benzene at the top, ethyl benzene in the middle and diethyl

benzene in the bottoms

Reboiler duty of this configuration came out to be 4481 kW, which is a 23% decrease from the

conventional reboiler duty.

The final configuration assembled is shown in figure 4.5.3.

The assembly in Figure 4.5.5 consists of two tray section, with the bottom product of the first

section being sent to the second. This configuration was meant to replace the Recycle Column

and the Product Column.

The initial tray section was used to separate benzene from the feed. The bottoms product of the

first column was introduced as feed into the second column. The second column sends out vapor

stream to the first column from the same stage on which it received the feed. Ethyl Benzene

comes out in the distillate and Diethyl Benzene from the bottoms of the second column.

Figure 4.5.4: 2 columns with a single reboiler

Reboiler duty of the configuration is 15% less than the conventional column configuration.

Table 4.5.1 : Summary of reboiler duty reduction on implementing complex column

Configuration Power Consumed(kW) % Saving

1 5164 10

2 5738 1.7

3 4481 23.23

4 4961 15

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Then the complex column whose reboiler duty was 23% less than the conventional configuration

was implementd in the flowsheet and got it converged. The new flowsheet appeared as follows :

Figure 4.5.5 : Flowsheet with the above complex column showed minimum reboiler duty

4.6 Control Structure

Controls were implemented to control flow rates, temperatures, pressures, compositions, levels,

Benzene to ethylene excess ratio and Benzene split ratio at different locations in the flow sheet.

Controls ensure the safe, stable and economical operation of the chemical processes in a plant.

The designing of the control system involved the identification of appropriate positions at which

the control had to be implemented and the type of the control, which depends on the variable to

be controlled. Once the positions were identified and controllers were set up, then parameters for

each controller were adjusted and tuned.

In each controller, the process variable and the controlled variable were specified, input values of

Kc, i and d depending upon the type of controller were defined i.e. Proportion(P), Proportional

Integral(PI) or Proportional Integral Derivative(PID). Then the range of the Process Varible(PV)

was defined. i.e. the range around the set point which the control can operate in. Finally the

controller mode was specified i.e. direct acting or reverse acting.

Reverse mode represents that the contoller action should be opposite to the change in variable.

For example, if we are heating the components in a heater and controlling the temperature of

outlet stream with the heat duty of the heater. So when the temperature of the outlet stream

increases beyond the set point, then we need to reduce the heat duty in order to bring the

temperature close to set point and thus there will be reduction in the control valve opening which

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would restrict the amount of heating fluid to flow. This kind of controller action is called reverse

action.

If we are cooling the system and temperature of the outlet stream rises, we will increase the

amount of coolant flowing through the valve by increasing the valve opening. So with increase in

temperature, valve opening increases and thus the controller is in the direct mode.

Proportional controllers were used to control level and Proportional Integral controllers for rest

of the variables. After defining the controllers, there came the task of tuning them. After running

the simulation in the dynamic mode, all the controllers were carefully observed and the

performance of each of them was analyzed by observing the variations in the PV and opening

percent of the valve (Output%) with time.

If PV was far away from the set point and fluctuating rapidly with time or output % was close to

0 or 100 or fluctuating with time, then this was an indication that the controller was not working

properly and there was a need to tune it. For tuning, the values of Kc and i were varied until all

controllers became stable with PV close to set point and Output % away from the extremes of 0

and 100.

23 | P a g e

Figure 4.6.1: Fully Optimized flowsheet with controls

Table 4.6.1 : The following controls were implemented in the flowsheet-

OBJECT TO

CONTROL

CONTROLLER

TYPE

PROCESS

VARIABLE Kc Ti

RANGE

OF PV

SET

POINT MODE To Control Flow of

Ethylene Flow Inlet Flow Rate 0.5 0.5 0-200 100 Reverse

Excess Ratio Flow

Fresh benzene Flow

rate 0.1 0.5 1.5-2 1.8 Reverse

To Contrlol Inlet

Temperature of

Alkylator Temperature Inlet Feed Temperature 0.5 5 35-55 45 Reverse

Control Exit

temperature from

alkylator Temperature

Outlet Temperature of

Alkylator 4 1 30-80 65 Direct

Temperature of Inlet

Stream to Separator Temperature Inlet Stream Temp 0.5 5 40-60 50 Direct

Separator Level Level Level 2 0 0-100% 50% Direct

Ethane Column Inlet

Pressure Pressure Inlet Pressure 0.5 5 700-850 770 Reverse

Ethane Column Inlet

Temperature Temperature Inlet Temperature 0.5 5 40-100 70 Direct

Ethane Column

Condensor Pressure Pressure Pressure of Vessel 12 0.8 550-800 670 Direct

24 | P a g e

Ethane Column

Condensor Level Level Level 2 0 0-100 85% Direct

Ethane Column

Condensor Temperature Temperature

Temperature of

Condensor 0.1 40 20-60 40 Reverse

Ethane Column Reboiler

Level Level Level 1 0 0-100% 50% Direct

Ethane Column Tray

Temperature Temperature

Tray Temperature

(Tray 4) 0.5 15 140-180 157 Reverse

Recycle Column

Condensor Pressure Pressure Condensor Pressure 5 1 140-300 240 Direct

Recycle Column

Condensor Level Level Level 2 0 0-100 50% Direct

L/F Ratio Controller of

the Recycle Column Flow Reflux Rate 0.5 0.5 .2-.8 0.5425 Reverse

Recycle Column

Reboiler Level Level Level 2 0 0-100 50% Direct

Recycle Column Tray

Temperature Temperature Temp of 18th Tray 0.5 30 110-150 135 Reverse

Product Column

Condensor Pressure Pressure Condensor Pressure 4 1 110-140 120% Direct

Product Column

Condensor Level Level Level 2 0 0-100 50% Direct

Product Column Tray

Temperature Temperature Temp of Tray 27 0.5 8 160-195 172 Reverse

Produt Column

Condensor Reboiler

Level Level Level 2 0 0-100% 64% Direct

Split Ratio Controller(

Benzene Split) Flow

Flow Rate of Bz to

Transalkylator 0.5 0.5 0-200 100% Reverse

Transalkylator Inlet

Temperature Temperature Inlet Temperature 0.4 10 140-180 160 Reverse

4.7 Controller Performance

Variation of Throughput

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Figure 4.7.1 : Variation of EB production Rate with 10% change in throughput

Figure 4.7.2: Variation of ethane vent rate with 10% change in throughput

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Figure 4.7.3 : Varition of DEB recycle rate with change in throughput

Figure 4.7.4: Varition of EB purity in the product with change in throughput

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Figure 4.7.5 : Varition of Benzene flow in the recycle with change in throughput

Variation of Ethane Percentage in the feed

Figure 4.7.6 : Variation of Ethyl Benzene purity with changes in mole percent of ethane in ethylene feed

28 | P a g e

Figure 4.7.7 : Varition of ethane vent rate with changes in mole percent of ethane in ethylene feed

Figure 4.7.8 : Varition of Benzene flow in the recycle with change in mole percent of ethane in feed

Figure 4.7.9 : Varition of DEB in the recycle with change in mole percent of ethane in feed

5 Economics

5.1 Sizing of the Heat Exchangers

Sizing of the equipments play a major role in the economics of the plant. Designers are always in

the search of reducing the size keeping the overall performance safe, stable and efficient. There

is always a question at the back of the mind that can we reduce volume of the reactors? Is there a

way to reduce the number of trays in the column? What should be the heat transfer rates across

heaters and coolers that could limit their size? and so on.

29 | P a g e

For heat exchangers such as heaters, coolers, condensers and reboilers areas were found using

the following formula :

Q = UATlm eq (5.1)

Q : Power supplied to the heater

U : Overall heat transfer rates obtained from literature[4]

A : Area of heater or cooler

Tlm : Log Mean Temperature Difference(LMTD)

For finding LMTD we assumed temperature of steam to remain constant for heater and

temperature of cold water rose from 30C to 40C for cooler. Using these assumptions and inlet

and outlet temperature, we found T1 and T2 and calculated Tlm from the following formula :

= (1 2)

ln (1

2)

..eq(5.2)

5.2 Size, Capital Cost and Operating Cost of the equipments

For finding capital cost of the equipments, we first obtained purchased cost(Cp) using formula :

log10Cp = K1 + K2log10(A) + K3[(log10A)2] ..eq(5.3)

where A is the capacity or size parameter and K1, K2 and K3 are constants

Then Pressure factor(Fp) was found using :

log10Fp = C1 + C2log10(P) + C3[(log10P)2] eq(5.4)

where unit of P is bar gauge (barg) and C1, C2 and C3 are constants.

Pressure factor is used in the overall capital cost because in purchase cost, it is assumed that the

equipment operates at the atmospheric pressure.Material factor(Fm) for different materials are

obtained since purchase cost assumes that the material is made up of carbon steel.

Then we obtained Bare Module Cost Factor(FBM) which accounts for direct and indirect costs

associated with the purchased cost(Cp). Final capital cost or bare module cost of the equipment is

calculated using :

FBM = (B1 + B2FMFP) eq(5.5)

CBM = CPFBM .eq(5.6)

where CBM is the bare module cost of the equipment

Table 5.2.1 : Size, Capital Cost and Operating Cost of all the major equipments

Equipment Size Capital Cost($) Operating Cost($/yr) Cooler before Alkylator 19.55 m2 2.397E+04 4868

Heater before Transalkylator 10.68 m2 3481 3.987E+04

30 | P a g e

Cooler before Ethane Column .35 m2 997 350.1

Heater before separator 71.57 m2 8.492E+04 5683

Alkylator No of Tubes = 50

Length = 10m

Volume = 20 m3

D= 0.22 m

3.677E+06 2.923E+04

Transalkylator Length = 20m

Volume = 5m3

D = .56m

2.534E+04 9162

Ethane Column (Vessel + Trays) No. of Trays = 5

Height = 3.048 m

7242 None

Ethane Column condenser .75 m2 1565 122

Ethane column reboiler .48 m2 2.087E+05 8478

Recycle column(Vessel + Trays) No of Trays = 34

Height = 20.73 m

6.3E+05 None

Recycle column condenser 161.3 m2 1.868E+05 4.991E+05

Recycle column reboiler 196.3 m2 1.592E+05 2.918E+06

Product column(Vessel + Trays) No of Trays = 35

Height = 21.34 m

2.851E+05 None

Product column condenser 33.55 m2 3.762E+04 2.246E+04

Product Column Reboiler 105.4 m2 1.149E+05 1.080E+06

5.3 Objective Function(J) Objective Function(or J function) is defined as follows :

=

Table 5.3.1 : J function calculations

Total Energy Cost $4.635E+06

Total Capital Cost $2.534E+06

Payback Period 3 years

Cost of Benzene $51.25 per kmol

Cost of Ethylene $33.66 per kmol

Selling Price of Ethyl Benzene $113.3 per kmol

Total Raw Material Cost per annum $6.724E+07

Revenue per annum $8.974E+07

J(Profit per annum) $1.702E+07

6 References [1] : Turton, Richard, et al. Analysis, synthesis and design of chemical processes. Pearson

Education, 2008.

31 | P a g e

[2] : Douglas, J. M., Conceptual Design of Chemical Processes (New York: McGraw-Hill,

1989).

[3] : Douglas, J. M., A Hierarchical Design Procedure for Process Synthesis, AIChE Journal, 31 (1985): 353.

[4] : Ebrahimi, Ali Nejad, et al. "Modification and optimization of benzene alkylation process for

production of ethylbenzene." Chemical Engineering and Processing: Process

Intensification 50.1 (2011): 31-36.

[5] : Ganji, Hamid, et al. "Modelling and simulation of benzene alkylation process reactors for

production of ethylbenzene." Petroleum and Coal 46 (2004): 55-63.

[6] : Qi, Zhiwen, and Ruisheng Zhang. "Alkylation of benzene with ethylene in a packed reactive

distillation column." Industrial & engineering chemistry research 43.15 (2004): 4105-4111.

[7] : Tiako Ngandjui, L. M., D. Louhibi, and F. C. Thyrion. "Kinetic analysis of diethylbenzene-

benzene transalkylation over faujasite Y." Chemical Engineering and Processing: Process

Intensification 36.2 (1997): 133-141.