PROCESS DESIGN OF
MALEIC ANHYDRIDE
PLANT
BY
WORIL TURNER DUDLEY
VIJAYA KRISHNA BODLA
TABLE OF CONTENTS
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
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
Environmental Analysis
Conclusion
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.
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
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.
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, nbutane 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.
Process Flow sheet
Purge (μ92)
R1
9. Splitter
μ41
μ42
CM2
4. Condenser
C1
µ31
µ01
µ1
1. Mixer
µ32
2. Reactor
μ61
µ51
3. Absorber
5. Mixer 2
6. Distillation
μp
C3
µ2
R1
μ62
µ81
μ71
7. Distillation
Final Column
CM1
8. Mixer 3
HE1
So
HE2
C2
μ72
R2
μS
R3
MASS BALANCE FLOW CHART
Purge (μ92)
R1
9. Splitter
μ41
μ42
4. Condenser
µ31
µ01
µ1
µ2
µ32
µ51
μ61
1. Mixer
2. Reactor
3. Absorber
5. Mixer 2
6. Distillation
μp
μ62
μ71
µ81
7. Distillation
8. Mixer 3
So
μ72
μS
Final Column
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.
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.
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
μp
μfs
S0
R1
μ92
Maleic Anhydride
57.3
0.1
0.0
0.1
0.0
Succinic Anhydride
0.0
2.3
1.3
5.7
1.4
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
0.0
0.0
0.0
1.0
1.0
Components
Vapor Fraction
Lo
2337.0
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 the
streams 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 enthalpy
correlations:
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) dT
To calculate the enthalpies of liquid mixtures the following correlation is used,
ΔHLk (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 negative
for an exothermic reaction.
ΔHv (T, y2) = ΔHout and ΔHv (T, y1) = ΔHin
With the stream enthalpies, and stream temperatures, the heating and cooling duties
can 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
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.
Sizing
Reactor:
Reacto
r
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
Mixer
Fl
Τ
(residence
time) (1/hr)
Temp
(K)
Molar
Density of
the flow
(kmol/m3)
Volume
(m3)
Diamete
r (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
Mixers:
Splitter:
Splitter
Fl
Τ
(residence
time) (1/hr)
Temp
(K)
Molar Density
of the flow
(kmol/m3)
Volume
(m3)
Diameter
(m)
Length
(m)
12270,8
0,033333
393
184,7362
4,42822
1,1214
4,48566 0,98719
Area (m2)
Compressors:
ηc
ηm
(Shaft
driven)
Compressor
P2
P1
T1
γ
µ
W
(kJ/hr)
CM1 (For
stream R1)
200
109
393
1,4
9816,7
2125869
0,8
0,9
CM2 (For Air
Compression)
200
103,2
298
1,4
2763
4981751
0,8
0,9
Wb (shaft
driven) (watts)
8,201658291
1,921971726
Heat Exchangers:
Heat
Exchangers
Q
U
T1
T2
Delta Tln
Area(A) (m2)
Amount
of
cooling
water(Km
ol/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
Delta Tln
Condenser:
Condenser
Q
U
T1
T2
C
2,64E+06
919,8756
400
394 54,39447
Area(A)
(m2)
Amount of cooling
water (Kmol/hr)
52,76186
467
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.
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.
Column 1
αlk/hk
(avg)
N1
25.33
3.54
N2
βlk
(ξlk)
βhk
(1-ξhk)
YN
NT
YR
R1
R2
R
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)
920.29
2.59
13841.91
L' (Kg/h)
σb(dyne/cm
)
Flv
Csb
Uf(ft/s)
m/s
Db(m)
253198.29
8.76
0.97
0.10
2.13
0.65
1.97
Flv
Csb
Uf(ft/s)
m/s
Db(m)
0.01
0.29
4.25
1.04
1.93
Top of the Column 1
ρl
(kg/m3)
ρg
(Kg/m3)
V'(Kg/h)
L' (Kg/h)
σb
(dyne/cm)
923.27
0.61
2565.37
1282.68
0.16
Column 2
αlk/hk
(avg)
N1
4.69
16.29
N2
βlk
(ξlk)
βhk
(1-ξhk)
YN
NT
YR
R1
R2
R
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
ρl(kg/m3)
ρg (Kg/m3)
V'(Kg/h)
L' (Kg/h)
568.87
2.27
87064.30
320560.95
Uv (m/s)
Dc (m)
0.71555296
4.35772
Column 3
αlk/hk
(avg)
N1
4.31
11.94
N2
βlk
(ξlk)
βhk
(1-ξhk)
YN
NT
YR
R1
R2
R
N
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
ρl (kg/m3)
ρg (Kg/m3)
V' (Kg/h)
L' (Kg/h)
634.78
2.28
11536.25
11769.68
Uv (m/s)
0.75364887
Dc (m)
1.54079
Condensers and Reboilers
Overall heat
transfer
Coefficient (U)
(kJ/hr.m2.0K)
Area (m2)
97.4
4292,767
2.075829299
373
509
1430,922
14.91611077
373
552
1430,922
16.07479274
Condenser
No.
Qc
(kJ/hr)
Tcond
(K)
Circulating Cooling water
Tin (K)
Tout (K)
Amount
(kmol)
C1
3,82E+05
372
298
350
C2
2,88E+06
473,89
298
C3
3,12E+06
474,58
298
Reboiler
No.
QB
(kJ/hr)
Treb
(K)
Circulating Steam
Tin (K)
Tout (K)
Amount
(kmol)
R1
7,73E+07
468,93
1000
488,93
R2
8,73E+07
536,54
1000
R3
2,92E+06
532,92
1000
Overall heat
transfer
Coefficient (U)
(kJ/hr.m2.0K)
Area
(m2)
1010
4292,767
101.7213
556,54
1170
1430,922
131.6394
552,92
39.2
1430,922
4.368935
Absorber
Trays
Tray Stack
Extra Feed
Space
23
11
1.5
Diseng Space
Skirt Ht
Total Ht
3
1.5
17
ρl (kg/m3)
ρg (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
Pumps
Pump 1
Right
Elbows
2.00
Leq
Gate
Valves
Leq
Check
Leq
Z1-Z2
64.00
1.00
7.00
1.00
170.00
4.00
I.D. (m)
Ac
(m2)
Length
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
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
Pump 2
Velocity
μ
Re
e
e/D
R/ρv2
Hf
8.977745
1.56E-05
5.42E+06
0.000046
4.60E-05
0.001125
79.13129
ΔPtotal (m)
Wp (KW)
104.13129
113.2027
Pump 3
Right
Elbows
Leq
2
Gate
Valves
10.47
1
Leq
Check
1.15
1
Leq
Z1-Z2
27.82
I.D. (m)
Ac
(m2)
Length
0.16
0.021
029
25
6.5
Velocity
μ
Re
e
e/D
R/ρv2
Hf
ΔPtotal (m)
Wp (KW)
7.97
5.06E-07
2.51E+07
0.000046
2.81E-04
1.75E-03
162.89
169.39
4.51
Right
Elbows
Leq
Gate
Valves
Leq
Check
Leq
Z1-Z2
I.D. (m)
2
64.00
1
7.00
0
0.00
0
1.00
Pump 4
Ac (m2)
0.7855
Length
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
Right
Elbows
0
Leq
Gate
Valves
Leq
Check
Leq
0.00
1
0.15
0
0.00
Z1-Z2
I.D.
(m)
Ac(m2)
Length
0
0.02
0.00038
25
Velocity
μ
Re
e
e/D
R/ρv2
Hf
ΔPtotal (m)
Wp
(KW)
10.11
5.90E-05
3.54E+04
0.000046
2.09E-03
2.75E-03
1230.65
1230.65
0.73
Leq
Gate
Valves
Pump 6
Right
Elbows
2
64.00
1
Leq
7.00
Check
Leq
Z1-Z2
I.D.
(m)
1
170.0
0
14
1.00
Ac (m2)
0.7855
Length
25
Velocity
μ
Re
e
e/D
R/ρv2
Hf
ΔPtotal (m)
Wp
(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
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
Heat Exchangers
HX
Area(ft2)
BC($US)
MF
MPF
UF
BMC
Reactor
8094.5
35314.62
3.29
2.529
3.862609 657343.2
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
R3
47.40
315.9157
1.83
0.85
3.862609 2050.035
211767.4
Total
$1,756,010.1
8
Pumps
#
(Hp)
D
(inches)
1
117
39
34917.21
7675.11
4.81
0.510
0.05
7.5-250
42592.324
2
150
39
34917.21
9908.91
5.41
0.312
0.10
1-7.5
44826.124
3
6.0434
6
11823.23
911.52
12734.749
4
17
39
34917.21
1319.25
36236.466
5
1
1
4419.90
369.21
4789.1126
6
271
39
34917.21
18687.5
53604.726
Capacity
Cost
($US)
Motor
($US)
a1
a2
a3
HP
Total Cost
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
$64,082.38
2
2
1131.152
3.11
1
3.862609
13588.2
UF
MF
MPF
BMC($US)
Mixers and Splitters
Mixer/Splitter
Area
Height
Diameter
BC
($US)
1
7.32
12.22
3.055
2518.5
3.8626
4.23
1
41149.506
2
20.71
20.54
5.13
4182.2
3.8626
3.18
1
51370.603
3
24.16
22.19
5.54
4789.6
3.8626
3.18
1
58831.994
4
10.62
14.71
3.67
3548.2
3.8626
4.23
1
57975.006
Total
209327.11
Costing of entire Project
Fix Capital
Capital Investment
Equipment + PI
Building and Site
Working Capital
Fixed and Working Capital
10,864,669.30
4,345,867.72
2,950,844.18
18,161,381.21
Raw Materials
Unit
Amount
Price ($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
Labour
$US/man*yr
15
40000
600,000.00
Manager
$US/man*yr
1
200000
200,000.00
Insurance
%Plant Cost
2
Lab Analyses
$US/man*yr
1
70000
70,000.00
0
0
0
0.017488189
177462.8666
Steam
908,069.06
363,227.62
Cooling Water
$US/Kmol/h
10147584
Plant Overheads
%Labour Cost
50
300000
Taxes
% Fix Capital
2
304210.7405
Revenue
$US(Kmol Product/h)
501948
Total Operating Cost
14,237,371.93
44.1261
22,149,007.64
Profit After Tax
7,911,635.72
ROI
Pay out Time
NPV
Rate of Return (NPV = 0)
IRR
0.435629627
2.210530747
56,420,932.00
$0.00
0.43562
NPV = 0
2E-07
N
i
2.7
0.1
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
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
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
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 $32.5/Kmol. This represent a 26% decrease
in current selling price.
High increase in interest rates. If the interest rates exceeds the MARR,
then the process no longer remains profitable. Doing the analyses with an
interest rate of 50%, the process becomes highly non-profitable, with a
highly negative NPV and a pay back period of over a hundred years.
Sensitivity analyses
1) A sharp increase in raw material cost. A 50% increase
in the price of butane was used.
Raw Materials
Unit
Amount
Price ($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
Labour
$US/man*yr
15
40000
600,000.00
Manager
$US/man*yr
1
200000
200,000.00
Insurance
%Plant Cost
2
Lab Analyses
$US/man*yr
1
70000
70,000.00
0
0
0
0.017488189
177462.8666
Steam
908,069.06
363,227.62
Cooling Water
$US/Kmol/h
10147584
Plant Overheads
%Labour Cost
50
300000
Taxes
% Fix Capital
2
304210.7405
Revenue
$US (Kmol Product/h)
501948
Total Op. Cost
15,265,870.03
44.1261
22,149,007.64
Profit After Tax
6,883,137.61
ROI
Pay out Time
NPV
NPV = 0(IRR)
IRR
0.379
2.526854178
46,725,368.29
($0.00)
0.37897
NPV = 0
5.59E-08
N
i
3.2
0.1
2) A 50% decrease in product cost.
Raw Materials
Unit
Amount
Price ($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
Labour
$US/man*yr
15
40000
600,000.00
Manager
$US/man*yr
1
200000
200,000.00
Insurance
%Plant Cost
2
Lab Analyses
$US/man*yr
1
70000
70,000.00
0
0
0
0.017488189
177462.8666
Steam
908,069.06
363,227.62
Cooling Water
$US/Kmol/h
10147584
Plant Overheads
%Labour Cost
50
300000
Taxes
% Fix Capital
2
304210.7405
Revenue
$US (Kmol
Product/h)
501948
Total Op. Cost
14,237,371.93
22.06305
11,074,503.82
Profit After Tax
-3,162,868.10
3) High increases in interest rates
Raw Materials
Unit
Amount
Price ($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
Labour
$US/man*yr
15
40000
600,000.00
Manager
$US/man*yr
1
200000
200,000.00
Insurance
%Plant Cost
2
Lab Analyses
$US/man*yr
1
70000
70,000.00
0
0
0
0.017488189
177462.8666
Steam
908,069.06
363,227.62
Cooling Water
$US/Kmol/h
10147584
Plant Overheads
%Labour Cost
50
300000
Taxes
% Fix Capital
2
304210.7405
Revenue
$US (Kmol
Product/h)
501948
Total Op Cost
14,237,371.93
44.1261
22,149,007.64
Profit After Tax
7,911,635.72
NPV
-2,338,192.29
Heat Integration
Tin
(K)
Flow
Tout
(K)
Enthalpy in
(kJ)
Enthalpy out
(kJ)
Available
Heat
No.
Streams
Condition
1
u2
Hot
12396.4
619.3
400
9.84E+06
3.07E+06
-6.77E+06
2
Lo
Hot
2336.97
536
400
-5.05E+08
-5.56E+08
-5.10E+07
3
He1cooling
water
Cold
14800
298
373
0
1.59E+07
1.59E+07
4
He2cooling
water
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+07
11
u31
Hot
12396.4
400
394
3.83E+07
3.56E+07
-2.70E+06
12
Ccooling
water
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 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.
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
12064.7
5060.28
1668.93
98.6746
5060.28
0.21655
0
176.289
0
Impact
Generated
3324.36
3363.39
1668.02
-350.017
3363.39
0.21655
0
-4720.64
0
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.