Ethyl Benzene Plant Design

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 1|Page

Table of Contents 1. Acknowledgement…………………………………………………………………………3 2. Executive Summary………………………………………………………………………..4 3. Introduction………………………………………………………………………………...4 3.1 General Philosophy behind design of chemical processes…………………………….4 3.2 Introduction to Ethyl Benzene…………………………………………………………6 3.3 Ethyl Benzene Reaction System……………………………………………………….6 3.4 Reaction Kinetics………………………………………………………………………7 3.5 Commercial methods for production of Ethyl Benzene………………………………..8 4. Process Simulation…………………………………………………………………………9 4.1 Fluid Property Package………………………………………………………………...9 4.2 Process Description……………………………………………………………………11 4.3 Process flow diagram………………………………………………………………….12 4.4 Degrees of freedom……………………………………………………………………13 4.5 Optimization…………………………………………………………………………...13 4.6 Idea for heat integration………………………………………………………………. 17 4.7 Complex Column Configuration………………………………………………………17 4.8 Control Structure………………………………………………………………………21 4.9 Controller Performance………………………………………………………………..24 5. Economics…………………………………………………………………………………27 5.1 Size of the Heat Exchangers…………………………………………………………...27 5.2 Size, Capital Costs and Operating Cost of the equipments……………………………28 5.3 Objective function(J)…………………………………………………………………..29 6. References…………………………………………………………………………………29

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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 cooperation.

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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. 4|Page

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. 2. 3. 4. 5.

Decide whether the process will be batch or continuous. Identify the input/output structure of the process. Identify and define the recycle structure of the process. Identify and design the general structure of the separation system. 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 AlCl 3 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 6|Page

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 → Benzene Ethylene C6H6 Benzene

+ 2C2H4 → Ethylene

…………………….1

C6H4 (C2H5)2 Diethyl Benzene

.....………………….2

C2H4 Ethylene

→

C6H4 (C2H5)2 DEB

…….……………….3

+ C6H5 (C2H5)2 DEB

→

2 C6H5C2H5 EB

……..........................4

C6H5C2H5 EB C6H6 Benzene

C6H5C2H5 EB

+

2 C6H5C2H5 EB

→

C6H6 + Benzene

C6H4 (C2H5)2 DEB

......…………………5

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 doesn’t gives us practical results. The kinetics used is as follows: r1 = 0.084*exp(-9502/RT)CE1.0CB0.32 r2 = 0.603*exp(-15396/RT)CE1.3CB0.33

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r3 = 0.00085*exp(-20643/RT)CE1.77CEB0.35 For reaction 4 and 5, we regressed data for CBZ0 and CDEB0 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 𝐾 1 1 1 ln ( )= ( − ) 𝐾571.15 𝑅 𝑇 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 - 300°C 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 Steady State Operating Degrees of Freedom Steady State Design Degrees of Freedom

43 17 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 13 | P a g e

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.

J vs Excess Ratio

1.80E+07

J (Economic Function)($)

1.70E+07

1.60E+07

1.50E+07

1.40E+07

1.30E+07

1.20E+07 1

1.5

2

2.5

1.8 Excess Ratio

3

3.5

4 J vs…

4.5

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 doesn’t 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.

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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.

J vs Split Ratio

1.75E+07

J (Economic Function)($)

1.70E+07 1.65E+07 1.60E+07 1.55E+07 1.50E+07

1.45E+07 1.40E+07 1.35E+07 1.30E+07 40

65

90

115

Split Ratio

140

165 190 J vs Split Ratio

215

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 doesn’t 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.

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1.80E+07

J vs Transalkylator Temperature

J (Economic Function)($)

1.60E+07 1.40E+07 1.20E+07 1.00E+07 8.00E+06 6.00E+06 4.00E+06 100

150 200 Temperature (oC)

250 300 J vs Transalkylator Temperature

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.

J (Economic Function)($)

1.72E+07

J vs Benzene Leakage

1.70E+07 1.68E+07

1.66E+07 1.64E+07 1.62E+07 1.60E+07 1.58E+07 1.56E+07 1.54E+07 0.00E+00 2.00E-07 4.00E-07 6.00E-07 8.00E-07 1.00E-06 1.20E-06 J vs Benzene… Benzne Leakage

Figure 4.4.4 : Variation of J with benzene leakage from the recycle column 16 | P a g e

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.

Without Heat Integration Heat Integration

Energy Consumption Savings 6181kW 1283kW $2.16 million

0

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

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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 doesn’t 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, τi and τd 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 21 | P a g e

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 τi were varied until all controllers became stable with PV close to set point and Output % away from the extremes of 0 and 100.

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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

To Control Flow of Ethylene

Flow

0.5

0.5

0-200

100

Reverse

Flow

Inlet Flow Rate Fresh benzene Flow rate

0.1

0.5

1.5-2

1.8

Reverse

Temperature

Inlet Feed Temperature

0.5

5

35-55

45

Reverse

Temperature

Outlet Temperature of Alkylator

4

1

30-80

65

Direct

Temperature Level

Inlet Stream Temp Level

0.5 2

5 0

40-60 0-100%

50 50%

Direct Direct

Pressure

Inlet Pressure

0.5

5

700-850

770

Reverse

Temperature

Inlet Temperature

0.5

5

40-100

70

Direct

Pressure

Pressure of Vessel

12

0.8

550-800

670

Direct

Excess Ratio To Contrlol Inlet Temperature of Alkylator Control Exit temperature from alkylator Temperature of Inlet Stream to Separator Separator Level Ethane Column Inlet Pressure Ethane Column Inlet Temperature Ethane Column Condensor Pressure 23 | P a g e

RANGE SET Kc Ti OF PV POINT MODE

Ethane Column Condensor Level Ethane Column Condensor Temperature Ethane Column Reboiler Level Ethane Column Tray Temperature Recycle Column Condensor Pressure Recycle Column Condensor Level L/F Ratio Controller of the Recycle Column Recycle Column Reboiler Level Recycle Column Tray Temperature Product Column Condensor Pressure Product Column Condensor Level Product Column Tray Temperature Produt Column Condensor Reboiler Level Split Ratio Controller( Benzene Split) Transalkylator Inlet Temperature

Level Temperature Level

2

0

0-100

85%

Direct

0.1

40

20-60

40

Reverse

1

0

0-100%

50%

Direct

Temperature

Level Tray Temperature (Tray 4)

0.5

15

140-180

157

Reverse

Pressure

Condensor Pressure

5

1

140-300

240

Direct

Level

Level

2

0

0-100

50%

Direct

Flow

Reflux Rate

0.5

0.5

.2-.8

0.5425

Reverse

Level

Level

2

0

0-100

50%

Direct

Temperature

Temp of 18th Tray

0.5

30

110-150

135

Reverse

Pressure

Condensor Pressure

4

1

110-140

120%

Direct

Level

Level

2

0

0-100

50%

Direct

Temperature

Temp of Tray 27

0.5

8

160-195

172

Reverse

Level

2

0

0-100%

64%

Direct

Flow

Level Flow Rate of Bz to Transalkylator

0.5

0.5

0-200

100%

Reverse

Temperature

Inlet Temperature

0.4

10

140-180

160

Reverse

4.7 Controller Performance Variation of Throughput

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Level Temperature of Condensor

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

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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.

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For heat exchangers such as heaters, coolers, condensers and reboilers areas were found using the following formula : Q = UAΔTlm

…………………eq (5.1)

Q : Power supplied to the heater U : Overall heat transfer rates obtained from literature[4] A : Area of heater or cooler ΔTlm : 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 30°C to 40°C for cooler. Using these assumptions and inlet and outlet temperature, we found ΔT1 and ΔT2 and calculated ΔTlm from the following formula :

∆𝑇𝑙𝑚 =

(∆𝑇1− ∆𝑇2)

…………………..eq(5.2)

∆𝑇1 ) ∆𝑇2

ln(

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 Cooler before Alkylator Heater before Transalkylator 29 | P a g e

Size

Capital Cost($) Operating Cost($/yr) 2

19.55 m 10.68 m2

2.397E+04 3481

4868 3.987E+04

Cooler before Ethane Column Heater before separator Alkylator

Transalkylator

Ethane Column (Vessel + Trays) Ethane Column condenser Ethane column reboiler Recycle column(Vessel + Trays) Recycle column condenser Recycle column reboiler Product column(Vessel + Trays) Product column condenser Product Column Reboiler

.35 m2 71.57 m2 No of Tubes = 50 Length = 10m Volume = 20 m3 D= 0.22 m Length = 20m Volume = 5m3 D = .56m No. of Trays = 5 Height = 3.048 m .75 m2 .48 m2 No of Trays = 34 Height = 20.73 m 161.3 m2 196.3 m2 No of Trays = 35 Height = 21.34 m 33.55 m2 105.4 m2

997 8.492E+04 3.677E+06

350.1 5683 2.923E+04

2.534E+04

9162

7242

None

1565 2.087E+05 6.3E+05

122 8478 None

1.868E+05 1.592E+05 2.851E+05

4.991E+05 2.918E+06 None

3.762E+04 1.149E+05

2.246E+04 1.080E+06

5.3 Objective Function(J) Objective Function(or J function) is defined as follows : 𝐽 = 𝑅𝑒𝑣𝑒𝑛𝑢𝑒 − 𝑅𝑎𝑤 𝑀𝑎𝑡𝑒𝑟𝑖𝑎𝑙 𝐶𝑜𝑠𝑡 − 𝐸𝑛𝑒𝑟𝑔𝑦 𝐶𝑜𝑠𝑡 −

𝐶𝑎𝑝𝑖𝑡𝑎𝑙 𝐶𝑜𝑠𝑡 𝑃𝑎𝑦𝑏𝑎𝑐𝑘 𝑝𝑒𝑟𝑖𝑜𝑑

Table 5.3.1 : J function calculations

Total Energy Cost Total Capital Cost Payback Period Cost of Benzene Cost of Ethylene Selling Price of Ethyl Benzene Total Raw Material Cost per annum Revenue per annum J(Profit per annum)

$4.635E+06 $2.534E+06 3 years $51.25 per kmol $33.66 per kmol $113.3 per kmol $6.724E+07 $8.974E+07 $1.702E+07

6 References [1] : Turton, Richard, et al. Analysis, synthesis and design of chemical processes. Pearson Education, 2008.

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[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 diethylbenzenebenzene transalkylation over faujasite Y." Chemical Engineering and Processing: Process Intensification 36.2 (1997): 133-141.

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