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Process Design Project 2007 Group 7

Process Design of anAcrylic Acid Plant

Group 7

Group Members:Jamillah David

Luke ElliottNorwind Khor

Oluwatoyin Olaleye

Jason SharpBasel SiddiqiZengcun Zhu

Process Design of an Acrylic Acid Plant........................................................................................................... 1

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Abstract......................................................................................................................................................... 4Level 0) Input Information ..............................................................................................................................4Level 1) Batch or Continuous Operation .......................................................................................................4

1.1) Production Rate ..................................................................................................................................51.2) Market Forces ....................................................................................................................................51.3) Operational Issues ..............................................................................................................................5

Level 2) Input-Output Flowsheet Structure ....................................................................................................5

2.1) Feed Stream Purification ....................................................................................................................52.2) Recycle by-products ...........................................................................................................................52.3) Gas Recycle and Purge ......................................................................................................................52.4) Number of Product streams ................................................................................................................52.3) Economic Potential............................................................................................................................. 62.4) Alternative Designs ............................................................................................................................6

Level 3) Recycle Structure ............................................................................................................................63.1) Number of reactors .............................................................................................................................63.2) Number of Recycle Streams ...............................................................................................................63.3) Excess Reactant................................................................................................................................. 73.4) Material Balances ...............................................................................................................................73.5) Reactor Heat Effects ..........................................................................................................................73.6) Compressor Design and Cost............................................................................................................ 8

3.7) Reactor Cost...................................................................................................................................... 83.8) Design Variable Optimisation .............................................................................................................9Level 4) General Separation System Structure ...........................................................................................10

4.1) General Structure ............................................................................................................................104.2) Vapour Recovery System .................................................................................................................104.3) Liquid Separation System .................................................................................................................10

5) Detailed Simulation using ASPEN ...........................................................................................................125.1) ASPEN Property Models ..................................................................................................................135.2) Process Optimisation ........................................................................................................................13

5.3) Process Adjustments ...........................................................................................................................156) Level 4 Economic Potential (EP4) ...........................................................................................................15

6.1) Capital cost of the separation system ...............................................................................................156.2) Operating cost of the separation system ..........................................................................................16

7) Heat Exchanger Network Synthesis ........................................................................................................167.1) Arrangement of Reactant Preheating ...............................................................................................167.2) Identification of Streams that can be Integrated ...............................................................................167.3) Network Synthesis ............................................................................................................................177.4) Integration of Distillation Columns and Process Streams .................................................................187.5) Final Heat Exchanger Network .........................................................................................................18

8) Level 5 Economic Potential (EP5) ...........................................................................................................189) Process Sensitivity Analysis ....................................................................................................................1810) Environmental considerations ...............................................................................................................19References ..................................................................................................................................................20Appendix A Selectivity Data .....................................................................................................................21Appendix B Data Tables ...........................................................................................................................22Appendix C Material balances required at level 2 ....................................................................................23

Appendix D Material balances required at level 3 ....................................................................................24Appendix E Hesss Law Calculation .........................................................................................................25Appendix F Pressure Sensitivity Analysis ................................................................................................27Appendix G Selection of Column Type .....................................................................................................28................................................................................................................................................................... 28Appendix H Sample Calculation for Costing a Separation System Vessel ............................................... 29

Required Information Obtained from Aspen ............................................................................................29Installed Cost........................................................................................................................................... 29Operating Cost........................................................................................................................................ 30

Appendix I Sizing the Flash Vessel.......................................................................................................... 32Appendix J Heat Transfer Co-efficient Selection ......................................................................................33

Temperature Interval Grid ........................................................................................................................34Energy Cascade ......................................................................................................................................35

Minimum Number of Units for Maximum Energy Recovery .....................................................................35Appendix L EP4 Price Break-Down .........................................................................................................36Appendix M EP5 Price Break-Down .........................................................................................................37Appendix N ASPEN .................................................................................................................................39

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AbstractThis report outlines the details of designing a process to produce 50,000 tonnes.yr-1 of acrylic acid with amass purity greater than 95%. It was firstly decided that a continuous process should be used. Next, a level2 analysis was carried out, where it was decided to employ a gas recycle and purge. The level 2 economicpotential was found to be in the region of 85106$.yr-1, indicating that this chemical route has a high potentialto be profitable. Carrying this design through to level 3 it was found that the optimum conversion and

percentage loss of propylene in the purge was 0.85 and 5% respectively. Feed mass fractions of 0.09propylene and 0.275 steam were chosen to maximise economic potential while maintaining safe operation.Using design heuristics, the general separation system was decided to consist of a flash vessel, vapourrecovery (absorber using demineralised water) and a liquid separation system. Studying calculated binarymixture ratios the liquid separation system was designed and consisted of a liquid-liquid extractor and twodistillation columns. Each of these units operating conditions were determined through optimisation usingASPEN. Finally, a heat integration network was designed to reduce the processes utility costs. The level 5economic potential of the designed process was estimated at 80.1106 $.yr-1.

Level 0) Input InformationAcrylic acid is an industrially important compound that is mainly used as an intermediate in the formation of awide variety of products such as plastics and paints. A plant was to be designed with the aim of producing

acrylic acid (C3H4O2) and the input information is provided below.

0.1) Reactions and conditions

a) Reactions: )(2)(243)(2)(63 5.1 gggg OHOHCOHC ++

[R1]

)(2)(2)(242)(2)(63 5.2 ggggg OHCOOHCOHC +++

[R2]

)(2)(2)(2)(63 335.4 gggg OHCOOHC ++ [R3]

b) Reactor conditions: Inlet Temperature = 310C and pressure = 3.5 bar.c) Phase of reaction system: Gas.d) Catalyst: The number of catalysts required is discussed in level 3 section 3.1.

e) Product distribution: Selectivity data as a function of conversion is provided in Appendix A.f) Kinetic information: At the reactor conditions the equilibrium of reactions R1, R2 and R3 lies far to

the right. Each reaction rate, r, can be evaluated using equation 1 and the values forEr and ko,r aretabulated in Appendix B, Table 8.

2exp, Oprop

R

r

ror PPRT

Ekr

=

[1]0.2) Required rate of production for acrylic acid: 50,000 metric tons per year.0.3) Required purity of acrylic acid: minimum 95% purity by mass.0.4) Raw materials specifications: Tabulated in Table 1.

Table 1: Raw material specifications.

Raw

Material

Mass Composition

(%)

Available

Pressure (bar)

Available

Temperature (oC) Price ($/kg)Propylene 6-9 3.50 -16.3 0.38Steam 25-30 3.50 310 139.36 10-4

Air 61-69 1.01 25.0 0.00

0.5) Processing constraints: the feed mass composition must be within 6 9% propylene, 25 30% steamand the remainder air to maintain operation outside the explosive limits of the combustion reactions.

Level 1) Batch or Continuous OperationTo determine whether batch or continuous operation should be applied to this process the following areashave been considered.

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1.1) Production Rate

Continuous operation is preferred to batch configuration for plants that are required to produce more than5000 tonne.yr-1 because batch processes have increased operating costs at higher production levels. In thiscase, 50,000 tonne.yr-1 of product is required, therefore continuous operation is favoured.

1.2) Market Forces

Acrylic acid has a stable demand throughout the year due to its wide use in industrial applications. The plantbeing designed is only required to produce a single product that has long term uses and demand. Thereforethe flexibility of a batch process is not required and continuous operation is more suitable.

1.3) Operational Issues

There are no operational issues that make a continuous process inappropriate. This is because the processhas a relatively high rate of conversion and does not involve the handling of solids and/or slurries. However,care must be taken in the selection of vessel and piping materials to counter the corrosive properties of thechemicals within the system which may weaken structures and result in solid deposits.

Level 2) Input-Output Flowsheet StructureThe level 2 flowsheet can be seen in Figure 1and the decisions made have been explained below.

Figure 1: Level 2 flowsheet

2.1) Feed Stream Purification

No feed purification is required as the propylene and air streams are pure. The oxygen feed is supplied to the

process as an air stream and hence contains a large proportion of inert nitrogen. It has been decided thatnitrogen will not be removed from this stream due to the high separation costs. In addition, the presence ofnitrogen reduces the amount of steam feed required to operate safely outside the explosive limits of theexothermic combustion reactions and may be utilised as a heat carrier.

2.2) Recycle by-products

At the reactors operating conditions the equilibrium of each reaction lies far to the right, effectively makingthem irreversible. Therefore, there is no benefit from recycling acetic acid (which has a monetary value) andother by-products. As water is inexpensive and may contain impurities it will not be recycled.

2.3) Gas Recycle and Purge

In keeping with the level 2 rule of recovering or using more than 99% of valuable reactants, it was decided

that propylene should be recycled. As propylene has a boiling point of -48o

C it is the boarder line betweenheavy and light components. In order to separate propylene from the lighter components exiting the reactor,distillation with refrigeration or a membrane process would be required. As refrigeration is very expensiveand no cost correlations are available for membrane separators, it was decided to consider propylene as alight component. Thus, a gas recycle and purge was utilised to recover the valuable reactant and prevent theaccumulation of light inert components and by-products in the process.

2.4) Number of Product streams

Every components boiling point and destination code can be seen in Table 2. Assuming distillation is theprimary separation process, product streams were determined by ordering the components by boiling pointand grouping neighbouring components with the same destination code. This results in four product streams,which can be seen in Figure 1.

Table 2: Component destination codesComponent Stream Number Normal Boiling Point (oC) Destination Code

Nitrogen (N2) 8 -195.8 2. Recycle and Purge

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Oxygen (O2) 8 -183.0 2. Recycle and PurgeCarbon Dioxide (CO2) 8 -78.5 2. Recycle and PurgePropylene (C3H6) 8 -48.0 2. Recycle and PurgeWater (H2O) 7 100.0 6. Excess WasteAcetic Acid (C2H4O2) 6 118.1 8. Valuable by-productAcrylic Acid (C3H4O2) 5 141.0 7. Primary Product

2.3) Economic Potential)/($2 yrstscomaterialrawvalueproductbyvaluesproductEP += [2]

( ) 1001

99

6363 =

F

yFLHCofLossPercentage

HC

p [3]

The economic potential at level 2 was calculated using equation 2 and the component prices shown inAppendix B. To determine the flowrates of the valuable components and raw materials in terms of the designvariables material balances were derived for the system assuming perfect separation, which can be seen in

Appendix C. The design variables at this level areconversion and the feed mass fractions of steam andpropylene. The selectivity for acrylic acid and acetic acid in

reactions R1 and R2 have been determined as a function ofconversion by fitting a polynomial to the providedexperimental data, which can be seen in Appendix A. Asthere is uncertainty in extrapolating beyond the range of theselectivity data, only conversions between 0.60 and 0.85will be considered. In order to use more than 99% of thevaluable reactant (propylene) a percentage loss of 1% inthe purge was defined for this component as shown byequation 3. Figure 2 shows the resulting plot of economicpotential against conversion. From this it can be seen thatthe most desirable conversion at this level is the lowestvalue of 0.6, since this results in the largest economicpotential. Most importantly Figure 2 shows that the products

and by-products are worth more than the raw materials, sothis design and chemical route will be progressed to level 3.Figure 2: Economic Potential at level 2 for

Lp = 1% versus conversion

2.4) Alternative Designs

a) Remove propylene from the light gases and recycle this as a pure component. This would requirethe cost of membrane separation to be lower than the loss of propylene in the purge.

Level 3) Recycle Structure

Figure 3: Level 3 flowsheet

3.1) Number of reactors

In industry a two step process is commonly used to produce acrylic acid from propylene as it has a higherconversion and yield than a one step process. This process requires two reactors as each reaction takesplace over different catalysts and at different operating conditions. For simplicity it was assumed that thedesired conversion and yield can be achieved in a single step process where propylene is converted intoacrylic acid using one catalyst. Thus, only one reactor is required for the purpose of this design project.

3.2) Number of Recycle Streams

Since there is only one reactor N2, O2, CO2 and C3H6 all have the same destinations. As these also haveboiling points in the same range, shown in Table 2, they will be grouped, resulting in one gaseous recyclestream that will require a gas compressor.

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where: =iF molar flowrate of stream i =ijy mole

fraction of component j in stream i

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3.3) Excess Reactant

An excess of oxygen will be used because air is relatively inexpensive and it can be seen from equation 1that increasing the partial pressure of oxygen will increase the rate of all reactions and hence the conversionof propylene. It can also be seen that this would not affect the selectivity as the rates of all three reactionsare directly proportional to the partial pressure of oxygen. Furthermore, increasing the flowrate of air to theprocess will also increase the amount of cooling within the reactor by heat carriers.

3.4) Material BalancesAt level 3 an estimate of the recycle and internal flows are required to size the reactor and compressors,which are needed to calculate the economic potential (EP3). In the absence of a designed separation system,perfect separation was assumed for this level 3 analysis. The material balances that have been performedcan be seen in appendix D. In addition to the design variables listed at level 2 (conversion (x) and the massfractions of steam and propylene in the feed), the percentage loss of propylene in the purge will also beconsidered. This is because the percentage loss is directly related to the gas recycle flowrate and thereforeboth the compressors and reactors size and cost. Hence, at level 3 the economic potential for a small lossof valuable reactant may not be as high as a larger loss due to the fact that a larger compressor and reactorwill be required.

3.5) Reactor Heat Effects

In order to examine the reactors heat effects an energy balance was performed using equation 4 . As theheats of each reaction were provided at 298K, Hesss law was used to calculate the enthalpy change acrossthe reactor and this calculation can be seen in Appendix E.

0=+ QH [4] where: =H enthalpy change and Q heat removed

3.5.1) Adiabatic Temperature Change

Using Maple and the Hesss law calculation the reactoroutlet temperature was determined in terms of thedesign variables such that equation 4 was satisfied forQ = 0. For adiabatic operation the temperature rise mustbe less than 10 to 15% of the inlet temperature. Figure 4

shows that operation at low percentage propylenelosses in the purge and lower conversion results infeasible adiabatic temperature rises. However,economic analysis has shown that such operatingconditions are unfavourable, due to high compressioncosts. As the adiabatic temperature rise is unacceptablyhigh for greater propylene losses and conversions it isnot economically desirable to operate the reactoradiabatically.

Figure 4: Adiabatic Temperature Change for different conversions and losses of propylene.

3.5.2) Heating and CoolingTo determine the required cooling duty for isothermal operation equation 4 has been solved for a reactor exittemperature of 310C. This results in a cooling duty of 15000kW for a 1% loss of propylene in the purge andconversion of 0.85. As this is greater than the maximum heat removal of 2550kW (~8106 Btu/hr) by coolingwater (Douglas, 1988) isothermal operation is not feasible.

3.5.3) Heat CarriersThe inert steam added to the reactor to ensure operation outside of the explosive limits is also a dilutant andheat carrier. An additional option to moderate the reactor temperature rise is to increase the gas recycleflowrate. This results in a higher flow of heat removing inert components through the reactor. However, aspreviously discussed in section 3.5.1 this results in high compressor costs which are economicallyundesirable. Therefore, this option is also not favourable.

As all of these operation modes are either unfeasible or undesirable one of the design rules will have to becompromised. By cooling the reactor with a molten salt it would be possible to remove 15000kW of heat.Thus, it has been decided to use this form of cooling to operate the reactor isothermally. Cost information for

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molten salts is not available so the cost of cooling water required was used instead to estimate the reactorsannual cooling cost.

3.6) Compressor Design and Cost

Two compressors are required for this process, one for the gas recycle and another for the inlet air stream.It has been decided to use centrifugal compressors as these have few moving parts, high energy efficiencyand a small size relative to other types of compressors. The installed cost of these were calculated using

Guthries correlation (equation 5) where the Marshall and Swift equipment cost index (M & S) for this year is1094 and Fc= 1.00 for centrifugal compressors (Douglas, 1988). The installed cost was converted to anannualised cost by dividing by a capital charge factor, which is related to the project life, cost of capital andhas been estimated as 3.

( )( ) ( )CFbhp

SMCostInstalledCompressor +

= 11.25.517

280

&($)

82.0[5]

=

11003.3 5

in

out

ininP

PQPhp [6]

The brake horse power (bhp) was determined by dividing the isentropic power requirement (hp) by 0.9. This

allows for irreversibilities, such as fluid friction in valves, friction of moving parts and fluid turbulence. Thehorse power (hp) was calculated using equation 6 and in doing so it was assumed that there is no pressuredrop across the mixer and heater preceding the reactor, such that Pout = 3.5 bar (7308 lb.ft -2) for bothcompressors. For the inlet air compressorPin is specified as 1.01 bar (2116 lb.ft -2). However, for the gasrecycle compressorPinwas assumed to be 0.8Pout to account for pressure drops across the reactor andseparation system, as this has yet to be designed. A sensitivity analysis, which can be seen in appendix Findicates that the total annual cost of the gas recycle compressor varies significantly with the pressure dropselected. Thus, if the gas recycle streams outlet pressure from the separation system determined at level 4is notably different from 0.8Pout the cost of this compressor will need to be re-evaluated.

The molar flowrate of both streams being compressed were converted into volumetric flowrates assumingperfect gas behaviour, where the temperature of the gas stream exiting the separation system was estimatedas 40C. was estimated by taking a molar weighted average of appropriate values for each component,

which were selected from Table 10, appendix B.

The compressors annual operating cost was determined by multiplying their power requirement andelectricity cost (0.04 $.hp-1.hr-1). The power requirement is equal to the bhp divided by 0.9 to allow forinefficiencies in converting input energy to shaft work.

3.7) Reactor Cost

The installed cost of the reactor is mainly dependent on its dimensions, as shown by Guthries costcorrelation for pressure vessels (equation 7). Therefore the reactor needed to be designed in terms of thedesign variables and the procedure by which this was done is outlined below.

( )

( ) ( )cFLD

SM

CostInstalledeactorR += 18.29.101280&

($)

802.0066.1

[7]

It was decided to use a tubular reactor for this process as these operate at a high reaction rate, resulting in alower volume to achieve the same conversion as other types of reactors. Furthermore, tubular reactors areusually more suitable for gas phase reactions. Due to the presence of H 2O and acidic chemicals within thereactor it will be made out of stainless steel to counter corrosion.

When designing the reactor it was assumed that the total molar flowrate remained constant as the flowrate ofinert materials, such as H2O and N2, through it is much larger than that of the limiting reactant. Using thissimplification the reactor volume for a desired conversion was calculated by modelling it as a series ofCSTRs using an iterative procedure in Maple. The conversion of each CSTR was fixed at dx = 0.01,maintaining a sufficient level of accuracy and the volume of each of these was calculated using equation 8,

where the rate of each reaction was evaluated at the exit conditions of the CSTR. This iteration was repeateduntil dx i was equal to the required conversion and the sum ofdVi evaluated to calculate the correspondingreactor volume. Assuming 6/ =DL , the reactors length, diameter and hence installed cost (using equation7) could be found in terms of the design variables. As the reactor operates at a pressure of 3.5 bar (50.75

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Where: P = Pressure (lb.ft -2)Q = Volumetric flowrate (ft3.min-1)

Where: pmc FFF = ,

D = reactor diameter (ft)L = reactor length (ft)

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lb.in-2) and is made out of stainless steelFm= 3.67 andFp= 1.05 (Douglas,1988). Similarly to section 3.6 thiscost was converted to an annualised cost by dividing by a capital charge factor of 3.

( )

idxx

p

irrr

dxyFdV

=++

=321

44

[8] (Derived from the CSTR design equation (Metcalfe, 1997))

The cooling cost of the reactor was estimated by calculating the flow rate of cooling water required to remove

the heat necessary for isothermal operation. By performing an energy balance on the cooling water streamusing the maximum cooling water temperature change (Tmax=15C) and its heat capacity (cp,w= 4.2 kJ.kg-1.C-1) the cooling water flowrate was determined. By multiplying this by the cooling water utility cost,tabulated in appendix B, the annual cooling cost in terms of the design variables was discovered.

3.8) Design Variable Optimisation

Generally design variables are not all fixed at level 3 and are retained until level 4. However, it has beendecided to optimise the design variables at this stage as time constraints make it difficult to do this at level 4.The inlet mass fractions of propylene and steam were optimised first, followed by the percentage loss ofpropylene in the purge and conversion simultaneously. This approach was taken so that the factors affectingthe economic potential least were optimised first.

Figure 5: EP3 for different conversions and massfraction of propylene in the feed

Figure 6: EP3 for different conversions and massfraction of steam in feed

Figure 5 clearly shows that the level 3 economic potential (EP3), which was calculated using equation 9,increases as the inlet mass fraction of propylene is increased. Therefore the optimum inlet mass fraction ofpropylene is the maximum value of 0.09. In contrast EP3 is insensitive to changes in the inlet mass fraction ofsteam, which can be seen in Figure 6. Since the mass fraction of propylene chosen is at the upper limit, asteam mass fraction at the lower limit was avoided to ensure an adequate error margin from the explosivelimits. Thus, a mid-point inlet steam mass fraction of 0.275 was selected.

)cos()cos(23 toperatingandcapitalreactortoperatingandcapitalcompressorEPEP = [9]

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Figure 7 displays EP3 as a function of conversion forvarious percentage losses of propylene in the purgeat the previously optimised inlet mass fractions. Thisfigure clearly indicates that the highest EP3 isproduced by an optimum conversion of 0.85 and apropylene loss of 4 or 5%. From closer examinationof this plot it was found that the optimum percentage

loss of propylene is 5%. These optimum designvariables result in a level 3 economic potential of83106 $/yr.

Figure 7: EP3 versus conversion for different Lp

Level 4) General Separation System Structure

4.1) General Structure

The stream leaving the reactor is at 310C and is gaseous. From design heuristics this stream is to becooled to 35C and a phase split attempted. The outlet vapour is sent to a vapour recovery system and thecondensed liquid to a liquid recovery system. It has been decided to cool the stream exiting the reactor to45C, as cooling water is available at 30C and a minimum approach of 10C must be maintained.

4.1.1) FlashAs the stream exiting the reactor contains water and acidic compounds with dipoles, non-ideal behaviourwould be exhibited. Therefore the short cut method is not applicable for this situation and the flash vesselwas modelled in ASPEN. The results from this can be seen in Appendix N. It was decided to operate the

flash vessel adiabatically at a pressure of 2.8bar. This was because lower pressures increase the costs ofcompressing the gas recycle and higher pressures are unfavourable for the subsequent separation units,which operate more effectively at lower pressures. A joule-thompson valve will be added downstream tofurther reduce the pressure to atmospheric. It was found that flashing to 2.8bar is a reasonable compromise,which provides an acceptable vapour-liquid split.

4.2) Vapour Recovery System

4.2.1) LocationUsing Aspen it has been found that a purge with a 5 percent loss of propylene results in acrylic and aceticacid losses of 1.48106 $/yr. This is significant and merits the use of a vapour recovery system on the purge.Although no materials in the gas recycle affects the reactor operation or degrade the product distribution, thepresence of acidic components may cause corrosion. In addition, not using a vapour recovery system wouldchange the level 3 material balances and the previously optimised design variables. Due to time constraintsre-evaluation of the material balances was not possible and therefore a vapour recovery system wasrequired to meet the specified acrylic acid production rate. Since the purge and gas recycle both require avapour recovery system, a single system was placed on the flash outlet due to economies of scale.

4.2.2) TypeThe only vapour recovery processes with available cost information in this project are condensation andabsorption with demineralised water. Other currently available technologies are adsorption, reaction systemsand membranes. As acrylic and acetic acid have an infinite solubility in water (Perry et al., 1997), absorptionis the more effective process and will be used. Using Table 16 in Appendix G it has been decided to usesieve trays in all columns designed for this process. Sieve trays were also chosen as these have a similarefficiency to bubble-cap trays but are simpler and cost 30 to 50% less (King,1980). Stainless steel is thechosen design material for all columns due to the corrosive properties of acrylic and acetic acid in thepresence of water.

4.3) Liquid Separation System

4.3.1) Removal of light components

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As Acrylic acid has a higher boiling point than all other components leaving the flash vessel it is expectedthat the light impurities would not significantly contaminate this product, assuming distillation is used in theliquid separation system. However, even if all the light components leaving the flash vessel remain in theacrylic acid product stream it would contain 96% acrylic acid (using and neglecting the vapour recovery). Asthis is greater than the product specification it is not necessary to remove the light ends.

4.3.2) Azeotropes and reactants

Propylene forms an azeotrope with water. This is not problematic since these components are mixed at thereactor inlet and thus separation is not important. The acetic acid-acrylic acid-water mixture is azeotropic innature; a feature which must be considered in the liquid separation system design.

4.3.3) Structuring the Liquid Separation SystemWith the cost information provided in this project the processes that are available for separating componentsin a liquid phase are flashing, distillation and liquid-liquid extraction with di-isopropyl ether (DIPE). The otheralternatives that have not been considered are crystallisation, ion exchange, liquid membranes andmicrofiltration. To identify which technique and separation sequence is most suitable, binary property ratios(rij) have been computed for each pairing of the three liquids required to be separated (H2O, C3H4O2,C2H4O2). The results from this are displayed in Table 3 and the properties used to compute these ratios canbe seen in appendix B. The table also displaysik (defined on the next page) and the higher this value is, thebetter the separation.

Table 3: Binary mixture ratios and separation analysis of different processesSeparation Process (k) kgr kfr Property (j) jr1 jr2 jr3 k1 k2 k3

Distillation (1atm.) 1.02 1.01 b.p. 1.11 1.05 1.06 10 3.8 4.9Distillation (1atm.) 1.50 1.05 Pvap 5.08 1.62 3.13 8.94 1.26 4.65Flash (1atm.) 1.40 1.23 b.p. 1.11 1.05 1.06 NA NA NAFlash (1atm.) 15.0 10.0 Pvap 5.08 1.62 3.13 NA NA NALiq-liq Exraction(42.22C, 2.5 bar) 2.0 1.2 solubility 100 135.6 1.35 123.5 168 0.188(62.08C, 2.5 bar) 2.0 1.2 solubility 38.4 65.4 1.70 46.5 80.2 0.626

Pair 1 = H2O-C3H4O2 Pair 2 = H2O-C2H4O2 Pair 3 = C2H4O2-C3H4O2

where:jB

jA

ij

p

pr =

kfkg

kfij

ikrr

rr

=

Table 3 indicates that each component pair can be separated by distillation or liquid-liquid extraction withDIPE. However, acrylic acid polymerises at temperatures greater than 90C, which is undesirable. As allthree components have normal boiling points greater than 90C atmospheric distillation would not be feasibleand vacuum distillation may need to be used. However, for the purpose of this project the polymerisationreaction has been neglected for simplicity. Table 3 also shows that water can be separated from both acrylicand acetic acid more effectively using liquid-liquid extraction with DIPE than distillation. This higheffectiveness justifies the cost of the solvent and will therefore be used rather than distillation. Furthermore,separating the acids from water requires azeotropic distillation, which will be much more complicated and

expensive due to the necessary use of multiple columns. As this separation has the largest value of,indicating it is the easiest, it will be performed first. In addition, water is corrosive which also favours its earlyremoval.

Table 3 shows that separating acrylic and acetic acid by liquid-liquid extraction with DIPE is worse thangood and distillation is more effective. As distillation also does not require a solvent it is generally cheaperthan liquid-liquid extraction and was therefore used for this separation. However, the DIPE must also beseparated from the acids. This was performed by distillation as it was the only process available for thisseparation. In addition, a difference in boiling points and vapour pressures exist (see appendix B, Table 12).As the boiling point and vapour pressure difference between the DIPE and both acids is greater than that ofthe two acids, removing DIPE is the easiest separation. Since this is also the most plentiful component andlightest it was removed first. This conflicts with the rule of removing products as distillates but it agrees withthe rule of removing recycles as distillates as DIPE was recycled.

Thus, the liquid separation system consists of a liquid-liquid extractor using DIPE to separate water from theflash liquid condensate, followed by a distillation column to remove DIPE from the acrylic and acetic acid(DISTIL1). Then a second distillation column separates the acrylic and acetic acids (DISTIL2).

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pjA or B= value of propertyj for component A or B, where mixture iis made up of A and B

ifik> 1 separation is better than good

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5) Detailed Simulation using ASPEN

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Figure 8: Final Process Flowsheet

DIPE TOP-UP

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The full stream table can be seen in Appendix N and specified values have been highlighted.

5.1) ASPEN Property Models

One key element of the Aspen simulation was the base and property methods which were used tocharacterise the fluid behaviour. Through accurately modelling the process streams, precision in the choiceof operating conditions and vessel design was improved. Owing to the presence of the polar components,namely water, acetic acid and acrylic acid, the system exhibited a high degree of non ideal behaviour. The

chosen base method was able to accurately model these particular stream characteristics. Additional fluidproperty consideration had to be made in each separation vessel and a complimentary property methodchosen to best replicate the conditions that the solution would face in a given vessel. The chosen basemethod was NRTL-RK and the property models employed were UNIF-LL, UNIFAC and UNIF-DMD in theliquid-liquid extractor, absorber and distillation columns respectively.

NRTL-RK: NRTL-RK is an activity coefficient method used to represent highly non-ideal liquid mixtures atlow pressures (less than 10 atm, which applies for this process). R.K. which stands for the Redlich Kwongequation of state is able to calculate the vapour phase thermodynamic properties of the system provided thatthe vapour phase non-ideality is small. This property is important for the modelling of N2, O2 and C3H6, whichwere the main components of the vapour phase. The NRTL component of the model describes vapour -liquid equilibrium (VLE) and liquid-liquid equilibrium (LLE) of strongly non-ideal solutions. Thus, it is suitablefor modelling the acidic solution which, owing to its polar nature, exhibits a high degree of non-ideality. The

model requires binary parameters such as equilibrium phase data, many of which are included in the AspenPhysical Property System databanks. The main disadvantage of this model is that it is not predictive; thebinary parameters are only valid over the temperature and pressure ranges under which they were collected.

UNIFAC: UNIFAC is used as a predictive model for the mixtures behaviour in the event that the binaryparameters for the system are not included in the Aspen databanks. It is particularly accurate at modellingVLE data. This feature is useful in the absorber since the system involves the dissolution of the acidic gasesin the water where accurate simulation of the acid gas into the water is vital.

UNIF-LL: UNIF-LL is based on the previously described UNIFAC property method and is able to accuratelymodel LLE data. This property is particularly important in the liquid-liquid extraction column for the accuratemodelling of the interactions between the organic solvent and acidic solution.

UNIF-DMD: Also based on the UNIFAC Method, it contains more temperature dependent terms of the group-group interaction parameters. It is also able to predict both VLE and LLE with a single set of parameters andto predict heats of mixing better. Thus, it is important in the distillation columns and flash vessel where bothVLE and LLE interactions are important and separation is based on relative volatility.

5.2) Process Optimisation

5.2.1) Vapour Recovery SystemIf this unit was optimised solely on its owneconomics then the best solution would be to use asmall column with a large flowrate of demineralisedwater. However, using large amounts of water atthis stage will lead to increased costs in

subsequent separation units; due to the increasedsize and solvent flowrate required to remove thisadditional water. Therefore, the column wasoptimised to achieve a high recovery of the acidswith a low flowrate of water without neglecting thesize of the column required.

From Figure 9 It can be seen that larger columnsachieve an extremely high recovery of acrylic acidfor low water flowrates. However, due to the highcosts of larger columns this option is not aseconomically promising. As a compromise it hasbeen decided to use a 7 stage absorber with a

water flowrate of 0.167kmol/s (600kmol/hr). Thissolution achieves an acrylic acid recovery greaterthan 99.9% while using a low water flowrate and arelatively inexpensive column.

99.0%

99.2%

99.4%

99.6%

99.8%

100.0%

0.1 0.2 0.3 0.4 0.5

RecoveryofAcrylicAcid

7 Stages

9 Stages

5 Stages

3 Stages

Figure 9: Absorber acrylic acid recoveryversus flowrate of water

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5.2.2) Liquid-Liquid ExtractorIn order to minimise the construction costs of separation units further downstream of the flash vessel it hasbeen decided to reduce the pressure from 2.8bar to 1atm using a valve. Furthermore, when the pressure isdropped the stream will cool, which increases the effectiveness of the extraction process as seen in Table 3.

98.680

98.685

98.690

98.695

98.700

98.705

2000 2200 2400 2600 2800 3

Liq-LiqextractorEPx106($

/yr)

10 Stages

9 Stages

11 Stages

0.99994

0.99995

0.99996

0.99997

0.99998

0.99999

1.00000

2000 2200 2400 2600 2800 30

AcylicAcidRecovery

11 Stage

10 Stage

9 Stages

Figure 10: Economic Potential of the liquid-liquidextractor for different DIPE flowrates and columns

Figure 11: Liquid-liquid extractor acrylic acid recoveryfor different DIPE flowrates and columns

The DIPE will be recycled, reducing the total flow to the plant, as it is very expensive. However, as thedistillation columns further downstream had yet to be optimised, an estimate for the fresh DIPE feed had tobe used for an economic analysis to be preformed. The flowrate of fresh feed DIPE was estimated to beequal to the flowrate of DIPE lost in the stream leaving the top of the extractor, as the majority was lost here.The economic potential of this unit was estimated by subtracting the annualised installed costs of theextractor and the annual cost of estimated fresh DIPE feed from the sale value of the acids leaving this unitin the product stream (stream 17, Figure 8). Figure 10 illustrates that higher flowrates of DIPE and a smallercolumn would give a slightly higher economic potential. It can also be seen that the extractors economicpotential is not very sensitive to the flowrate of DIPE within the range of data shown. Considering the factthat the DIPE will have to be removed from the acid products in the next stage of the liquid separationsystem a lower flowrate will be used; drastically reducing the size and cost of the following distillationcolumn. However, it is important not to neglect the fact that a high recovery of the acrylic acid is required, toensure the product specification is met. In order to achieve a high recovery with a lower flowrate of DIPE aslightly larger column is required, which can be seen in Figure 11. Although this slightly increases the cost ofthe extractor it is likely to reduce the distillation costs by a much bigger margin, as this extraction process isvastly more effective than the distillation system, as shown in Table 3. Accounting for all these factors it hasbeen decided to use an 11 stage liquid-liquid extractor with an inlet DIPE flowrate of 2200kmol/hr(0.611kmol/s), achieving a very high recovery of acrylic acid.

5.2.3) Distillation Columns

4.44

4.46

4.48

4.50

4.52

4.54

4.56

4.58

1.1 1.15 1.2 1.25 1.3

Reflux ratio divided by the minimum reflux ratio

CostofDISTIL1

x106

($/yr)

0.062

0.064

0.066

0.068

0.070

0.072

0.074

0.076

CostofDISTIL2x106($/yr)DISTIL1

DISTIL2

Figure 12: Cost of both distillation columns against reflux ratio/minimum reflux ratio

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The DIPE and acid mixture leaving the liquid-liquid extractor contained a significant fraction of lightcomponents, which will readily vaporise in DISTIL1 and thus exit in the distillate. As refrigeration would berequired to totally condense the distillate a partial condenser will be used with a vapour fraction that isachievable with cooling water. Since the majority of the light components are no longer present in the feed toDISTIL2 a total condenser can be utilised, thus reducing the number of stages required. This is alsobeneficial as a liquid distillate product is required. Initially the recoveries in both columns were set at a highlevel of 0.99 for the light key (LK) and 0.01 for the heavy key (HK). These recoveries gave a high purity of

acrylic acid; however, the production rate was slightly below the requirement. In order to increase thisflowrate the light key recovery in DISTIL2 was reduced to 0.93. As these recoveries met acrylic acidsproduct specification the reflux ratio (RR) of each column was optimised at these conditions. This was donepurely on an economic basis where the annual installed column cost and utility costs of the condenser andreboiler were minimised. Figure 12 shows that the optimum reflux ratio for DISTIL1 is in the region of 1.15times the minimum reflux ratio (RRmin). It can also be seen that the cost of DISTIL2 decreases with refluxratio and no minimum is achieved. As it is undesirable to operate distillation columns with a reflux ratio ofless 1.1 times the minimum it has been decided to set a reflux ratio of 1.15 times the minimum for bothcolumns. These reflux ratios result in DISTIL1 requiring 42 stages and DISTIL2 requiring 17 stages. It wasnoticed that the diameters calculated using equation 16 (Appendix H) were mechanically infeasible. Thesewere therefore sized up to 3.66m and 2.13m for DISTIL1 and DISTIL2 respectively.

5.3) Process Adjustments

After completing the optimisation section the total process was reviewed. Although the acrylic acid had avery high mass purity of 99.0% and a total flowrate meeting the product specification, the mass purity of theacetic acid stream was only 52.3%. This purity of acetic acid was unacceptable for sale and must thereforebe improved. It was found that the main impurity in the acetic acid stream was DIPE. By reducing theflowrate of DIPE into the liquid-liquid extractor to 1250kmol/hr and increasing the light key recovery inDISTIL1 to 0.9997 a more acceptable acetic acid mass purity of 87.7% was achieved. The downside to doingthis is that the size of DISTIL1 increased to 48 stages. However, this additional cost is justified by the sale ofacetic acid. In order to meet the acrylic acid production rate the recovery of the light key in DISTIL2 had to bedecreased to 0.844 and the heavy key was decreased to 0.001. This increased the size of DISTIL2 to 23stages due to the higher recovery of the bottom component. This resulted in an acrylic acid production rate of50046tonnes/yr and mass purity of 99.0%.

Table 4: Final separation system set-up

Vessel Pressure(bar) Solvent Flowrate(kmol/hr) Number ofStages RR(RRmin-1)LKRecovery HKRecovery Partial CondenserVapour Fraction

FLASH 2.80 NA NA NA NA NA NAABSORB1 2.80 600 7 NA NA NA NALLEXT 1.01 1250 11 NA NA NA NADISTIL1 1.01 NA 48 1.15 0.9997 0.01 0.005DISTIL2 1.01 NA 23 1.15 0.844 0.001 0.00 (total)

Table 5: Final key stream results

Stream Key Component Key Component mass purity (%) Total Stream Flowrate (tonne/yr)

5 Acrlic acid 99.0 50,0006 Acetic acid 87.7 2,9701 Propylene 100 33,800

2 Air 100 237,0003 Steam 100 103,00021 Demineralised water 100 88,100

Top - up DIPE 100 352

6) Level 4 Economic Potential (EP4)

systemseparationofstcooperatingsystemseparationofstcocapitalEPEP = 34[10]

Using equation 10 the economic potential at level 4 was found to be 77.2106$/yr. The extra costs involvedin addition to level 3 have been explained below and example calculations can be seen in Appendix H and L.

6.1) Capital cost of the separation system6.1.1) Base vessel installed costThe installed cost of all the units enclosures in the separation system were calculated using Guthries costcorrelation for pressure vessels, reactors and columns (equation 7). The flash vessels volume was

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calculated by using the exiting vapour velocity and assuming a residence time of 10 seconds to allow forliquid and vapour disengagement (Appendix I). The height and diameter of this vessel was deduced from thevolume by assuming a height to diameter ratio of 6, which applies for such a vessels. The height of theabsorber, distillation columns and liquid-liquid extractor were calculated from the number of stages byassuming a 0.610m (2ft) standard tray spacing. In addition, extra height was added to the columns toaccount for tray inefficiencies and extra space required at the top and bottom for effective separation. Thediameter of the absorber and distillation columns were calculated from the volumetric vapour flowrate within

these vessels using a correlation, which can be seen in Appendix H, equation 16. As this could not beapplied for the liquid-liquid extractor a height to diameter ratio of 6 was assumed to calculate its diameter.

6.1.2) Internal tray installed costThe installed costs of the absorbers and distillation columns trays were calculated using equation 11.

cHFDSM

CostInstalledTray55.1

7.4280

&$,

= [11] where:

mtsc FFFF ++=

6.1.3) Reboiler and condenser installed costs for the distillation columnsThe installed costs of both of these units were calculated using equation 12, where A is the heat transferarea (ft2). Simple and commonly used kettle reboilers are to be installed on both distillation columns. For

such units UTcan be assumed as 11250 Btu.(hr.ft2

)-1

(Douglas, 1988), which allowed the heat transfer areato be calculated. Shell and tube heat exchangers are to be used as condensers for both distillation columns,as these apply for the operating pressure and temperature ranges (Hewitt, 2007). These heat exchangerswere sized using known log-mean temperature differences and heat transfer coefficients displayed inAppendix J (Hewitt, 2007). This method of costing was also applied to all coolers in the process.

( )cFA

SMCostInstalledExchangerHeat +

= 29.23.101

280

&$,

65.0[12] where:

mpdc FFFF +=

6.1.4) Furnace installed costThe furnaces installed costs were calculated using equation 13, where Q is the required duty in 106 Btu.hr-1.

( ) ( )cFQ

SMCostInstalledFurnace +

= 27.11052.5

280

&$,

85.03[13] where:

pmdcFFFF ++=

6.2) Operating cost of the separation system

The DIPE and demineralised water annular costs were calculated using the flowrates displayed in Table 5and their utility costs (Appendix B). The cooling water, steam and fuel oil utility costs were calculated foreach heat exchanger, furnace and reboiler by performing an energy balance on each of these units.

7) Heat Exchanger Network Synthesis

7.1) Arrangement of Reactant Preheating

The reactant preheating and mixing systems structure can be seen in Figure 8. First the compressed air andsteam streams were mixed together. This was to ensure that when propylene is added the resulting mixtureis outside of its explosive limits. The air and steam stream was then heated to avoid any steamcondensation. Following this, propylene was safely added and the entire stream was heated to 310C. Thepropylene stream was not heated before being mixed as an explosion could occur if any propylene leaks inthe furnace. Before this stream is fed to the reactor it was mixed with the gas recycle, which was also heatedto 310C.

7.2) Identification of Streams that can be Integrated

All process streams that require a heater or cooler have been considered for heat integration in order to

minimise the process utility costs. These can be seen in, where the supply temperatures were eitherprovided in Table 1 or calculated in ASPEN, see Appendix N. The heating or cooling requirements (duties) ofeach of these streams were also determined using APSEN and by dividing by the corresponding streamstemperature change, heat capacity flowrates were calculated. The heat released by the reactor was also

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considered in the heat integration (stream R) as this is also an available energy source. The reactor wasmodelled as a hot stream with a target temperature of 1C lower than its operating temperature, such that itsheat capacity flowrate is equal to its duty, which was calculated in Maple at level 3.

Table 6: Heat Integration streams information

StreamSupply

Temperature (C)Target

Temperature (C)Heating or

Cooling Duty (W)Heat Capacity Flowrate

(Fcp) (W/C)

24 cold 231 310 1230000 -1550022 cold 269 310 749000 -18400

10C cold 64.4 310 4340000 -17700

R hot 310 309 14900000 14900000

11 hot 310 45.0 19400000 73200

6 hot 100 40.0 11900 198

5 hot 141 40.0 197000 1960

7.3) Network Synthesis

The maximum energy recovery design was the first to be considered, where counter-current shell and tubeheat exchangers with a minimum approach temperature of 10C were used. The available utilities werecooling water at 30C (with a maximum temperature rise of 15C), saturated steam at various pressures and

fuel oil (see appendix B, Table 11). The first stage in this design was to determine the temperature intervals.The temperature interval grid can be seen in appendix K and the results have been summarised in Table 7.

Table 7: Energy surplus or deficit in each temperature interval

Temperature Interval (C) Fcphot - Fcpcold (W/C) Energy Surplus or Deficit in Interval (W)

310 300 -51500 -515000

300 299 14900000 14900000

299 269 21700 648000

269 231 40100 1540000

231 131 55600 5550000

131 90.0 57500 2340000

90.0 64.4 57700 1480000

64.4 35.0 75400 222000035.0 30.0 2160 10800

Using Table 7 a cascade view of the temperature intervals, which can be seen in Appendix K, wasconstructed. From analysing this energy cascade diagram it was found that the minimum hot utility duty toproduce a feasible cascade was 515kW, which resulted in a minimum cold utility duty of 28700kW. From thisa cold stream pinch temperature (temperature at which no heat is transferred between intervals) of 300C(310C hot stream) was also determined. For the maximum energy recovery design only hot utilities arerequired above the pinch temperature (hot side) and only cold utilities are required below the pinch (coldside). Thus, the heat integration network designs of the hot and cold sides were carried out separately.

7.3.1) Cold Side DesignImmediately below the pinch temperature hot and cold streams can only be integrated if Fcp hot Fcpcold, so

that the minimum approach temperature is maintained. At the pinch either stream 11 or the reactor can beused to heat streams 24, 22 and 10C. However, in order to do this the hot stream will need to be split. Byweighting the split fraction by streams 24, 22 and 10Cs duties the outlet temperature of the stream used toheat these will be equal and cause less stresses within materials when the split streams are mixed backtogether. This results in molar hot stream split fractions of 0.185, 0.097 and 0.718 to heat streams 24, 22 and10C respectively. When the heat capacity flowrates of each split was considered it was found that Fcphot Fcpcold for some split fractions of stream 11, which is not feasible, but this was not the case for the reactor.Thus, the reactor will be used to fully heat streams 24, 22 and 10C to the cold stream pinch temperature(300C). A cooler of duty 9080kW will be required to remove the remainder of the reactors heat and streams11, 6 and 5 will require coolers of duty 19400kW, 11.9kW and 197kW respectively. This results in a total coldutility duty of 28700kW. As this is equal to the minimum cooling duty previously determined, the number ofunits (7) is equal to the minimum (see Appendix K) and only cold utilities are required this meets all threecriteria of the maximum energy recovery system.

7.3.2) Hot Side DesignAs the hot side only contains cold streams requiring heating from the cold stream pinch temperature (300C)to 310C no integration can be performed. Thus, streams 24, 22 and 10C require heaters of duty 155kW,

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184kW and 177kW respectively. This results in a total hot utility requirement of 515kW. As this is equal to thepreviously discovered minimum heating duty, the number of units (3) is equal to the minimum (see AppendixK) and only hot utilities are required this meets all three criteria of the maximum energy recovery system.

7.3.3) Reducing the number of unitsIn order to reduce the total number of units used in this system loops between heat exchangers and utilitieshave been studied. Although loops do exist between the reactor heat exchangers and the hot utility, no

heaters could be removed as the reactor cannot heat the cold streams above 300C due to the minimumapproach temperature. The elimination of a heat exchanger is not economically beneficial as all heatexchangers have significantly high duties, which merit the units installed cost.

7.4) Integration of Distillation Columns and Process Streams

7.5) Final Heat Exchanger Network

Figure 14: Heat exchanger network

8) Level 5 Economic Potential (EP5)

tegrationinheattoduesavingsstcoutilityEPEP +=45

[14]

Using equation 14 the level 5 economic potential was found to be 80.1106$/yr. The utility costs of both thereboilers and condensers for the two distillation columns were completely eliminated by heat integration. Inaddition, the heat integration network also reduced the duty required to be supplied by the hot utility forheating streams 24, 22 and 10C. The costs of the additional heat exchangers required for this network havebeen calculated using the method described previously in section 6.1.3 and a cost break down is shown inAppendix M.

9) Process Sensitivity AnalysisA sensitivity analysis was preformed on the ASPEN simulation to analyse the effects of changing theflowrates of the inputs to the system. The input flowrates of Propylene (1), DIPE (18) and water (21) were

increased and decreased by 10% and the purity and flowrates of the acid products were calculated. Fromthis analysis, it was observed that the purity of the acrylic acid stream is only drastically changed by reducingthe flowrate of propylene by 10%. Thus, strict control over this flowrate is important. The full table can beseen in Appendix O.

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0

50

100

150

200

250

300

350

0 5 10 15 20 25 30

Enthalpy x106 (W)

Temperature(C)

GCCDistil 1Distil 2

The grand composite curve (GCC), which isgenerated from the cascade, is a graph oftemperature against enthalpy and indicates howmuch surplus energy there is over each temperatureinterval. The grand composite curve for this processis displayed in Table 7. From this it can be seen thatthe temperature enthalpy representation of bothdistillation columns fit within the process streamgrand composite curve. This indicates that surplusenergy can be used by the distillation columnsreboilers at high temperatures and returned by thecondensers at lower temperatures. Hence, theheating and cooling of both distillation columns iseliminated.

Figure 13: Grand composite curve and distillation columns temperature enthalpy diagrams

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10) Environmental considerations

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References

Douglas, J., M., 1988, Conceptual Design of Chemical Processes (McGraw-Hill, Singapore)

Felder, R., M. and Rousseau, R., W., 1986, Elementary principles of chemical processes (Wiley,New York, Chichester)

Happel, J. and Jordan, D.G., 1975, Chemical Process Economics, 2nd Edition (Dekker, New York),

pp 454 Hewitt, G., F., 2007, Process Heat Transfer Imperial College Option Course Lecture Notes

King, C.J., 1980, Separation Processes, 2nd Edition (McGraw-Hill, USA), pp 604 - 605

Metcalfe, I.S., 1997, Chemical Reaction Engineering A First Course (Oxford University Press, GreatBritain), pp 11

Perry, R.H., Green, D.W., 1997, Perrys Chemical Engineers Handbook, 7th Edition (McGraw-Hill,USA) pp 2-28

{1} www.cheric.org/kdb/ (23-01-07)

{2} www.chrismanual.com/A/ACR.pdf (23-01-07)

{3} en.wikipedia.org/wiki/Acetic_acid_(data_page) (23-01-07)

Nomenclaturebhp Break horse power hpCpi Heat capacity of component i j.mol-1.k-1

D Diameter mEPi Economic potential at level i $.yr-1

rE Activation energy of reaction r kJ/kmol

Fi Molar flowrate of stream i mol.s-1

flp Feed mass fraction of propylene NAfls Feed mass fraction of steam NAH Entalpy WH Height mHK Heavy key NA

rok , Rate constant of reaction r kmol/(m3.(kPa)2.yr)

L Length mLK Light key NAM&S Marshall and swift equipment cost index NAP Pressure kPa, bar, atmPprop Partial pressure of propylene kPaPo2 Partial pressure of Oxygen kPaQ Heat WR Universal gas constant J.mol-1.K-1

RRmin Minimum reflux ratio NARR Reflux ratio NA

rr Rate of reaction r kmol.m-3.yr-1

T Temperature C, KTR Temperature of reaction Kyij Mole fraction of componetjin stream i NA

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Appendix A Selectivity Data

y =-0.2036x2+0.4441x +0.67

0.870

0.875

0.880

0.885

0.890

0.895

0.900

0.905

0.910

0.915

0.60 0.65 0.70 0.75 0.80 0.85

SelectivityofReaction1(s1

Figure 15: Relationship of the selectivity of reaction 1 with respect to conversion

y =0.1516x2- 0.3335x +0.242

0.00

0.02

0.04

0.06

0.08

0.10

0.12

0.60 0.65 0.70 0.75 0.80 0.85

SelectivityofReaction2(s2

Figure 16: Relationship of the selectivity of reaction 2 with respect to conversion

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Appendix B Data Tables

Table 8: Kinetic data for equation 1

Reaction Er(kJ/kmol) k0,r(kmol/(m3.(kPa)2.yr))

R1 63000 1.30 x 109

R2 84000 7.20 x 109

R3 100000 1.48 x 109

Table 9: Commodity Prices

Component Price ($/kg)

Propylene 0.38Process Steam 139.36 x 10-4

Acrylic Acid 1.92Acetic Acid 0.97

Table 10: values for different gas structures (Happel et al., 1975)Gases

Monatomic 0.4

Diatomic 0.29More complex gases (CO2, CH4) 0.23Other R/Cp

Table 11: Utility Costs

Utility Cost Units

Cooling Water 8.1210-6 $.kg -1

Saturated steam at 41.4bar and 253C 9.9710-3 $.kg-1





Compressor power cost 0.04 $.(hp.hr)-1

Fuel oil for combustion in furnaces 3.7910-9 $.J-1

Demineralised water 16.132 $.(kmol.hr -1.yr)-1

Di-isopropyl Ether (DIPE) 0.89 $.kg-1

Table 12: Properties of the components in the liquid separation system (Perry et al., 1997)

Property ComponentsH2O C3H4O2 C2H4O2 DIPE

Boiling point (b.p.) at 1 atm. (C) 100 141 118 69Vapour Pressure (Pvap) at 50C (mmHg) 91.7 18.1 56.6 401Solubility in DIPE (KD)(42.22C, 2.5 bar) 0.02039 2.048 2.765 3459(62.08C, 2.5 bar) 0.02869 1.102 1.875 1226

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Appendix C Material balances required at level 2

Figure 17: Level 2 process diagram with the mix and feed stream to the reactor displayed

Subscript p has been used to represent propylene.

Mole balance around the reactor and separation system for:

C3H4O2: pyxFsF 441500 +=

C2H4O2: pyxFsF 442600 +=

Defining the feed mass fraction design variables:

H2O(g):321

3

182942

18

FFF

Fms ++

=

C3H6:321

1

182942

42

FFF

Fmp ++

=

Define a term for 99.5% percentage loss of propylene in the purge:

1001

441 63

= F

yxFF

L

HC

p

These 5 equations have been solved in Maple to find the molar flowrates of acrylic acid, acetic acid,propylene feed and input steam in terms of the design variables. Converting these flowrates into massflowrates allows the economic potential to be found in terms of the design variables. The Maple code for thiscan be seen on the attached CD.

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Appendix D Material balances required at level 3Perfect separation in the separation system has been assumed at this level.

Mole balance around the whole process for:

C3H4O2: 63441500 HCyxFSF +=

C2H4O2: 63442600 HCyxFSF += C3H6: 6363 449910 HCHC yxFyFF =

H2O: 636363 442144244173 )1(30 HCHCHC yxFSSyxFSyxFSFF +++=

N2: 22 99220 NN yFyF =

O2:

63

636322

4421

44244192

)1)(2/9(

)2/5()2/3(920

HC

HCHCNO

yxFSS

yxFSyxFSyFyF

= CO2: 63632 442144299 )1(300 HCHCCO yxFSSyxFSyF ++=

Mole balance around the mixer for:

C3H6: 6363 10104410 HCHC yFyFF += N2: 222 101044220 NNN yFyFyF +=

O2: 222 101044220 OOO yFyFyF +=

H2O: OHyFF 2430 =

CO2: 22 10104400 COCO yFyF +=

Physical Constraints:

Stream 9: 163222 9999=+++ HCCOON yyyy

Stream 4: 163222 4444=+++

HCCOON yyyy

Equations of the splitter:

Total mole balance: 91080 FFF =

N2mole fraction: 222 1098 NNN yyy ==

O2mole fraction: 222 1098 OOO yyy ==

CO2mole fraction: 222 1098 COCOCO yyy ==

C3H6: mole fraction: 222 1098 OOO yyy ==

Taking the mid-point of the allowable mass fraction ranges for the feed to the reactor:

H2O(g): 275.0182942

18

321

3 =++= FFFF

ms

C3H6: 075.0182942

42

321

1

63=

++=

FFF

Fm HC

Define a term for the percentage loss of propylene in the purge:

1001

441 63

63

=

F

yxFFL

HC

HC

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Appendix E Hesss Law Calculation

The enthalpy diagram, shown in Figure 18, was constructed assuming the gas stream behaves ideally so theenthalpy only varies with temperature.

Figure 18: Enthalpy Diagram for the reactions that occur within the reactor

Table 13: Conditions for each point displayed in the above enthalpy diagramLetter Components Temperature (C) Phase

a N2, O2, CO2, H2O, C3H6 (or p) 310 g

b H2O 100 g

c H2O 100 ld N2, O2, CO2, H2O, C3H6 25 l&ge N2, O2, CO2, H2O, C3H6, C3H6O2,

C2H6O225 l&g

f H2O 100 lg C2H6O2 118.1 lh C3H6O2 141 li H2O 100 gj C2H6O2 118.1 gk C3H6O2 141 gl N2, O2, CO2, H2O, C3H6 Tout g

Enthalpy balance equations for each change shown in Figure 18, where the variable heat capacities can beseen in Table 14 and the heats of vaporisation can be seen in Table 15.

( )

( )vapOHC

OHC

T

pp

T

NN

T

OO

T

COco

RRR

OHOH

vap

OHOH

OHOH

ppNNOOCOco

HFH

dTCpFH

dTCpydTCpydTCpydTCpyFH

HHHH

dTCpyFH

HyFH

dTCpyFH

dTCpydTCpydTCpydTCpyFH

243

243

222222

22

22

22

222222

58

141

15

57

25

8

25

8

25

8

25

886

3322115

25

100

444

443

100

310

442

25

310

4

25

310

4

25

310

4

25

310

441

=

=

+++=

++=

=

=

=

+++=

dTCpFHT

OHC=141

59 243

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dTCpFH OHC=1.118

15

610242

( )

dTCpFH

HFH

dTCpFH

dTCpFH

HFH

T

OH

vap

OH

OH

T

OHC

vap

OHC

=

=

=

=

=

100

715

714

100

25

713

1.118

612

611

2

2

2

242

242

From Figure 18 it can be deduced that the change in enthalpy across the reactor (moving from position a to l)can be found by summing all of the enthalpies shown above (H1toH15).

Table 14: Heat capacities of each component required in the above enthalpy equations

Component Phase Heat Capacity (kJ.(kmol.K)-1) Reference

H2O Gas 33.46+0.688*10-2T+0.7604*10^-5T2-3.593*10-9T3

(Felder et al.,1986)

N2 Gas 29+0.219910-2T+0.572310-5T2-2.87110-9T3

O2 Gas 29.1+1.15810-2T-0.607610-5T2+1.31110-9T3

C3H6 Gas 59.58+17.7110-2T-10.1710-5T2+24.610-9T3

CO2 Gas 36.11+4.23310-2T-2.88710-5T2+7.46410-9T3

C2H4O2 Gas 4.94910-8(T+273)3-1.75310-4(T+273)2+2.54910-1(T+273) + 4.84 {1}C3H4O2 Gas 6.97510-8(T+273)3-2.35210-4(T+273)2 + 0.3191(T+273) + 1.742 {1}H2O Liquid 75.4 (Felder et al.,

1986)C3H4O2 Liquid 138.7 {2}C2H4O2 Liquid 123.1 {3}

Table 15: Heats of vaporisation required for the above enthalpy equations

Component Heat of vaporisation at 298K (kJ/kmol) Reference

H2O 40650 (Felder et al., 1986)C3H4O2 46024 {1}C2H4O2 23681 {1}

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Appendix F Pressure Sensitivity Analysis

Figure 19: Gas recycle compressor cost against conversion fordifferent inlet pressures at a 1% loss of propylene

Figure 1 indicates that the gas recycle compressors cost varies significantly with the selected inlet pressureand hence the assumed pressure drop across the reactor and separation system. It can also be seen thatthe cost is more sensitive at lower conversions. This sensitivity must be kept in mind when analysing theeconomic potentials at level 3.

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Appendix G Selection of Column Type

Table 16: Selection Guide* for distillation-column internals (King, 1980)

ConditionApplication

for this

process

Tray Packed ColumnSieve

or

valve

Bubble-

cap

Counterflow Random StackedDisk and

doughnutLow Pressure (50% of critical)

No 3 2 2 2 0 0

High turndown ratio No 2 3 0 1 2 1Low Liquid flowrates

No 1 3 0 1 2 0

Foaming Systems No 2 1 2 3 0 1Internal towercooling

No 2 3 1 1 0 0

Suspended solids No 2 1 3 1 0 1Dirty orpolymerizingsolution

No(neglected)

2 1 3 1 0 2

Multiple feeds orsidestreams

No 3 3 2 1 0 1

High liquid flowrates

No 2 1 3 3 0 2

Small-diametercolumns

No 1 1 1 3 2 1

Column diameter1 to 3m

Yes 3 2 2 2 2 1

Larger-diameter

columns

No 3 1 2 2 1 1

Corrosive fluids Yes 2 1 2 3 1 2Viscous fluids No 2 1 1 3 0 0Low pressure drop(efficiencyunimportant)

No 1 0 0 2 2 3

Expanded columncapacity

No 2 0 2 2 3 0

Low cost(performanceunimportant)

Yes 2 1 3 2 1 3

Reliability of design Yes 3 2 1 2 1 1

* 0 = do not use, 1 = evaluate carefully, 2 = usually applicable, 3 = best selection

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Appendix H Sample Calculation for Costing a Separation System VesselFor this example both the annualised installed cost and operating cost of the second distillation columnDISTIL2 has been shown.

Required Information Obtained from Aspen

Table 17: Required information to cost the second distillation column DISTIL2

Refluxratio/min

refluxratio

ActualRefluxRatio

Numberof

Stages

HeatDuty of

theReboiler

(W)

ReboilerTemp.(C)

CoolingDuty of theCondenser

(W)

CondenserTemp.(C)

Density ofgas

(kmol/m3)

Distillateflowrate(kmol/s)

Molecularweight gas

(g/mol)

1.1 6.80 19 480000 141 475000 100 7.19 0.00240 73.4

Installed Cost

First the distillation column was required to be sized. The calculations for the height also apply for theabsorber and liquid-liquid extractor but the diameter calculations only apply for the distillation columns andabsorber.

Calculating the column heightThe column height was initially calculated by multiplying the tray spacing of 0.610m (2ft) by the number ofstages, which neglects the extra space required in the ends and plate efficiency. This is equal to 11.6m forthe example being considered.

The extra spacing for the ends was than accounted for by adding an additional 15% to the previouslycalculated height (Douglas,1988). The plate efficiency was assumed to be 90% and this was included in theheight by dividing through by 0.9.

Actual column height mm

8.149.0

15.16.11=

= , in general

9.0

610.015.1 =N

H , N = number of

stages

Calculating the Column DiameterThe diameter of the distillation column was calculated using equation 16, where V= molar vapour flowrate

(mol.hr-1), MG = molecular mass of the gas (lb.mol-1) andm = gas molar density (mol/ft3). The molar vapourflowrate was determined by performing material balances on the top of the column using the reflux ratio. Byconverting units it was found thatDis equal to 0.715ft for this example.

41

0164.0

=

m

GMVD

[16]

Calculating the base vessel installed costThis method for calculating the installed cost of the vessel enclosure applies for all units in the separationsystem. Now the distillation columns height and diameter are known equation 7, which has been shownbelow, can be used to calculate the installed cost of the basic vessel enclosure. As the distillation column ismade out of stainless steel and operates at a pressure of 1atmFm = 3.67 andFp = 1 (Douglas,1988). For thisexample the installed cost is equal to $36700.

( )( ) ( )cFHD

SMCostInstalledVeseel += 18.29.101

280

&($) 802.0066.1 [7]

Calculating the tray installed cost

The following calculations apply for both distillation columns and the absorber. Using equation 11, shown onthe next page, the installed tray cost can be calculated. For the considered example Fs = 1, as the trayspacing is equal to 24in. (0.610m),Ft = 0.0, as sieve trays are being used and Fm = 1.7, as the trays will be

- 29 -

Where: pmc FFF = ,

D= diameter (ft)H= height length (ft)

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made out of stainless steel (Douglas,1988). Thus, Fc= 2.7 and the tray installed cost is for this example is$1430.

cHFDSM

CostInstalledTray55.17.4

280

&$,

= [11] where:

mtsc FFFF ++=

Calculating the reboiler installed costFirstly the reboiler was sized using equation 17, where UT has been estimated as 11250 Btu.(hr.ft2)-1

(Douglas, 1988). By converting the heat duty displayed in Table 17 into the appropriate units, the heatexchange area in the reboiler was found to be 146ft 2 for this example. Using equation 12, shown below, theinstalled cost of the reboiler was found to be $61440, asFd = 1.35 for a kettle reboiler,Fp = 0 as the pressure

is at 1 atm. and Fm = 2.81 as stainless steel and carbon steel would be used for each side of the heatexchanger (Douglas).

TU

QA

= [17]

( )cFASM

CostInstalledExchangerHeat +

= 29.23.101

280

&$,

65.0[12] where:

) mpdc FFFF +=

Calculating the condenser installed costSimilarly to the reboiler, the installed cost of the condenser can also be calculated using equation 12, where

FpandFm remain the same butFdwould change to 1 as floating head shell and tube heat exchangers are tobe employed (Douglas). The area of these heat exchangers were found using equation 17 but with a log-mean temperature difference, as the cooling water would change in temperature. The heat transfer co-efficient selection has been discussed in Appendix J. For this example, a heat transfer co-efficient of764W.m-2.K-1 has been selected and the log-mean temperature difference is equal to 62.3C, resulting in aheat transfer area of 9.98m2 (107ft2). Using equation 12 it can be deduced that the condenser installed costis equal to $42086. This cost calculation was also applied to all other coolers in the process and heatexchangers.

All of these installed costs for the distillation column were then summed and divided by a capital chargefactor of three to convert these into an annualised cost, which results in a total annualised installed cost of47200$.yr-1. This can also be applied to all units in the separation system.

Operating Cost

The distillation columns operating costs are made up of both the heating and cooling utilities required for thereboiler and condenser. Other units that involve the use of solvents such as the absorber and liquid-liquidextractor also need to include the cost of these.

Reboiler operating costAs the reboiler in DISTIL2 operates at a temperature of 140.75C saturated steam at 10.3 bar was thechosen utility as this has a temperature of 186C, which is required to ensure the minimum approachtemperature is maintained. By performing an energy balance for this utility and using the cost informationdisplayed in Table 11 the flowrate of steam required to supply 480000W of energy was found to be0.241kg.s-1. This corresponds to an annualised cost of 52900$/yr, which was calculated using theassumption that the plant operates 8150 hours in a year.

Condenser operating costThe flowrate of the cooling water can be calculated through an energy balance using the maximum coolingwater temperature change (Tmax=15C) (if this is feasible for the condensers operating temperature) and itsheat capacity (cp,w= 4.2 kJ.kg-1.C-1). For this examples condenser duty of 475000W, 7.54kg.s-1 of coolingwater would be required. Using Table 11 this can be converted to an annualised cost, which was found to be1800$.yr-1.

Thus, the total operating cost for DISTIL2 is 54700$.yr-1

for these example conditions. By summing theannualised installed cost and operating cost a units total annualised cost can be found. For this example thetotal annualised cost is equal to 97400$.yr-1. Although this example was for a distillation column it covers

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nearly all of the additional utility costing required at level 4 and the same methods have been applied to otherunits.

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Appendix I Sizing the Flash Vessel

From ASPEN it was found that the:

Volumetric flow rate of the gas within the flash vessel,13

90.7= smVg

Vapour fraction within the flash vessel, 778.0=gx

To size the flash vessel a residence time, , of 10 seconds was assumed. This should be sufficient for theliquid and vapour phases to separate as the mixture fed to the flash vessel is two phase from the quenching.

The volume of the vapour within the flash vessel,3

79mVVV ggg ==

The total volume of the flash vessel,3102mV

x

VV

g

g ==

To ensure the flash vessel would operated as it is required, this volume was sized up to 150m3

Assume that the height to diameter ratio is 6, 6=DH

Solving for the diameter results in mDV

D 17.36

43 ==

Thus, mH 0.19=

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Appendix J Heat Transfer Co-efficient SelectionThe heat transfer co-efficients were determined for each heat exchanger using ESDU data tables (Hewitt,2007). The following information was used to determine the heat transfer co-efficients:

type of heat exchanger (this was always shell and tube) duty divided by appropriate temperature difference (Q/T) cold side fluid hot side fluid

Table 18: Heat transfer co-efficients selected for each heat exchanger

Heat Exchanger Q/T (W.K-1)Closest Cold

Side FluidMatch

Closest HotSide Fluid

Match

Heat TransferCo-efficient(W.m-2.K-1)

DISTIL1 Condenser1750000

Treated CoolingWater

CondensingHyrdro-carbon

764

DISTIL2 Condenser8900


CondensingHyrdro-carbon

764

ACECOOL450


Low ViscosityOrganic Liquid

714

ACRCOOL 15500 Treated CoolingWater Low ViscosityOrganic Liquid 714

QUENCH223000


MediumPressure Gas

484

10C R (HE)5660

MediumPressure Gas

MediumPressure Gas

300

22 R (HE)25800

MediumPressure Gas

MediumPressure Gas

300

24 R (HE)32000

MediumPressure Gas

MediumPressure Gas

300

Reactor Cooler33300


MediumPressure Gas

484

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Appendix K Heat Integration

Temperature Interval Grid

Figure 20: Temperature interval grid (where the hot streams account for the minimum approach temperature)

The corresponding cascade is displayed on the next page.

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Energy Cascade

Figure 21: Cascade view of the temperature intervals, where energy flows are in W

From Figure 21 it can be seen that the minimum hot utility duty is 515,000W and the minimum cold utilityduty is 28,700,000W. This figure also displays that the cold stream pinch temperature is 300C, whichcorresponds to a hot stream pinch temperature of 310C.

Minimum Number of Units for Maximum Energy Recovery

The minimum number of units for maximum energy recovery can be calculated using equation 15. Wherex

Umin

is the minimum number of units for the x side design (hot or cold),x

sN is the number of streams on

the x side design andx

utN is the number of utilities available on the x side design.

1min +=x

ut

x

s

xNNU

[15]

For the code side design: 7117minmin

=+=coldcold

UU

For the hot side design: 3113 minmin =+=hothot UU

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Appendix L EP4 Price Break-Down

EP4 77,242,720.83

EP3 83,079,692.78

Annualisedinstalled costof column and

trays ($/yr)

Annualsolvent

cost ($/yr)

Total cost($/yr)

ABSORB1 111,024.77 9,679.20 120,703.97

Annualisedinstalled cost

of column($/yr)

Cost ofDIPE lostin stream20 ($/yr)

Cost of DIPElost in

stream 7($/yr)

Total cost($/yr)

LLEXT 58,492.82 92,460.23 220,501.95371,455.0

0

Annualisedinstalled costof column and

trays ($/yr)

CostReboiler

Utility($/yr)

Cost CoolingUtility ($yr)

Reboilerarea (ft2)

ReboilerInstalled

cost ($/yr)

Condenser area (ft2)

Condenserinstalled cost

($/yr)

Totalcost($/yr)

DISTIL1 615,006.842,060,643.

08 66,013.285669.6610

04220,959.7

2 24592.5 480,772.263,443,39

5.19

DISTIL2 181,962.53 66,341.16 2,265.86168.85214

67 22,510.25 125.469 15,558.45288,638.

25

Utility cost

($/yr)

Heatexchange

area (ft2)

Annualisedinstalled cost

of column($/yr)

Totalcosts

($/yr)

QUENCH 73,396.42 4957.8 169,767.87243,164.3

0

ACECOOL 45.01 6.794 2,337.76 2,382.76

ACRCOOL 746.18 232.986 23,2

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