PROCESS DESIGN OF MALEIC ANHYDRIDE PLANT

1
PROCESS DESIGN OF MALEIC ANHYDRIDE PLANT

• BY
• WORIL TURNER DUDLEY
• VIJAYA KRISHNA BODLA

2
TABLE OF CONTENTS

• Introduction
• The five processes selected
• Product and Process Selected for Design
• Screening of Process Alternatives
• Material Balance
• Energy Balance
• Equipment Sizing
• Equipment Costing
• Heat Intergration
• Economic evaluation
• 11. Environmental Analysis
• 12. Conclusion

3
INTRODUCTION

• The process design project involves designing a
  process plant for producing a particular product.
• A list of five products are selected, then from
  that list one product is selected for the design
  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 an economic
  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 cooling
  required.
• An environmental analysis was also done to
  determine the environmental impact of effluent
  discharge streams.

4
Product Names Raw Materials Process References
Maleic Anhydride n-butane and air Huntsmann fixed bed maleic anhydride process, Kirk Othmer Encyclopaedia of Chemical Technology by Timothy R. Felthouse, Joseph C. Burnett, Ben Horrell, Micheal J. Mummey and Yeong-Jen Kuo
Citric Acid Sucrose or dextrose Fermentation of sugars to produce citric acid, Shreves Chemical Process Industries 5th Edition by George T. Austin, page 598
Acetaldehyde Oxygen, water and ethylene Oxidation of ethylene to produce acetaldehyde, Shreves Chemical Process Industries
Ammonia Nitrogen and Hydrogen Ammonia process, Shreves Chemical Process Industries 5th Edition by George T. Austin page 306
Cinnamic Aldehyde Water and aldol Cinnamic aldehyde production by aldol condensation, Shreves Chemical Process Industries 5th Edition by George T. Austin page 494
5
PRODUCT SELECTED

• The 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 a variety of chemical
  reactions.
• Maleic anhydride's largest use today is in the
  production of unsaturated polyester resins.
• Another significant use is in the manufacture of
  alkyd resins, which are in turn used in paints
  and coatings.
• Other applications where maleic anhydride is used
  include the production of agricultural chemicals,
  maleic acid, copolymers, fumaric acid, lubricant
  additives, surfactants and plasticizers.
• Future applications are anticipated to be
  numerous given the versatility and usefulness of
  the product.

6

• 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)

7
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 even looking at
  raw material costs, benzene is rejected as the
  raw material for the maleic 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 remain in
  the vapor phase. We then consider separating
  water from maleic anhydride by exploiting the
  differences between their physical properties.
• A solvent can be used for the product recovery,
  by contacting the product gases with a liquid
  solvent and then separating the maleic anhydride
  from the solvent. A number 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 maleic anhydride
  produced.

8
Process Flow sheet
9
MASS BALANCE FLOW CHART
10
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 1.5 times the amount
  required
• Using the yield of Maleic Anhydride, percentage
  conversion of butane and the 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 to form Carbon Dioxide
  as for Carbon Monoxide
• Split factors are then specified for all the
  separation equipment as well as for the purge
• The absorber is specified to be an isothermal
  absorber
• It is assumed that the solvent entering the
  column doesnot contain any Maleic Anhydride
• For the absorber mass balance model, the Kremser
  equation is used to determine the number of
  stages, using the split factor for the key
  component recovery.

11

• The remaining split factors for the other
  components are calculated from the Kremser
  relationship
• To solve the mass balance model, the flow sheet
  is partitioned into two modules and the recycle
  broken by tearing the inlet stream to the
  reactor.
• Once the component flows to the reactor are
  calculated, from the mass balance model, we can
  sequentially calculate all the other flowrates
• With the flowrates calculated, distillation
  column temperatures can be calculated.
• Temperature for the vapor leaving top of the
  column is found from a dew point calculation.
  The temperature in the condenser and reboiler is
  calculated from a bubble point calculation
• It is assumed that the distillation column
  operates at one atmosphere of pressure.

12
SUMMARY - MASS BALANCE
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
Succinic Anhydride 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
13
SUMMARY - MASS BALANCE
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 0.0 440.4 110.1  
Butane 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 
14
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 the streams are
  considered.
• It is assumed that there is no ?H of mixing, or
  pressure effects on ?H.
• 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 enthalpy correlations
• To calculate the enthalpies of vapor mixtures the
  following correlation is used,
• ?Hv (T, y) ?Hf ?HT ?k yk Hf, k (T1) ?k yk
  ? Cpo, k (T) dT
  
• To calculate the enthalpies of liquid mixtures
  the following correlation is used,
• ?HLk (T) ?Hof, k ?Cpo, k (T) dT - ?Hkvap
• For the reactor the following expression is used,
• QR µ2?Hv (T, y2) µ1?Hv (T, y1)
  
  
  
• QR is the heat of reaction, which is positive for
  an endothermic reaction and negative for an
  exothermic reaction.
• ?Hv (T, y2) ?Hout and ?Hv (T, y1) ?Hin
• With the stream enthalpies, and stream
  temperatures, the heating and cooling duties can
  now be calculated
  
  

15
SUMMARY - ENERGY BALANCE
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
16
Sizing

• All the process equipments are sized for cost
  considerations based on the procedure given in
  the text book.
• Splitter is sized as a reverse mixer that the
  output flow is considered for sizing other than
  the input flow considered for mixers.
• For the heat exchangers the area of heat transfer
  and the amount of cooling water required have
  been calculated. The values of overall heat
  transfer coefficients are obtained from the book.
  
• Since nitrogen is the one carrying the maximum
  heat, it is considered as the main component from
  which heat has to be removed.
• Sizing of compressors is based on the assumption
  that the expansion is ideal, isentropic and
  adiabatic giving a Gamma value of 1.4 for an
  ideal system.
• Reactor Design is done by assuming a Space
  velocity or residence time and also that the
  reactor volume is twice that of the volume
  occupied by the catalyst.

17
Sizing
Reactor
Reactor Volume (m2) Outside tube diameter (m) Inside tube diameter (m) Inner cross sectional area (m2) Tube Length (m) Inside volume of each tube (m2) No. of tubes Outside surface area of each tube (m2)
Multi Tubular 4.492676 0.035 0.02892 0.0006565 6.09 0.003998 1123.625 0.669291
Mixers
Mixer Fl ? (residence time) (1/hr) Temp (K) Molar Density of the 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
18
Splitter
Fl ? (residence time) (1/hr) Temp (K) Molar Density of 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 ? µ W (kJ/hr) ?c ?m (Shaft driven) Wb (shaft driven) (watts) Wb (shaft driven) (watts)
CM1 (For stream R1) 200 109 393 1,4 9816,7 2125869 0,8 0,9 8,201658291 8,201658291
CM2 (For Air Compression) 200 103,2 298 1,4 2763 4981751 0,8 0,9 1,921971726  
19
Heat Exchangers
Heat Exchangers Q U T1 T2 Delta Tln Area(A) (m2) Amount of cooling water(Kmol/hr)
HE1 8,38E07 919,8756 619,3 400 163,6839 556.5563 14800
HE2 1,19E07 1430,922 536 400 130,1257 63.90984 2100
Condenser
Condenser Q U T1 T2 Delta Tln Area(A) (m2) Amount of cooling water (Kmol/hr)
C 2,64E06 919,8756 400 394 54,39447 52,76186 467
20

• For the distillation columns, the number of trays
  and the reflux ratio were determined 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 which when used in ICAS was
  giving tray number in excess of 100, the method
  of Underwood was used to determine the minimum
  reflux ratio and the heuristic of the reflux
  ratio being 1.2 the minimum reflux ratio used
  to get the 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
  the height 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 to determine the maximum
  allowable vapor velocity based on the column
  crosssectional area.

21

• For the absorber and flash drum, number of
  theoretical stages calculated by the 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 Vj and 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 column diameter as done for the
  absorber and the distillation column.

22
Column 1
alk/hk (avg) N1 N2 ßlk (?lk) ßhk (1-?hk) 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
?l (kg/m3) ?g (Kg/m3) V'(Kg/h) L' (Kg/h) sb(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
?l (kg/m3) ?g (Kg/m3) V'(Kg/h) L' (Kg/h) sb (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
23
Column 2
alk/hk (avg) N1 N2 ßlk (?lk) ßhk (1-?hk) 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
Trays Tray Stack Extra Feed Space Diseng Space Skirt Ht Total Ht
50 24.5 1.5 3 1.5 30.5
?l(kg/m3) ?g (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
24
Column 3
alk/hk (avg) N1 N2 ßlk (?lk) ßhk (1-?hk) 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
Trays Tray Stack Extra Feed Space Diseng Space Skirt Ht Total Ht
20 9.5 1.5 3 1.5 15.5
?l (kg/m3) ?g (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
25
Condensers and Reboilers
Condenser No. Qc (kJ/hr) Tcond (K) Circulating Cooling water Circulating Cooling water Circulating Cooling water Overall heat transfer Coefficient (U) (kJ/hr.m2.0K) Area (m2)
Condenser No. Qc (kJ/hr) Tcond (K) Tin (K) Tout (K) Amount (kmol) Overall heat transfer Coefficient (U) (kJ/hr.m2.0K) Area (m2)
C1 3,82E05 372 298 350 97.4 4292,767 2.075829299
C2 2,88E06 473,89 298 373 509 1430,922 14.91611077
C3 3,12E06 474,58 298 373 552 1430,922 16.07479274
Reboiler No. QB (kJ/hr) Treb (K) Circulating Steam Circulating Steam Circulating Steam Overall heat transfer Coefficient (U) (kJ/hr.m2.0K) Area (m2)
Reboiler No. QB (kJ/hr) Treb (K) Tin (K) Tout (K) Amount (kmol) Overall heat transfer Coefficient (U) (kJ/hr.m2.0K) Area (m2)
R1 7,73E07 468,93 1000 488,93 1010 4292,767 101.7213
R2 8,73E07 536,54 1000 556,54 1170 1430,922 131.6394
R3 2,92E06 532,92 1000 552,92 39.2 1430,922 4.368935
26
Absorber
Trays Tray Stack Extra Feed Space Diseng Space Skirt Ht Total Ht
23 11 1.5 3 1.5 17
?l (kg/m3) ?g (Kg/m3) V' (Kg/h) L' (Kg/h) Uv (m/s) Dc (m)
1157.01 1.28 351433.04 231038.67 1.36 4.23
27
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/?v2 Hf ?Ptotal (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
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 25.00 1.00 0.79 25.00
Velocity µ Re e e/D R/?v2 Hf ?Ptotal (m) Wp (KW)
8.977745 1.56E-05 5.42E06 0.000046 4.60E-05 0.001125 79.13129 104.13129 113.2027
28
Pump 3
Right Elbows Leq Gate Valves Leq Check Leq Z1-Z2 I.D. (m) Ac (m2) Length
2 10.47 1 1.15 1 27.82 6.5 0.16 0.021029 25
Velocity µ Re e e/D R/?v2 Hf ?Ptotal (m) Wp (KW)
7.97 5.06E-07 2.51E07 0.000046 2.81E-04 1.75E-03 162.89 169.39 4.51
Pump 4
Right Elbows Leq Gate Valves 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/?v2 Hf ?Ptotal (m) Wp (KW)
8.79 1.00E-06 8.27E07 0.000046 4.60E-05 1.25E-03 11.86 11.86 12.58
29
Pump 5
Right Elbows Leq Gate Valves 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
Velocity µ Re e e/D R/?v2 Hf ?Ptotal (m) Wp (KW)
10.11 5.90E-05 3.54E04 0.000046 2.09E-03 2.75E-03 1230.65 1230.65 0.73
Pump 6
Right Elbows Leq Gate Valves Leq Check Leq Z1-Z2 I.D. (m) Ac (m2) Length
2 64.00 1 7.00 1 170.00 14 1.00 0.7855 25
Velocity µ Re e e/D R/?v2 Hf ?Ptotal (m) Wp (KW)
8.80 5.90E-05 1.40E06 0.000046 4.60E-05 2.75E-03 176.27 190.27 202.10
30
Costing and Project Evaluation
Distillation Columns, Flash Drum and Absorber
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
31
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 1,756,010.18
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 76341.38  
C1 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  
32
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
33
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 41149.506
2 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
34
Costing of entire Project
Fix Capital 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/manyr 15 40000 600,000.00
Manager US/manyr 1 200000 200,000.00
Insurance Plant Cost 2   363,227.62
Lab Analyses US/manyr 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
35
ECONOMIC EVALUATION

• With all the equipment size and cost, we now
  proceed to assess the economic viability of the
  project
• The capital investment is calculated. The capital
  Investment which is all the cost incurred at the
  beginning of the plant life is composed of two
  components Fix capital and working capital.
• The equipment cost plus 25 contingency,
  represents a part of the fix capital 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 costs are continuous
  over the entire life of the plant. These costs
  are broken down into the following parts
• Raw material costs

36

• Cost of utilities
• Labour
• Supervision
• Laboratory analyses
• Maintenance
• Plant Overheads/Supplies e.g. Office supplies and
  spares and sales costs etc.
• Taxes
• Insurance
• The net revenue generated by operating the plant,
  will be the amount made by selling the product
  produced, minus all the operating expenses
• Steam utility and electricity was not included in
  utility cost, because with heat integration, it
  was obvious that there are large amounts of heat
  available 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 rate of return, MARR,
  was computed
• The minimum payback period, at NPV 0 was
  computed

37

• The process is found to be highly profitable
• The MARR is 43.5, well above the 10 interest
  rate used for computing the NPV.
• Pay back period is computed to be 2.7 years
• NPV is computed to be highly positive

38
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 the process
  remains profitable is 32.5/Kmol. This represent
  a 26 decrease in 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 an interest rate of 50, the process becomes
  highly non-profitable, with a highly negative NPV
  and a pay back period of over a hundred years.

39
Sensitivity analyses

• A sharp increase in raw material cost. A 50
  increase
• in the price of butane was used.

Raw Materials Unit Amount Price (US) Total
n-Butane Kmol/h 876000 3.5222538 3,085,494.33
Succinic Anhydride Kmol/h 11563.2 800.592 9,257,405.41
Maintenance Plant Cost 5   908,069.06
Labour US/manyr 15 40000 600,000.00
Manager US/manyr 1 200000 200,000.00
Insurance Plant Cost 2   363,227.62
Lab Analyses US/manyr 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 Op. Cost 15,265,870.03
Revenue US (Kmol Product/h) 501948 44.1261 22,149,007.64
      Profit After Tax 6,883,137.61
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
40
2) A 50 decrease in product cost.
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/manyr 15 40000 600,000.00
Manager US/manyr 1 200000 200,000.00
Insurance Plant Cost 2   363,227.62
Lab Analyses US/manyr 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 Op. Cost 14,237,371.93
Revenue US (Kmol Product/h) 501948 22.06305 11,074,503.82
      Profit After Tax -3,162,868.10
41
3) High increases in interest rates
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/manyr 15 40000 600,000.00
Manager US/manyr 1 200000 200,000.00
Insurance Plant Cost 2   363,227.62
Lab Analyses US/manyr 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 Op Cost 14,237,371.93
Revenue US (Kmol Product/h) 501948 44.1261 22,149,007.64
      Profit After Tax 7,911,635.72
NPV
-2,338,192.29
42
Heat Integration
No. Streams Condition Flow Tin (K) Tout (K) Enthalpy in (kJ) Enthalpy out (kJ) Available Heat
1 u2 Hot 12396.4 619.3 400 9.84E06 3.07E06 -6.77E06
2 Lo Hot 2336.97 536 400 -5.05E08 -5.56E08 -5.10E07
3 He1cooling water Cold 14800 298 373 0 1.59E07 1.59E07
4 He2cooling water Cold 2100 298 373 0 1.59E07 1.59E07
5 C1 Cold 67.6 298 350 0 1.19E07 1.19E07
6 C2 Cold 509 298 373 0 1.59E07 1.59E07
7 C3 Cold 552 298 373 0 1.59E07 1.59E07
8 R1 Hot 15900 1000 488.93 8.10E07 6.34E07 -1.76E07
9 R2 Hot 18000 1000 556.54 8.10E07 6.57E07 -1.53E07
10 R3 Hot 602 1000 552.92 8.10E07 6.55E07 -1.54E07
11 u31 Hot 12396.4 400 394 3.83E07 3.56E07 -2.70E06
12 Ccooling water Cold 467 298 373 0 1.59E07 1.59E07
43
The PA tool box of ICAS was used to generate the
Pinch Diagrams after giving all 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 stream obtained from the
cascade diagram. The results shows an additional
of 3 heat exchangers are needed to satisfy the
condition. The heat duties have been added up
and found that the process has excess heat than
required in the process. This can be attributed
to the highly exothermic reactions in the
reactor.
44
Environmental Impact Analysis
Streams In Streams Out
µ01 µ92
S0 µ61
µp
Total PEI HTPI HTPE ATP TTP GWP ODP PCOP AP
Input Sum 8740.31 1696.89 0.90263 448.692 1696.89 0 0 4896.93 0
Output Sum Output Sum 12064.7 5060.28 1668.93 98.6746 5060.28 0.21655 0 176.289 0
Impact Generated Impact Generated 3324.36 3363.39 1668.02 -350.017 3363.39 0.21655 0 -4720.64 0
45
Analysis The Report generated gives a higher
value of the Total Potential Environmental Impact
suggesting that the process has to be modified
for environmental purposes. 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 that output stream 3 has
quiet high values of the total PEI. It is because
of the release of the purge gas from the splitter
directly into the atmosphere. As a process
improvement step, we can use incinerator to
convert the Carbon monoxide to carbon dioxide
before it is released into the atmosphere. As an
alternative a scrubber can be used to scrub all
the harmful gases and prevent them from entering
into the atmosphere. The other 2 outlet streams
mostly contain water other than the product, so
they have 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 required product yield cannot be attained.