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PROCESS DESIGN OFMALEIC ANHYDRIDE

PLANTBY

WORIL TURNER DUDLEYVIJAYA KRISHNA BODLA



TABLE OF CONTENTS

1. Introduction2. The five processes selected3. Product and Process Selected for Design4. Screening of Process Alternatives

5. Material Balance6. Energy Balance7. Equipment Sizing8. Equipment Costing

9. Heat Intergration10. Economic evaluation11. Environmental Analysis12. Conclusion



INTRODUCTION

The process design project involves designing a process plant for producing aparticular product.

A list of five products are selected, then from that list one product is selected for thedesign project.

Selection of the best process for the design from a list of alternatives is then done.

Material and energy balances are done, equipment sizing and costing, and then aneconomic evaluation of the process.

Different tools of process optimization were considered for cost savings.

Heat integration is done for the process to calculate the additional heating or coolingrequired.

An environmental analysis was also done to determine the environmental impact of effluent discharge streams.



Product Names Raw Materials Process References

Maleic Anhydride n-butane and air Huntsmann fixed bed maleic

anhydride process, Kirk OthmerEncyclopaedia of ChemicalTechnology by Timothy R.

Felthouse, Joseph C. Burnett, BenHorrell, Micheal J. Mummey and

Yeong-Jen Kuo

Citric Acid Sucrose or dextrose Fermentation of sugars to producecitric acid, Shreves Chemical

Process Industries 5th Edition byGeorge T. Austin, page 598

Acetaldehyde Oxygen, water and ethylene Oxidation of ethylene to produceacetaldehyde, Shreves Chemical

Process Industries

AmmoniaNitrogen and Hydrogen Ammonia process, Shreves

Chemical Process Industries 5thEdition by George T. Austin page

306

Cinnamic Aldehyde Water and aldol Cinnamic aldehyde production byaldol condensation, Shreves

Chemical Process Industries 5thEdition by George T. Austin page

494



PRODUCT SELECTEDThe unique nature of maleic anhydride's chemical structure results in a highly

reactive and versatile raw material.

Its unsaturated double bond and acid anhydride group lend themselves to avariety of chemical reactions.

Maleic anhydride's largest use today is in the production of unsaturatedpolyester resins.

Another significant use is in the manufacture of alkyd resins, which are in turnused in paints and coatings.

Other applications where maleic anhydride is used include the production ofagricultural chemicals, maleic acid, copolymers, fumaric acid, lubricantadditives, surfactants and plasticizers.

Future applications are anticipated to be numerous given the versatility andusefulness of the product.



REACTIONS INVOLVED

C4H10 + 3.5 O2 C4H2O3 + 4 H2O∆H = -1236 kJ/mol (-295.4 kcal/mol)

C4H10 + 6.5 O2 4 CO2 + 5 H2O

∆H = -2656 kJ/mol (-634.8 kcal/mol)

C4H10 + 4.5 O2 4 CO + 5 H2O

∆H = -1521 kJ/mol (-363.5 kcal/mol)



SCREENING OF PROCESS ALTERNATIVES

There are two predominant raw materials for producing maleic anhydride, n-butane and benzene. Benzene however is a major environmental concern,because it is deemed as carcinogenic, so on environmental grounds, without evenlooking at raw material costs, benzene is rejected as the raw material for themaleic anhydride manufacture.

The process is a high temperature process so all the components leaving the

reactor are gases, so several separation options exist. The gases can be flashed,to recover water and maleic anhydride as liquids, while the other gases will remainin the vapor phase. We then consider separating water from maleic anhydride byexploiting the differences between their physical properties.

A solvent can be used for the product recovery, by contacting the product gaseswith a liquid solvent and then separating the maleic anhydride from the solvent. Anumber of alternatives exist for the solvent.

The conversion of butane is 85%, so recycling the unreacted butane is an option.

It is decided to use a process in which a solvent is used for absorbing the maleicanhydride produced.



Process Flow sheet

2. Reactor 3. Absorber 5. Mixer 2

7. Distillation

1. Mixer

µ01 µ1

µ2

µ31

µ32 µ51

µ81

4. Condenser

μ42

μ41

6. Distillation

8. Mixer 3

μ61

μ62 μ71

μ72 So

R1

9. Splitter

Purge (μ92)

Final Column

μp

μS

HE1

HE2

C2

R2

C1

R1

R3

C3

CM1

CM2



2. Reactor 3. Absorber 5. Mixer 2

7. Distillation

1. Mixer

µ 01 µ 1 µ 2

µ 31

µ 32 µ 51

µ 81

4. Condenser

μ42

μ41

6. Distillation

8. Mixer 3

μ61

μ62

μ71

μ72 So

R1

9. Splitter

Purge (μ92)

Final Column

μp

μS

MASS BALANCE FLOW CHART



MASS BALANCE

Assume an inlet flow of 100 Kmol/h of butane

Assume compressed air is fed in a ratio, where the amount of Oxygen is 1.5times the amount required

Using the yield of Maleic Anhydride, percentage conversion of butane andthe reaction stoichiometry of the reaction material balance relations are

written for the reactor For the side reactions, it is assumed that equal amounts of butane reacts toform Carbon Dioxide as for Carbon Monoxide

Split factors are then specified for all the separation equipment as well asfor the purge

The absorber is specified to be an isothermal absorber

It is assumed that the solvent entering the column doesnot contain anyMaleic Anhydride

For the absorber mass balance model, the Kremser equation is used todetermine the number of stages, using the split factor for the keycomponent recovery.



The remaining split factors for the other components are calculated from the

Kremser relationshipTo solve the mass balance model, the flow sheet is partitioned into twomodules and the recycle broken by tearing the inlet stream to the reactor.

Once the component flows to the reactor are calculated, from the massbalance model, we can sequentially calculate all the other flowrates

With the flowrates calculated, distillation column temperatures can becalculated.

Temperature for the vapor leaving top of the column is found from a dewpoint calculation. The temperature in the condenser and reboiler iscalculated from a bubble point calculation

It is assumed that the distillation column operates at one atmosphere of

pressure.



Component μ01 μ1 μ2 μ31 μ32 μ41 μ42 μ51 μ61 μ62 μ71 μ72

Maleic Anhydride 0.0 0.1 57.8 0.3 57.5 0.1 0.2 57.7 0.3 57.4 57.4 0.1

SuccinicAnhydride 0.0 5.7 5.7 98.9 2243.7 7.1 91.8 2335.5 0.0 2335.5 2.3 2333.2

Nitrogen 2238.0 11187.0 11187.0 11187.0 0.0 11187.0 0.0 0.0 0.0 0.0 0.0 0.0

Oxygen 525.0 965.0 550.5 550.5 0.0 550.5 0.0 0.0 0.0 0.0 0.0 0.0

Butane 100.0 113.4 17.0 16.8 0.2 16.7 0.1 0.3 0.3 0.0 0.0 0.0

Carbon Dioxide 0.0 61.6 77.1 77.1 0.0 77.0 0.0 0.1 0.1 0.0 0.0 0.0

Carbon Monoxide 0.0 61.7 77.1 77.1 0.0 77.1 0.0 0.0 0.0 0.0 0.0 0.0

Water 0.0 284.2 424.1 375.2 48.9 355.2 20.0 68.9 68.6 0.3 0.3 0.0

Total (Kmol/h) 2863.0 12678.7 12396.4 12382.9 2350.4 12270.8 112.1 2462.6 69.2 2393.3 60.1 2333.3

Pressure(Kpa) 200.0 200.0 200.0 150.0 109.0 109.0 109.0 101.0 101.0 101.0 101.0 101.0

Temperature (K) 300.0 350.0 700.0 400.0 395.0 395.0 395.0 395.0 369.0 514.0 475.0 536.6

Vapor Fraction 1.0 1.0 1.0 1.0 0.0 1.0 0.0 0.0 0.0 0.0 0.0 0.0



Components μp μfs S0 R1 μ92 Lo

Maleic Anhydride 57.3 0.1 0.0 0.1 0.0

Succinic Anhydride 0.0 2.3 1.3 5.7 1.4 2337.0

Nitrogen 0.0 0.0 0.0 8949.6 2237.4

Oxygen 0.0 0.0 0.0 440.4 110.1Butane 0.0 0.0 0.0 13.4 3.3

Carbon Dioxide 0.0 0.0 0.0 61.6 15.4

Carbon Monoxide 0.0 0.0 0.0 61.7 15.4

Water 0.3 0.0 0.0 284.2 71.0

Total (Kmol/h) 57.7 2.4 0.0 9816.7 2454.2

Pressure(Kpa) 101.0 101.0 101.0 101.0 101.0

Temperature (K) 474.6 532.9 394.0 394.0 394.0

Vapor Fraction 0.0 0.0 0.0 1.0 1.0 0.0



ENERGY BALANCE

The heat contents of all the streams are evaluated and heating and cooling duties for all the heat exchangers in the process determined.Kinetic and potential energies are neglected, and only enthalpy changes for thestreams are considered.It is assumed that there is no ΔH of mixing, or pressure effects on ΔH.

A standard reference of 298 K and one 1 atm or 101 Kpa of pressure is chosen.The enthalpy of each stream is now considered in turn, using the following enthalpycorrelations:To calculate the enthalpies of vapor mixtures the following correlation is used,ΔHv (T, y) = ΔHf + ΔHT = ∑k yk Hf, k (T1) + ∑k yk ∫ Cpo

, k (T) dTTo calculate the enthalpies of liquid mixtures the following correlation is used,ΔHL

k (T) = ΔHof, k + ∫Cpo

, k (T) dT - ΔHkvap

For the reactor the following expression is used,

QR = μ2ΔHv (T, y2) – μ1ΔHv (T, y1)QR is the heat of reaction, which is positive for an endothermic reaction and negativefor an exothermic reaction.ΔHv (T, y2) = ΔHout and ΔHv (T, y1) = ΔHin With the stream enthalpies, and stream temperatures, the heating and cooling dutiescan now be calculated



Stream μ0 μ1 μ2 μ31 μ32 μ41 μ42 R1 μ92

Flow (Kmol/h) 2863.0 12678.7 12396.4 12382.9 2350.4 12270.8 112.1 9816.7 2454.2

Pressure (Kpa) 200 200 200 150 150 109 109 200 101

Temperature

(K) 300 372 619.3 400 400 393 393 393 393

Enthalpy

(KJ/h) -1.24E07 -8.9E07 -4.78E08 -1.45E08 -1.28E09 -9.62E07 -5.56E07 -7.66E07 -1.96E07





SizingReactor:

Reacto

r

Volume

(m2)

Outside

tubediameter(m)

Inside

tubediameter(m)

Inner

crosssectionalarea (m2)

Tube

Length(m)

Inside

volumeof eachtube(m2)

No. of

tubes

Outside

surfacearea ofeachtube(m2)

MultiTubular

4.492676 0.035 0.02892 0.0006565 6.09 0.003998 1123.625 0.669291

Mixers:

Mixer Fl Τ

(residencetime) (1/hr)

Temp(K)

MolarDensity ofthe flow(kmol/m3)

Volume(m3)

Diameter (m)

Length(m)

Area(m2)

M1 2863 0,083333 372 188,1388 2.53624 0,931295 3,725181 0,680838

M2 2462.5 0,083333 399 22,13813 12.0502 1,565631 6,262522 1,924191

M3 2337 0,083333 399 15,38897 15.1862 1,691115 6,764462 2,244998

S litt



Splitter:

Fl Τ

(residencetime) (1/hr)

Temp (K)

Molar Densityof the flow (kmol/m3)

Volume(m3)

Diameter(m)

Length(m)

Area (m2)

Splitter 12270,8 0,033333 393 184,7362 4,42822 1,1214 4,48566 0,98719

Compressors:

Compressor

P2 P1 T1 γ µW

(kJ/hr) ηc

ηm(Shaftdriven)

Wb (shaftdriven) (watts)

CM1 (Forstream R1) 200 109 393 1,4 9816,7 2125869 0,8 0,9 8,201658291

CM2 (For AirCompression)

200 103,2 298 1,4 2763 4981751 0,8 0,9 1,921971726



Heat Exchangers:

HeatExchangers

Q U T1 T2 Delta Tln Area(A) (m2) Amountofcooling

water(Kmol/hr)

HE1 8,38E+07 919,8756 619,3 400 163,6839 556.5563 14800

HE2 1,19E+07 1430,922 536 400 130,1257 63.90984 2100

Condenser:

Condenser Q U T1 T2 Delta Tln Area(A)(m2)

Amount of coolingwater (Kmol/hr)

C 2,64E+06 919,8756 400 394 54,39447 52,76186 467



For the distillation columns, the number of trays and the reflux ratio weredetermined by the method of Westerberg, assuming ideality.

ICAS PDS was used to determine the number of trays for comparative

checks.

In cases where the method of Westerberg was giving a reflux ratio whichwhen used in ICAS was giving tray number in excess of 100, the method of Underwood was used to determine the minimum reflux ratio and theheuristic of the reflux ratio being 1.2 * the minimum reflux ratio used to getthe reflux ratio.

An overall column efficiency of 80% is assumed.

The column height is calculated using specified values for the tray spacing,extra feed space, disengagement space, skirt height and calculating theheight of the tray stack from the number of trays and the value of the tray

spacing.

The column diameter is calculated by using the Souder Brown equation todetermine the maximum allowable vapor velocity based on the columncrosssectional area.



For the absorber and flash drum, number of theoretical stages calculated bythe Kremser equation.

Column efficiency is however much lower than distillation columns,generally around 20%, which was the figure used.

The column diameter for the absorber is determined where total flows Vjand Lj are largest. This is at the bottom of the column.

The diameter is then determined as for distillation column.

The solvent recovery unit is a flash drum.

The vapor velocity is calculated and is used to determine the columndiameter as done for the absorber and the distillation column.



Column 1

αlk/hk

(avg) N1

N2

βlk

(ξlk) βhk

(1-ξhk) YN

NT

YR

R1

R2

R

25.33 3.54 3.54 0.995 0.995 0.8 3.54 0.8 0.13 0.13 0.13

Trays Tray Stack Extra Feed Space Diseng Space Skirt Ht Total Ht

7 3.6 1.5 3 1.5 9.6

Bottom of the Column 1

ρl

(kg/m3)ρg

(Kg/m3) V'(Kg/h) L' (Kg/h)σb(dyne/cm

) Flv Csb Uf(ft/s) m/s Db(m)

920.29 2.59 13841.91 253198.29 8.76 0.97 0.10 2.13 0.65 1.97

Top of the Column 1

ρl

(kg/m3)ρg

(Kg/m3) V'(Kg/h) L' (Kg/h)σb

(dyne/cm) Flv Csb Uf(ft/s) m/s Db(m)

923.27 0.61 2565.37 1282.68 0.16 0.01 0.29 4.25 1.04 1.93



Column 2

αlk/hk

(avg) N1 N2 βlk

(ξlk) βhk

(1-ξhk) YN NT YR R1 R2 R

4.69 16.29 16.29 0.999 0.999 0.80 16.29 0.80 0.85 0.85 0.85


50 24.5 1.5 3 1.5 30.5

ρl(kg/m3) ρg (Kg/m3) V'(Kg/h) L' (Kg/h) Uv (m/s) Dc (m)

568.87 2.27 87064.30 320560.95 0.71555296 4.35772



αlk/hk

(avg) N1 N2βlk

(ξlk) βhk

(1-ξhk) YN NT YR R1 R2 R N

4.31 11.94 17.53 0.99 1.00 0.80 16.41 0.80 0.75 0.94 0.90 20.51

Column 3


20 9.5 1.5 3 1.5 15.5

ρl (kg/m3) ρg (Kg/m3) V' (Kg/h) L' (Kg/h) Uv (m/s) Dc (m)

634.78 2.28 11536.25 11769.68 0.75364887 1.54079



Condensers and Reboilers

CondenserNo.

Qc (kJ/hr)

Tcond (K)

Circulating Cooling water Overall heattransferCoefficient (U)(kJ/hr.m2.0K)

Area (m2)

Tin (K) Tout (K) Amount(kmol)

C1 3,82E+05 372 298 350 97.4 4292,767 2.075829299

C2 2,88E+06 473,89 298 373 509 1430,922 14.91611077

C3 3,12E+06 474,58 298 373 552 1430,922 16.07479274

ReboilerNo.

QB (kJ/hr)

Treb (K)

Circulating Steam Overall heattransferCoefficient (U)(kJ/hr.m2.0K)

Area(m2)

Tin (K) Tout (K) Amount(kmol)

R1 7,73E+07 468,93 1000 488,93 1010 4292,767 101.7213

R2 8,73E+07 536,54 1000 556,54 1170 1430,922 131.6394

R3 2,92E+06 532,92 1000 552,92 39.2 1430,922 4.368935





Pumps

Pump 1

Right

Elbows Leq

Gate

Valves Leq Check Leq Z1-Z2 I.D. (m)

Ac

(m2) Length

2.00 64.00 1.00 7.00 1.00 170.00 4.00 1.00 0.79 25.00

Velocity μ Re e e/D R/ρv2 Hf ΔPtotal (m) Wp (KW)

7.06 5.69E-05 1494573.9 4.60E-05 4.60E-05 1.50E-03 75.74 79.74 87.15

Pump 2

RightElbows Leq

GateValves Leq Check Leq Z1-Z2 I.D.(m) Ac(m2) Length

2.00 64.00 1.00 7.00 1.00 170.00 25.00 1.00 0.79 25.00


8.977745 1.56E-05 5.42E+06 0.000046 4.60E-05 0.001125 79.13129 104.13129 113.2027



Pump 3

RightElbows Leq

GateValves Leq Check Leq Z1-Z2 I.D. (m)

Ac(m2) Length

2 10.47 1 1.15 1 27.82 6.5 0.16

0.021

029 25


7.97 5.06E-07 2.51E+07 0.000046 2.81E-04 1.75E-03 162.89 169.39 4.51

Pump 4

RightElbows Leq

GateValves Leq Check Leq Z1-Z2 I.D. (m) Ac (m2) Length

2 64.00 1 7.00 0 0.00 0 1.00 0.7855 25

Velocity μ Re e e/D R/ρv2 Hf ΔPtotal (m) Wp

(KW)

8.79 1.00E-06 8.27E+07 0.000046 4.60E-05 1.25E-03 11.86 11.86 12.58



Pump 5

RightElbows Leq

GateValves Leq Check Leq Z1-Z2

I.D.(m) Ac(m2) Length

0 0.00 1 0.15 0 0.00 0 0.02 0.00038 25


(KW)

10.11 5.90E-05 3.54E+04 0.000046 2.09E-03 2.75E-03 1230.65 1230.65 0.73

Pump 6

RightElbows Leq

GateValves Leq Check Leq Z1-Z2

I.D.(m) Ac (m2) Length

2 64.00 1 7.00 1170.0

0 14 1.00 0.7855 25


(KW)

8.80 5.90E-05 1.40E+06 0.000046 4.60E-05 2.75E-03 176.27 190.27 202.10



Costing and Project Evaluation

Column

# Type

Height

(Ft) Diameter BC($US) UF MF MPF BMC($US)1 D. C 31.68 3.531 6342.824 3.86261 4.23 1 103634.4

2 D. C 91.5 14.52 66095.23 3.86261 4.23 1 1079919

3 D. C 51.15 5.082 13704.26 3.86261 4.23 1 223911.6

4 Abs 56.1 13.959 42668.86 3.86261 4.23 1 697159.4

5 Abs 56.1 13.959 42668.86 3.86261 4.23 1 697159.4

6 F.D. 56.1 14.025 42880.72 3.86261 4.23 1 700620.9

7 F.D. 56.1 14.025 42880.72 3.86261 4.23 1 700620.9

Stack Ht BC UF MF MPF BMC($US) Total($US)

11.88 485.068 3.862609 1 1.4 2623.079 106257.44

80.85 24214.76 3.862609 1 1.4 130945 1210864.1

31.35 2108.044 3.862609 1 1.4 11399.57 235311.14

36.3 10517.86 3.862609 1 1.4 56876.9 754036.34

36.3 10517.86 3.862609 1 1.4 56876.9 754036.34

Distillation Columns, Flash Drum and Absorber



Heat Exchangers

HX Area(ft2) BC($US) MF MPF UF BMC Total

Reactor 8094.5 35314.62 3.29 2.529 3.862609 657343.2$1,756,010.1

8

H1 5930 28848.06 3.29 2.529 3.862609 536975.2

H2 688 7113.175 3.29 0.85 3.862609 86272.79

H3 570 6294.332 3.29 0.85 3.862609 76341.38C1 26.36 311.4985 1.83 0.85 3.862609 2021.371

C2 63.14 318.0975 1.83 0.85 3.862609 2064.192

C3 83.14 320.2053 1.83 0.85 3.862609 2077.871

R1 1095.00 9621.664 3.29 2.529 3.862609 179096.8

R2 1417.00 11376.83 3.29 2.529 3.862609 211767.4

R3 47.40 315.9157 1.83 0.85 3.862609 2050.035



Pumps

#Capacity

(Hp)D

(inches)Cost($US)

Motor($US) a1 a2 a3 HP Total Cost

1 117 39 34917.21 7675.11 4.81 0.510 0.05 7.5-250 42592.324

2 150 39 34917.21 9908.91 5.41 0.312 0.10 1-7.5 44826.124

3 6.0434 6 11823.23 911.52 12734.749

4 17 39 34917.21 1319.25 36236.466

5 1 1 4419.90 369.21 4789.1126

6 271 39 34917.21 18687.5 53604.726

Total $194,783.5



Compressors

Number Capacity BC($US) MF MPF UF BMC Total

1 11 4203.395 3.11 1 3.862609 50494.18 $64,082.38

2 2 1131.152 3.11 1 3.862609 13588.2

Mixers and Splitters

Mixer/Splitter Area Height Diameter

BC($US) UF MF MPF BMC($US)

1 7.32 12.22 3.055 2518.5 3.8626 4.23 1 41149.5062 20.71 20.54 5.13 4182.2 3.8626 3.18 1 51370.603

3 24.16 22.19 5.54 4789.6 3.8626 3.18 1 58831.994

4 10.62 14.71 3.67 3548.2 3.8626 4.23 1 57975.006

Total 209327.11

C ti f ti P j t



Costing of entire Project

Fix Capital Capital Investment

Equipment + PI Building and Site Working Capital Fixed and Working Capital

10,864,669.30 4,345,867.72 2,950,844.18 18,161,381.21

Raw Materials Unit Amount Price ($US) Total

n-Butane Kmol/h 876000 2.3481692 2,056,996.22

Succinic Anhydride Kmol/h 11563.2 800.592 9,257,405.41

Maintenance % Plant Cost 5 908,069.06

Labour $US/man*yr 15 40000 600,000.00

Manager $US/man*yr 1 200000 200,000.00Insurance %Plant Cost 2 363,227.62

Lab Analyses $US/man*yr 1 70000 70,000.00

Steam 0 0 0

Cooling Water $US/Kmol/h 10147584 0.017488189 177462.8666

Plant Overheads %Labour Cost 50 300000

Taxes % Fix Capital 2 304210.7405

Total Operating Cost 14,237,371.93

Revenue $US(Kmol Product/h) 501948 44.1261 22,149,007.64

Profit After Tax 7,911,635.72

ROI Pay out Time NPV Rate of Return (NPV = 0) IRR NPV = 0 N i

0.435629627 2.210530747 56,420,932.00 $0.00 0.43562 2E-07 2.7 0.1



ECONOMIC EVALUATION

With all the equipment size and cost, we now proceed to assess theeconomic viability of the project

The capital investment is calculated. The capital Investment which is all thecost incurred at the beginning of the plant life is composed of twocomponents: Fix capital and working capital.

The equipment cost plus 25% contingency, represents a part of the fixcapital investment. The other component is the cost for building and site,this is generally 40% of the bare module cost

The working capital is all the funds require to operate the plant due to

delays in payment and maintenance of inventories The other cost to consider is the cost of operating the plant. These costsare continuous over the entire life of the plant. These costs are brokendown into the following parts:

Raw material costs



Cost of utilitiesLabour SupervisionLaboratory analysesMaintenancePlant Overheads/Supplies e.g. Office supplies and spares and sales costsetc.TaxesInsuranceThe net revenue generated by operating the plant, will be the amount made

by selling the product produced, minus all the operating expensesSteam utility and electricity was not included in utility cost, because with heat integration, it was obvious that there are large amounts of heat avail able for the process that could be used for generating steam and electricity to operate the plant The project was evaluated in terms of the following markers:

Net Present Worth An internal rate of return, IRR, also refers to as the minimum attractive rateof return, MARR, was computed

The minimum payback period, at NPV = 0 was computed



The process is found to be highly profitable

The MARR is 43.5%, well above the 10% interest rate used for computingthe NPV.

Pay back period is computed to be 2.7 yearsNPV is computed to be highly positive



Sensitivity Analyses

Sensitivity analyses were done, using the following markers:

A sharp increase in raw material cost. A 50% increase in the price of butane was used. The process remained profitable

A 50% decrease in product price. The product was no longer profitable.This indicates that the profitability of the process is highly sensitive to sale

price of the product. The minimum price the product can be sold for and theprocess remains profitable is $32.5/Kmol. This represent a 26% decreasein current selling price.

High increase in interest rates. If the interest rates exceeds the MARR,then the process no longer remains profitable. Doing the analyses with aninterest rate of 50%, the process becomes highly non-profitable, with a

highly negative NPV and a pay back period of over a hundred years.

Sensitivity analyses



Sensitivity analyses

1) A sharp increase in raw material cost. A 50% increasein the price of butane was used.


n-Butane Kmol/h 876000 3.5222538 3,085,494.33



Labour $US/man*yr 15 40000 600,000.00

Manager $US/man*yr 1 200000 200,000.00

Insurance %Plant Cost 2 363,227.62


Steam 0 0 0




Total Op. Cost 15,265,870.03

Revenue $US (Kmol Product/h) 501948 44.1261 22,149,007.64


ROI Pay out Time NPV NPV = 0(IRR) IRR NPV = 0 N i

0.379 2.526854178 46,725,368.29 ($0.00) 0.37897 5.59E-08 3.2 0.1



2) A 50% decrease in product cost.


n-Butane Kmol/h 876000 2.3481692 2,056,996.22

SuccinicAnhydride Kmol/h 11563.2 800.592 9,257,405.41


Labour $US/man*yr 15 40000 600,000.00

Manager $US/man*yr 1 200000 200,000.00Insurance %Plant Cost 2 363,227.62


Steam 0 0 0




Total Op. Cost 14,237,371.93

Revenue$US (KmolProduct/h) 501948 22.06305 11,074,503.82

Profit After Tax -3,162,868.10



3) High increases in interest rates


n-Butane Kmol/h 876000 2.3481692 2,056,996.22



Labour $US/man*yr 15 40000 600,000.00

Manager $US/man*yr 1 200000 200,000.00

Insurance %Plant Cost 2 363,227.62


Steam 0 0 0




Total Op Cost 14,237,371.93

Revenue$US (KmolProduct/h) 501948 44.1261 22,149,007.64


NPV

-2,338,192.29

H t I t ti



Heat Integration

No. Streams Condition FlowTin(K)

Tout(K)

Enthalpy in(kJ)

Enthalpy out(kJ)

AvailableHeat

1 u2 Hot 12396.4 619.3 400 9.84E+06 3.07E+06 -6.77E+062 Lo Hot 2336.97 536 400 -5.05E+08 -5.56E+08 -5.10E+07

3He1coolingwater Cold 14800 298 373 0 1.59E+07 1.59E+07

4He2coolingwater Cold 2100 298 373 0 1.59E+07 1.59E+07

5 C1 Cold 67.6 298 350 0 1.19E+07 1.19E+07

6 C2 Cold 509 298 373 0 1.59E+07 1.59E+07

7 C3 Cold 552 298 373 0 1.59E+07 1.59E+07

8 R1 Hot 15900 1000 488.93 8.10E+07 6.34E+07 -1.76E+07

9 R2 Hot 18000 1000 556.54 8.10E+07 6.57E+07 -1.53E+07

10 R3 Hot 602 1000 552.92 8.10E+07 6.55E+07 -1.54E+0711 u31 Hot 12396.4 400 394 3.83E+07 3.56E+07 -2.70E+06

12Ccoolingwater Cold 467 298 373 0 1.59E+07 1.59E+07



The PA tool box of ICAS was used to generate the Pinch Diagrams after givingall the streams input data.

The Diagrams shows an additional cooling of 1.0811E11 kJ/hr.

So this is the amount of excess heat which can be used for other purposes.

The pinch point is at 394K for the hot stream and 383K for the cold streamobtained from the cascade diagram.

The results shows an additional of 3 heat exchangers are needed to satisfy thecondition.

The heat duties have been added up and found that the process has excessheat than required in the process. This can be attributed to the highlyexothermic reactions in the reactor.

E i t l I t A l i



Streams In Streams Out

µ01 µ92

S0 µ61

µp

Environmental Impact Analysis

TotalPEI HTPI HTPE ATP TTP GWP ODP PCOP AP

InputSum 8740.31 1696.89 0.90263 448.692 1696.89 0 0 4896.93 0

OutputSum 12064.7 5060.28 1668.93 98.6746 5060.28 0.21655 0 176.289 0

ImpactGenerated 3324.36 3363.39 1668.02 -350.017 3363.39 0.21655 0 -4720.64 0

Analysis:



Analysis:

The Report generated gives a higher value of the Total Potential EnvironmentalImpact suggesting that the process has to be modified for environmentalpurposes. The high value of the PEI is because of the excess amounts of carbon

dioxide released into the atmosphere.

By analyzing all the individual output streams, it can be clearly observed thatoutput stream 3 has quiet high values of the total PEI. It is because of the releaseof the purge gas from the splitter directly into the atmosphere.

As a process improvement step, we can use incinerator to convert the Carbonmonoxide to carbon dioxide before it is released into the atmosphere. As analternative a scrubber can be used to scrub all the harmful gases and preventthem from entering into the atmosphere.

The other 2 outlet streams mostly contain water other than the product, so theyhave less environmental impact.

Changing the solvent in the absorption column from Succinic anhydride to water can increase the environmental attractiveness of the process but the requiredproduct yield cannot be attained.

Documents

processdesignslides.204153727 (1)