1. 1
Conceptual Plant Design for the
Production of Dimethyl Carbonate
Ramiro Ramirez
Russell Wong
Group 20
June 3, 2015
Executive Summary
This report will provide a technical and profitability assessment associated with the construction
and operation of plant producing dimethyl carbonate (DMC) by the oxidative carbonylation of
methanol by oxygen and carbon monoxide at a rate of 150 million kilograms per year. This is an
eco-friendly alternative to the conventional method which utilizes phosgene, a highly toxic and
undesirable reagent. The biodegradability and low toxicity of this molecule combined with a
shift to more eco-friendly processes in the global chemical market is reflected by appreciable
growth and stability within the global DMC market.
The proposed design will utilize a single gas-liquid-solid slurry reactor along with proprietary
additives to provide the desired amount of DMC. Three distillation columns and a vapor
recovery system will be utilized to overcome the presence of azeotropes in the effluent and
deliver a pure material stream of DMC. Integration of all associated costs and economic factors
and on the basis of a two-year construction period and ten year operating time yields a Total
Capitalized Investment required equal to $54 million dollars. The Net Present Value of the
proposed project is equivalent to $54 million dollars with a relative annual growth of this value
normalized to the total capital investment equal to 5.7% each year at an expected industry tax
rate of 25%. This project will also provide a return on investment before taxes equal to 33% each
year and an estimate of the Internal Rate of Return (IRR) equal to 12.5%.
Further analysis providing comprehensive technical and economic considerations is provided.
Additional modeling of the system, plant wide control systems as well as sensibility and safety
analysis will reflect the feasibility of the proposed design. Given the conclusions of this base
case conceptual design, further research of alternative, more complex, separations systems is
recommended to decrease total capital and annual operating costs.

2. 2
Table of Contents
Executive Summary 1
Introduction and Market Overview 3
Production Chemistry 3
Plant Structure and Operating Conditions 4
Reactor System Modeling and Design Specifications 6
Separations System Modeling and Design Specifications 7
Economic Analysis 10
Discounted Cash Flow Analysis 11
Sensitivity Analysis 12
Risk and Safety Precautions 13
Process Control 13
Process Alternatives 15
Conclusion 16
References 17
Appendices
Appendix A: Production Chemistry and Design 17
Appendix B: Reactor Design at Various Operating Conditions 19
Appendix C: Separation System Design and Considerations 24
Appendix D: Economic Analysis 29
Appendix E: Sensitivity Analysis 40
Appendix F: Process Flow Diagrams 43
Appendix G: Matlab Code 47
Team Member Work Statements 66

3. 3
1. Introduction and Market Overview
Large volume demand for DMC is the result of its use as a methylating and carbonylating
reagent in the production of polycarbonates. These resulting products are low in toxicity, low
cost, and have desirable physical attributes making them amongst the most versatile and widely
used materials. Common uses for polycarbonates include vehicle parts, medical equipment,
housing materials, optical storage, and containers and packaging. DMC is also utilized as a
volatile organic compound exempt solvent, replacing conventionally used esters, ketones and
glycol ethers in formulation. Due to its high oxygen content, DMC is also utilized as a fuel
additive. Growth in this market is reflected by improved production methods and the eco-
friendly nature of the chemical, ensuring stability and profitability in the global DMC market.
The proposed project is a plant producing DMC through a more environmental friendly
approach, the oxidative carbonylation of methanol with carbon monoxide and oxygen, at a rate
of 150 million kilograms per year. Technical and economic analysis is conducted on the basis of
a 2-year construction period and 10-year operating time. For the purpose of economic analysis an
enterprise rate of 12.0%, a construction rate of 6.0%, a bond rate 4.0% and tax rates of 25%
(corporate tax rate) and 48% will be observed with relevant costs for involved utilities, chemicals
and equipment made available in Appendix C. Further detailed analysis and modeling of the
overall plant design will provide technical specifications as well as risk and profitability
assessment associated with the construction and operation of the proposed plant.
2. Production Chemistry
The observed reaction set for oxidative carbonylation of methanol by oxygen and carbon
monoxide can be shown by the following equations:
2 𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙 + 𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒 +
!
!
𝑂𝑥𝑦𝑔𝑒𝑛 → 𝐷𝑖𝑚𝑒𝑡ℎ𝑦𝑙 𝐶𝑎𝑟𝑏𝑜𝑛𝑎𝑡𝑒 + 𝑊𝑎𝑡𝑒𝑟 (1)
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒 +
1
2
𝑂𝑥𝑦𝑔𝑒𝑛 → 𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒 (2)
The reaction set occurs in the liquid phase over a heterogeneous catalyst comprising of
cuprous chloride (CuCl) as well as other proprietary additives. The formation of water in the
liquid phase is known to poison this catalyst, placing an operable limit of less than 15% water by

4. 4
mass in the liquid reaction compound. Oxygen concentration throughout the reaction is to be
maintained at a maximum of 4-mol% in vapor to ensure a non-explosive CO-O2 mixture. The
reaction set has an overall exothermic nature requiring heat exchange equipment in order to
maintain isothermal conditions. The activity of the various species in the reactor can be modeled
through a series of design equations and thermodynamic relations made available in Appendix A.
3. Plant Structure and Operating Conditions
Figure 1: Process mass and price flowsheet for the production of dimethyl carbonate.
E-1
Fresh Oxygen
6.4e3 kg/h
$20.6 MM/yr
Fresh Carbon Monoxide
11.3e3 kg/h
$17 MM/yr
Vapor Recycle
129e3 kg/h
Compressor
$2.20 MM
$0.24 MM/yrFresh Methanol
12.6e3 kg/h
$53.8 MM/yr
P-11
P-15
CSTR
$0.13 MM
$0.66 MM/yr
Cooler A
$0.05MM
Negligible/yr
Flash Drum A
$0.08 MM
Heat Exchanger A
$0.10 MM
Flash Drum B
$0.04 MM
P-25
Heater A
$0.28 MM
$0.63 MM/yr
Wastewater
3.6e3 kg/h
Negligible/yr
Heat Exchanger B
$0.21 MM
Heater B
$0.12 MM
$1.10 MM/yr Cooler C
$0.02 MM
$0.48 MM/yr
99.8 wt% Dimethyl Carbonate
18.0e3 kg/h
$136 M/yr
P-48
Cooler B
$0.08 MM
$0.35 MM/yr
P-57
Vapor Recovery System
$2.00 MM/yr
Column 1 (44 stages)
(MeOH+DMC/Water)
$3.88 MM
$4.50 MM/yr
Carbon Dioxide Purge
8.9e3 kg/h
DMC/MeOH Recycle
14.6e3 kg/h
Column 2 (21 stages)
(MeOH + DMC/DMC)
$2.67 MM
$8.95 MM/yr
Column 3 (30 Stages)
(MeOH + DMC/MeOH)
$5.87 MM
$7.87 MM/yr
Methanol Recycle
41.3e3 kg/h
Cooler D
Negligible Cap
$0.06 MM/yr

5. 5
A single gas-liquid-slurry reactor provides the best and most cost effective conditions for
this heterogeneous reaction. Pure methanol is provided into the reactor along with a vapor stream
of carbon monoxide and oxygen at the appropriate molar ratio. Unreacted reagents in the effluent
stream are to be completely recycled and combined with the fresh feed. Upon application of the
relevant design equations, it is determined the optimal operating conditions are achieved at high
pressures and temperatures, which for this process are limited at 40 bar and 130 °C (See
Appendix A). Reactor effluent is to be introduced to a separation system consisting of a vapor
and liquid recovery system, where the liquid recovery system comprises of three distillation
columns in order to provide a final 99.8-wt% DMC material stream as shown in Figure 1.
A relationship between the conversion of the limiting reagent, oxygen, and DMC
selectivity can be determined through analysis of the design equations and is shown by the
following figure:
Figure 2: Selectivity with respect to DMC versus the conversion of oxygen at molar ratios of
methanol to vapor (4-mol% O2, 96-mol% CO) in a slurry reactor at P = 40 bar and T = 130 °C.
Despite the decrease in selectivity with lower molar ratios of methanol, minimizing the
use of methanol by choosing a lower molar ratio within a reasonable rate will result in decreased
separation costs associated with high recycle rates of methanol. Observing the net present value
of the project under the same parameters as seen in Figure 3 can show this relationship.
0 0.2 0.4 0.6 0.8 1
0
0.2
0.4
0.6
0.8
1
1.2
1.4
1.6
1.8
2
Conversion (O2
)
Selectivity
MR = 1/3
MR = 1
MR = 3

6. 6
Figure 3: Net present value percent of the proposed design at various liquid/vapor mole ratios.
Based on the preliminary analysis of these parameters, a liquid/vapor molar ratio of 1/3
will be observed for this design. An increase in this value will compromise economic integrity
while choosing a smaller value will decrease selectivity, increasing the production of CO2 within
the plant and compromising environmental considerations. All subsequent analysis will be
conducted using this molar ratio, an oxygen conversion of 100%, a reactor temperature equal to
130°C and a reactor pressure of 40 bar.
4. Reactor System Modeling and Design Specifications
Matlab and Aspen HYSYS software has been utilized to demonstrate the performance of the
plant under the previously determined operating conditions. Various idealities are assumed in the
development of the Matlab model, which can account for the variance in numbers reported
between the software programs. Full HYSYS flowsheets and Matlab code are made available in
Appendices F and G.
A DMC material stream is required at a flowrate of 4.96 kg/s and purity of 99.8-wt% on
the basis that the plant operates continuously for 8400 hours a year. At the specified conditions,
the reactor specifications and inlet and outlet compositions of the reactor are as follows.
0 0.2 0.4 0.6 0.8 1
-25
-20
-15
-10
-5
0
5
10
15
20
25
Conversion (O2
)
NPV%
MR = 1/3
MR = 1
MR = 3

7. 7
Table 1: Values of species entering and leaving the reactor at the specified reactor conditions.
Reactor
Inlet
Reactor
Outlet
Matlab
HYSYS
Matlab
HYSYS
Species
Flowrate,
[kg/s]
Species
Flowrate,
[kg/s]
Fresh
Methanol
3.52
3.63
Methanol
10.7
15.1
Recycled
Methanol
10.7
14.99
Oxygen
0.00
0.00
Fresh
Oxygen
1.64
1.787
Carbon
Monoxide
33.1
30.7
Recycled
Oxygen
0.00
0.00
Carbon
Dioxide
2.09
2.61
Fresh
Carbon
Monoxide
2.86
3.12
Water
0.99
1.00
Recycled
Carbon
Monoxide
33.1
30.7
Dimethyl
Carbonate
4.96
6.40
Dimethyl
Carbonate
0.00
1.4
Total
Flowrate,
[kg/s]
51.9
55.7
Total
Flowrate,
[kg/s]
51.9
55.8
Reactor
Properties
Matlab
HYSYS
Volume,
[m^3]
3.0
3.6
Heat
Load,
[kW]
2.32E+04
Reactor measurements and conditions are consistent with this operation and provide the
desired amount of DMC. A higher rate of DMC production is required in the HYSYS model in
order to make up for the amount that is recycled back into the reactor due to non-idealities within
the separation system. There is adequate correspondence between the Matlab and HYSYS
models, which describe the system; both models reach approximately 100% oxygen conversion
and a DMC selectivity with respect to oxygen conversion equal to 1.07 and 0.98 respectively.
5. Separation System Modeling and Design Specifications
Given that the reaction requires large vapor and liquid streams, two separate recovery systems
are required to properly recycle all of the products. Multiple flash drums at low pressures and
temperatures attempt to fully separate all the carbon monoxide, carbon dioxide, and oxygen into
the vapor recovery system. In order create these conditions, a heat exchanger cools the vapor
outlet of the reactor and heats the initial vapor feed stream, and multiple throttle valves reduce
liquid pressures. Once the vapors are separated, they are sent to the vapor recovery system. The
vapor recovery system is treated as a utility, where the “black box” separation selectively
removes the undesired carbon dioxide from the other vapors at a single effective operating cost
(See Appendix C and D).

8. 8
The liquid recovery system is very complex due to the azeotropic nature between the
methanol and DMC and the methanol and water. Multiple separations were designed in Aspen
Plus, and a specific three-column separation setup is determined to be economically viable (See
Appendix C). These column specifications are initially determined by Aspen Plus before being
integrated into the HYSYS model. Reflux and reboil ratios are specified to be 1.5 times the
minimum ratios determined by Aspen Plus in order to account for fluctuations in the feed and
other non-theoretical conditions.
First, the reactor products are fed into a column at atmospheric pressure with a partial
condenser to remove the wastewater. The vapor distillate is brought to the vapor recovery
system, and the liquid distillate is pressurized, heated, and fed to the second column to be
separated. Specifically, this liquid column distillate is pressurized to be 20 bar, causing the feed
stream to be in a different distillation region, allowing for separation of 99.8-wt% DMC out of
the bottom of the column. The second column distillate cannot be fed back into the reactor feed
stream due to the inability of the first column to properly remove the water, so a third column is
required to create a pure methanol recycle feed stream. The DMC/methanol distillate from the
second column is cooled and fed into the third column, where a 99.9-mol% methanol stream is
attained from the bottom of the third column and recycled to the reactor feed. The third column’s
liquid distillate is recycled back into the second column after being pressurized and heated due to
its composition similarities to the second column feed stream, and the vapor distillate is sent to
the vapor recovery system.
Table 2: Column specifications required to accomplish separations are as follows:
Column
Specifications Column
1 Column
2 Column
3
Separation
Occuring AB
||
C AB
||
B AB
||
A
Number
of
Stages 44 21 30
Reflux
Ratio 4 3.5 3
Boilup
Ratio 30 105 10
Column
Pressure,
[atm] 1 20 1
Condensor
Temperature,
[C] 63.0 165 63.4
Reboiler
Temperature,
[C] 99.7 221 64.2
Condensor
Heat
Load,
[kW] 7.70E+04 1.40E+05 1.60E+05
Reboiler
Heat
Load,
[kW] 7.55E+04 1.60E+05 1.24E+05
Diameter,
[m] 5.7 6.20 3.80
Height,
[m] 14.6 7.44 10.2

9. 9

10. 10
6. Economic Analysis
The fixed capital investment is calculated by determining the initial costs of all major equipment
pieces utilized throughout the design. These installed costs are dominated by the three distillation
columns due to the difficulty of separation of the reactor effluent. Multiplication of capital costs
of all equipment by a factor of 2.28 will reflect the fixed capital costs of the design (See
Appendix D). Annual operating costs primarily consist of the purchase of raw materials as well
as the operation of the separation system.
Table 3: Capital and annual costs for proposed design.
Sale of dimethyl carbonate at the current market value of $0.90/kg generates annual
revenue of $136 million dollars resulting in a total profit before taxes equal to $18 million dollars
per year. In order to obtain a more comprehensive value of the total capital investment, the fixed
capital investment must be added to two months’ worth of raw materials and the start-up capital
equal to 10% of the fixed capital investment.
Table 4: Capital investment summary with the financing of the fixed capital
Equipment Cap Cost, [$MM] Op Cost, [$MM/yr]
CSTR 0.13 0.66
Vap Recovery System 0.00 2.00
2 Flash Drums 0.12 0.00
Column 1 3.88 4.50
Column 2 2.67 8.95
Column 3 5.87 7.87
Gas Compressor 2.20 0.24
Coolers (A,B,C,D) 0.16 0.89
Heaters (A,B) 0.40 1.77
Process Heat Exhangers (A,B) 0.30 0.00
Methanol 53.8
Carbon Monoxide 17.0
Oxygen 20.6
Total 15.7 118
Cost, [$MM]
Fixed Capital Investment 36.0
Working Capital 14.4
Start-Up Capital 3.60
Total Capital Investment 54.0

11. 11
7. Discounted Cash Flows and Detailed Economic Analysis
Discounted cash flows (DCF), a method described in Evaluating Plant Profitability in a Risk-
Return Context by Professor Mellichamp, determined the economic profitability of this design.
The economic parameters utilized to describe this design were conducted observing a tax rate of
25%, finance rate of 4%, enterprise rate of 12%, construction rate of 6%, 2 years of construction,
and 10 years of operation. Initial investment costs, the calculation of these parameters, and a
more detailed analysis is made available in Appendix D.
By calculating these parameters under various operating conditions, the profitability of
the plant can be optimized. Calculation of these parameters was conducted via a Matlab
conceptual design and validates the economic integrity of the established design parameters.
Figures 5 & 6: Optimization of Matlab model for NPV% (25% tax rate) and total capital
investment for various liquid/vapor molar ratios. Reactor operating at P = 40 bar and T = 130°C.
Use of the Aspen HYSYS model is used to provide the final cost estimates. The
simplified thermodynamics assumed in the programming of the Matlab model compared to the
more comprehensive fluid properties taken into account by the HYSYS model accounts for the
economic discrepancy between the models.
0 0.2 0.4 0.6 0.8 1
-25
-20
-15
-10
-5
0
5
10
15
20
25
Conversion (O2
)
NPV%
MR = 1/3
MR = 1
MR = 3
0 0.2 0.4 0.6 0.8 1
0
50
100
150
200
250
300
Conversion (O2
)
TotalCapitalInvestment,[$MM]
MR = 1/3
MR = 1
MR = 3

12. 12
Table 5: Economic parameters for Matlab and HYSYS models.
8. Sensitivity Analysis
A sensitivity analysis of the process’ NPV% reveals sensitivity to fluctuations in the value of the
raw materials, especially of methanol, and the product, dimethyl carbonate. This analysis reveals
that the process can tolerate an approximate decrease in DMC value of 10% or increase in
methanol value of 25% before reaching an NPV% of zero, or essentially the break-even point.
Figure 7: Sensitivity of project profitability to fluctuations in feed and product values.
There is a large risk in financing this process because of the large dependency of the
value of the final product, but this risk can be seem throughout all commodity chemical plants.
The decision to invest in this process is dependent on the stability of this value and future
projections of what this value may be. Plots for NPV% sensitivity versus enterprise rate, tax rate,
and finance rate can be found in Appendix E.
Economic Parameter Matlab HYSYS
Total Capital Investment, [$MM] 53.8 53.8
Profit Before Taxes 21.5 17.9
Return on Investment Before Taxes, [%/yr] 39.8 33.3
Net Present Value, [$MM at 25% Tax Rate] 70.3 54.2
Net Present Value Percent, [25% Tax Rate] 8.65 5.7
Net Present Value, [$MM at 45% Tax Rate] 51.8 40.5
Net Present Value Percent, [45% Tax Rate] 6.34 4.2
-40 -30 -20 -10 0 10 20 30 40
-10
-5
0
5
10
15
20
Percent Change in Price, [%]
NPV%
Dimethyl Carbonate
Methanol

13. 13
9. Risks and Safety Precautions
Possible economic risks that could result in the diminishing profitability of the plant include the
possible fluctuations in the market. Any lowering of the cost in DMC or increase in methanol or
carbon monoxide can dramatically damage the profitability of the plant due to the small price
difference of the two chemicals. Any increase in energy costs could possibly reduce the
profitability as well, considering that the separations are very energy intensive. In a failing
economy, the demand, and therefore price, of DMC would be reduced, leading to a negative
impact on the profitability of the plant. A different method of creating DMC, other than the
specified environmental-friendly alternative and separation system, could lead to a higher supply
of chemical, and once again, lower profits. These are not the only possible economic risks with
the plant, and many unknown variables could impact plant profitability.
The production of DMC from a non-phosgene route via green chemistry has a few safety
hazards. The largest risk comes from the explosive potential of the oxygen in the reaction, so the
concentration of oxygen does not exceed 4 mole% at any time in the plant. The exothermic
nature of the reaction at 130°C and 40 bar could pose a serious threat, so proper emergency
cooling and depressurization systems should be installed around the reactor. The chloride-based
copper catalyst poses a small risk because it can create a small amount of corrosive and toxic
hydrogen and/or methyl chloride, which is the reason why all components of the plant are
composed of stainless steel.
10. Process Control
The overall object of this process is to create 150 million kg/yr of 99.8-wt% DMC. In order to
reach this goal, the process begins with the ratio flow control on the feed amount of oxygen to
carbon monoxide for the vapor inlet, as well as the ratio flow control of vapor inlet to the liquid
methanol inlet. Further, feedback control on all of the pressures and temperatures throughout the
plant should be emphasized due to azeotropic nature of the reactor product mixture and the
separation system.

14. 14
Figure 8: Piping and instrumentation diagram for the proposed design, displaying control loops.

15. 15
Table 6: Controlled and manipulated variables for production of dimethyl carbonate.
Table 7: Design constraints for production of dimethyl carbonate.
11. Process Alternatives
Time was a major factor in the optimization of this plant. The process could have been run at
different conditions to change the conversion and risks, but this particular set of conditions were
chosen to ensure a low liquid water concentration in the reactor to ensure the water does not
poison the acidic catalyst. Although there is creation of 2% excess DMC, this was specifically
chosen to ensure enough DMC to sell. Operating at the maximum reactor temperature and
1 Molar ratio of carbon monoxide to oxygen in reactor vapor feed of 24
2 Molar ratio of vapor to liquid methanol in reactor feed of 3
3 Dimethyl carbonate flow rate of 150 million kg/yr
4 Reactor temperature of 130°C
5 Reactor pressure of 40 bar
6 Column 1 and 3 pressure of 1 bar
7 Column 2 pressure of 20 bar
Loop Control Type Controlled Manipulated Mechanism
F1 Feedback Carbon Monoxide Flow Rate Set Point/Fresh Carbon Monoxide V1
F2 Ratio Molar Ratio Feed of Oxygen Set Point/Fresh Oxygen V2
F3 Ratio Molar Ratio Feed of Methanol Set Point/Fresh Methanol V3
F4 Feedback Column 1 Vent Rate Vent Flow Rate V4
F5 Ratio Column 1 Reflux Ratio Steam and Flow Rates V5 + V6 + V7
F6 Feedback Column 1 Bottom Comp Steam and Flow Rates V8 + V9
F7 Ratio Column 2 Reflux Ratio Steam and Flow Rates V10 + V11 + V12
F8 Feedback Column 2 Bottom Comp Steam and Flow Rates V13 + V14
F9 Ratio Column 3 Reflux Ratio Steam and Flow Rates V15 + V16 + V17
F10 Feedback Column 3 Bottom Comp Steam and Flow Rates V18 + V19
P1 Feedback Reactor Input Pressure Power N/A
P2 Feedback First Flash Liquid Pressure Throttle Relief Valve V20
P3 Feedback Reactor Outlet Liquid Pressure Throttle Relief Valve V21
P4 Feedback Second Flash Liquid Pressure Throttle Relief Valve V22
P5 Feedback Product DMC Pressure Throttle Relief Valve V23
T1 Feedback Reactor Temp Coolant V24
T2 Feedback Flash Temp Coolant V25
T3 Feedback Column 1 Feed Temp Steam V26
T4 Feedback Column 2 Feed Temp Steam V27
T5 Feedback Column 3 Feed Temp Coolant V28
T6 Feedback Product DMC Temp Coolant V29

16. 16
pressure could be dangerous, but the excess of carbon monoxide should dilute the oxygen
enough to decrease the risk of explosions.
The vapor separation system could have been designed to a greater extent, compared to
the given “black box” single operating cost. The liquid separation system could have been
designed with different, more complex splits, such as side streams, to tackle the azeotropic
nature of the reactor products. These side streams could remove the need of a secondary or
tertiary column, which would greatly reduce the capital and operating costs of the plant. Also,
the methanol recycle stream comes out of the third, and last column, which is not recommended
due to the need to send all of this excess methanol throughout all of the separation system
compared to removing it earlier. The energy integration of the plant could be further explored to
optimize and reduce the heating and cooling costs of the plant. Currently, a large amount of
energy is being utilized to heat up and cool down the inputs and outputs of the secondary column
that separates out the pure DMC product, which could be reduced with the heat integration
optimization. These alternatives were not thoroughly investigated due to the lack of complex
separation knowledge and time.
12. Conclusion
The final proposed plant design would provide 150 million kilograms of 99.8% pure dimethyl
carbonate per year. The system relies on a single gas-liquid-solid slurry reactor operating with a
cuprous chloride catalyst as well as three distillation columns and a vapor recovery system. A
comprehensive economic analysis has yielded a Net Present Value for the project equal to $ 54
million dollars at an industry tax rate of 25% and has an annual growth on the Net Present Value
equal to 5.7%. The total capital investment required for this operation is equal to $54. Given the
results of this technical and economic analysis, further research and analysis of a separation
system specific to this process, such as the inclusion of entrainers, complex splits or column side
streams should be explored in order reduce capital and operating costs.

17. 17
References
[1] Douglas, J. M. Conceptual Design of Chemical Processes. N.p.: McGraw-Hill, 1988. Print.
[2] Mellichamp, D. A. Evaluating Plant Profitability in a Risk-Return Context. N.p: Department
of Chemical Engineering, UCSB, 2012. Print.
[3] Doherty, Michael F., and Michael F. Malone. Conceptual Design of Distillation Systems.
Boston: McGraw-Hill, 2001. Print.
Appendix
A:
Reaction
Set
Chemistry
and
Design
Equations
-­‐ Observed
reaction
set
for
the
formation
of
Dimethyl
Carbonate
2 𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙 + 𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒 +
1
2
𝑂𝑥𝑦𝑔𝑒𝑛 → 𝐷𝑖𝑚𝑒𝑡ℎ𝑦𝑙 𝐶𝑎𝑟𝑏𝑜𝑛𝑎𝑡𝑒 + 𝑊𝑎𝑡𝑒𝑟 𝐴. 1
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒 +
1
2
𝑂𝑥𝑦𝑔𝑒𝑛 → 𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒 (𝐴. 2)
-­‐ Respective
Rates
of
Reaction,
[mol/L*s]
𝑟! = 𝑘! 𝐶!"
!
𝐶!!
!
!
(𝐴. 3)
𝑟! = 𝑘! 𝐶!!
!
!
(𝐴. 4)
-­‐ Corresponding
‘k’
values:
𝑘! = 1.4 𝑥 10!!
exp −
24000
𝑐𝑎𝑙
𝑔𝑚𝑜𝑙𝑒
𝑅𝑇
(𝐴. 5)
𝑘! = 5.6 𝑥 10!"
exp −
22700
𝑐𝑎𝑙
𝑔𝑚𝑜𝑙𝑒
𝑅𝑇
(𝐴. 6)
Note:
R
=
1.987
cal/mol*K
,
‘T’
is
measured
in
Kelvin,
and
‘Ci’
is
the
concentration
of
the
species
present
in
the
liquid
phase.

18. 18
Heats
of
Reaction
𝑅𝑥𝑛 1: − 6 𝑥 10!
𝑘𝐽
𝑘𝑔𝑚𝑜𝑙𝑒 (𝐴. 7)
𝑅𝑥𝑛 2: − 5.7 𝑥 10!
𝑘𝐽
𝑘𝑔𝑚𝑜𝑙𝑒 (𝐴. 8)
Concentration
of
gaseous
compounds
in
the
liquid
phase
can
be
determined
through
Henry’s
Law:
𝑝! = 𝐾! 𝑥! (𝐴. 9)
Note:
‘pi’
represents
the
partial
pressure
in
‘bar’,
‘xi’
is
the
fraction
present
in
the
liquid
phase
and
‘KH’
is
the
Henrys
Constant:
Table
A.1:
Henry’s
constant
for
reaction
species.
Gas
Henry’s
Constant,
[bar]
Oxygen
3179
Carbon
Monoxide
3107
Carbon
Dioxide
158
Total
molar
flowrates
for
all
species
as
determined
by
Level
2
mole
balances:
𝐹!" −
𝐹!"
2
− 𝑃!"!
= 0 (𝐴. 10)
𝐹!!
−
𝐹!"
4
−
𝑃!"!
2
= 0 (𝐴. 11)
𝑃!"# =
𝐹!"
2
(𝐴. 12)
𝑃!!! =
𝐹!"
2
(𝐴. 13)
CSTR
Design
Equation:
𝑐!! − 𝑐! + 𝜏𝑟! = 0 𝑖 = 1,2, … 𝑛 (𝐴. 14)

19. 19
Appendix
B:
Reactor Design at Various Operating Conditions
Figure B.1: Selectivity with respect to DMC versus oxygen conversion at various molar ratios of
liquid to vapor feed into a CSTR operating at T = 130 °C and P = 40 bar.
Figure B.2: Volume of CSTR reactor required to produce desired DMC operating with a
liquid/vapor ration of 1/3 and T = 130 °C and P = 40 bar.
0 0.2 0.4 0.6 0.8 1
0
0.2
0.4
0.6
0.8
1
1.2
1.4
1.6
1.8
2
Conversion (O2
)
Selectivity
MR = 1/3
MR = 1
MR = 3
0 0.2 0.4 0.6 0.8 1
2
2.2
2.4
2.6
2.8
3
3.2
3.4
3.6
3.8
Conversion (O2
)
ReactorVolume,[m3
]

20. 20
Figure B.3: Recycle rate of reactants required to achieve DMC production.
Figure B4: Fresh feed rate of reactants required to achieve desired DMC production.
0 0.2 0.4 0.6 0.8 1
0
50
100
150
200
250
Conversion, (O2
)
RecycleFlowrate,[kg/s]
Methanol
Oxygen
Carbon Monoxide
0 0.2 0.4 0.6 0.8 1
0
50
100
150
200
250
300
Conversion (O2
)
TotalCapitalInvestment,[$MM]
MR = 1/3
MR = 1
MR = 3

21. 21
Figure B.5: Total mass flowrate required to enter the reactor in order to achieve required DMC.
Figure B.6: Total capital investment required to develop proposed design at various liquid/vapor
ratios. Reactor operating at T = 130C and P = 40 bar.
0 0.2 0.4 0.6 0.8 1
50
100
150
200
250
300
350
400
Conversion (O2
)
FlowrateIntoReactor,[kg/s]
0 0.2 0.4 0.6 0.8 1
1
1.5
2
2.5
3
3.5
4
4.5
5
Conversion (O2
)
FreshFeedRate,[kg/s]
Methanol
Oxygen
Carbon Monoxide

22. 22
Figure B.7: Net present value of the proposed design operating under various liquid/vapor molar
ratios.
Figure B.8: Net present value percent for the proposed design operating under various
liquid/vapor molar ratios.
0 0.2 0.4 0.6 0.8 1
-25
-20
-15
-10
-5
0
5
10
15
20
25
Conversion (O2
)
NPV%
MR = 1/3
MR = 1
MR = 3
0 0.2 0.4 0.6 0.8 1
-600
-500
-400
-300
-200
-100
0
100
200
300
400
Oxy Conversion
NPVProj
,[$MM]
MR = 1/3
MR = 1
MR = 3

23. 23
Figure B.9: Return on investment before taxes for the proposed design operating under a
liquid/vapor molar ratio of 1/3.
Figure B.10: Measure of the water present in the liquid phase in the proposed reactor.
0 0.2 0.4 0.6 0.8 1
-80
-60
-40
-20
0
20
40
60
80
Conversion (O2
)
ReturnonInvestmentBeforeTaxes,[%/yr]
0 0.2 0.4 0.6 0.8 1
0
1
2
3
4
5
6
7
Conversion (O2
)
MassPercentWaterinReactor

24. 24
Appendix
C:
Separation
System
Design
and
Considerations
Vapor
Recovery
System
The
following
method
represents
a
very
rough
estimation
of
the
associated
costs
of
a
vapor
recovery
system
capable
of
removing
CO2
from
reactor
vapor
effluent.
Model
assumes
idealities
in
the
system
and
uniform
temperature
and
pressure
throughout
the
process.
Total
Capital
and
Operating
Cost
can
be
lumped
into
a
single
operating
cost
as
determined
by
the
following
equation:
𝐶 = 𝜆 𝜖 𝑊!"# (𝐶. 1)
Note:
‘C’
is
the
total
annual
operating
cost,
‘λ’
is
a
scalar
factor
with
a
value
of
6,
and
‘WMin’
is
the
minimum
amount
of
work
required
for
separation
to
occur.
The
value
of
‘WMin’
can
be
determined
by
the
following
expression:
𝑊!"#
𝐹𝑅𝑇!"##
= 1 − 𝑧!"!
ln
1
1 − 𝑧!"!
𝜉
+ 𝑧!"!
1 − 𝜉 ln
1 − 𝜉
1 − 𝑧!"!
𝜉
− 𝜉𝑧!"!
ln 𝑧!"!
(𝐶. 2)
Note:
𝑓 =
𝑌
𝐹
(𝐶. 3)
𝜉 =
𝑓
𝑧!"!
(𝐶. 4)

25. 25
Liquid Recovery System
For a tertiary mixture, Aspen Plus is utilized to model the columns necessary for the liquid
separations, and Matlab with multiple correlations of the HYSYS specifications were utilized to
size the columns.
Corresponding heat loads in the reactor and condenser may be determined by the following
equations:
𝑄! = 𝜆! 𝑉! (𝐶. 5)
𝑄! = 𝜆! 𝑉! (𝐶. 6)
The cross sectional area of the column may be determined by the following equation:
𝐴 =
𝑀!
𝑝! 𝑝!
1
𝜙!"##$
𝐴
𝐴!
𝑉 (𝐶. 7)
Where ‘A’ is the cross sectional area in square meters, ‘Mv’ is the molar weight of the vapor,
‘𝜙!"##$’ is the desired fraction of flooding velocity, ‘(A/An)’ is the fraction available for flow
and ‘V’ represents the vapor rate.
The corresponding height of the column can then be determined using the following equation:
𝐻 = 𝐻!"# + 𝐻! 𝑁 (𝐶. 8)
Where ‘Hmin’ is equal to three times the tray spacing, ‘Ht’ and added to the over height of the
column, ‘H’.

26. 26
Figure C.1: Aspen Plus diagram for the first distillation column in the separation system, which
separates out the water from the methanol and dimethyl carbonate.
Table C.1: Aspen Plus composition results for the first column
Table C.2: Aspen Plus first column specifications.
T e rna ry
M a p
(M o le
B a s is )
WATER
(99.649 C)
MEOH
(64.201 C)
DMC
(89.806 C)
0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
0.10.20.30.40.50.60.70.80.9
DIST1
BOT1
Feed Stream
Column 1 (AB/C) Methanol Dimethyl Carbonate Water
Input Composition 0.765 0.117 0.118
Distillate Composition 0.867 0.133 Negligible
Bottom Composition Negligible Negligible 1.000
Number of Stages 41.6
Feed Stage 19.2
Pressure (bar) 1.00
Vapor/Liquid Flowrate 3.55
Reflux Ratio 3.00
Reboil Ratio 30.0

27. 27
Figure C.2: Aspen Plus diagram for the second distillation column in the separation system,
which separates out the 99.8-wt% pure dimethyl carbonate from the methanol and dimethyl
carbonate.
Table C.3: Aspen Plus composition results for the second column.
Table C.4: Aspen Plus second column specifications.
T e rna ry
M a p
(M o le
B a s is )
WATER
(212.476 C)
MEOH
(165.921 C)
DMC
(222.297 C)
0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
0.10.20.30.40.50.60.70.80.9
DIST2
BOT2
Feed Stream
Column 2 (AB/B) Methanol Dimethyl Carbonate Water
Input Composition 0.867 0.133 Negligible
Distillate Composition 0.910 0.090 Negligible
Bottom Composition Negligible 0.998 Negligible
Number of Stages 19.8
Feed Stage 10.4
Pressure (bar) 20.0
Vapor/Liquid Flowrate 3.42
Reflux Ratio 2.00
Reboil Ratio 72.2

28. 28
Figure C.3: Aspen Plus diagram for the third distillation column in the separation system, which
separates out the 99.8% pure methanol from the methanol and dimethyl carbonate.
Table C.5: Aspen Plus composition results for the third column.
Table C.6: Aspen Plus third column specifications.
Appendix D: Economic Analysis
T e rna ry
M a p
(M o le
B a s is )
WATER
(99.649 C)
MEOH
(64.201 C)
DMC
(89.806 C)
0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9
0.10.20.30.40.50.60.70.80.9
DIST3
BOT3
Feed Stream
Column 3 (AB/A) Methanol Dimethyl Carbonate Water
Input Composition 0.910 0.090 Negligible
Distillate Composition 0.897 0.103 Negligible
Bottom Composition 0.998 0.002 Negligible
Number of Stages 29.6
Feed Stage 14.7
Pressure (bar) 1.00
Vapor/Liquid Flowrate 4.49
Reflux Ratio 2.00
Reboil Ratio 8.21

29. 29
Raw Material Values
Typical Steam Prices
Initial Equipment Costing
Relations found in Appendix E of Conceptual Design of Chemical Processes by James Douglas
were used to find the installed equipment costs for every piece of equipment.
Reactor
The shell and tube reactor was split into two different parts, a heat exchanger and pressure
vessel. In particular, this design can be broken down into the heat exchanger and 500 different
pressure vessels, which are each of the tubes.
Substance Cost/Unit
Dimethyl Carbonate $0.90/kg
Methanol $0.49/kg
Carbon Monoxide $0.18/kg
Oxygen $0.38/kg
Cooling Water $0.08/1000 kg
Wastewater $0.06/1000 kg
Heating Fuel $3.00/MMBtu
Pressure (psia) Temperature (°C) Cost ($/1000 kg) ΔH (kJ/kg)
30 121 2.38 2213
50 138 3.17 2159
100 165 4.25 2067
200 194 5.32 1960
500 242 6.74 1755
750 266 7.37 1634

30. 30
Heat Exchanger
The heat exchanger’s installed cost can be modeled as:
Installed Cost, $ =
𝑀&S
280
101.3𝐴!.!"
2.29 + 𝐹! (𝐷. 1)
where the M&S is the Marshall and Swift index, or 1600 in modern day, A is the heat exchanger
area, and 𝐹! is the correction factor, defined as:
𝐹! = 𝐹! + 𝐹! 𝐹! (𝐷. 2)
where 𝐹! is the factor that accounts for the design type, 𝐹! is the factor that accounts for
pressure, and 𝐹! is the factor that accounts for the shell and tube material. Here, the 𝐹! is 0.80
because of the fixed-tube sheet design, 𝐹! is 0.00 because of the pressure only being 2 bar or
approximately 30 psi, and 𝐹! is equal to 1.00 for carbon steel on carbon steel, resulting in a 𝐹! of
0.80. The areas, A, for the HYSYS heat exchangers were found through a correlation between
typical heat transfer coefficients, U, heat transferred, Q, and average change in temperature,
ΔTavg. This assumes cooling with cooling water with ΔT=20°C, and heating with steam.
𝑄 = 𝑈 ∗ 𝐴 ∗ ∆𝑇!"# (𝐷. 3)
Pressure Vessel
The pressure vessel’s installed cost can be modeled as:
Installed Cost, $ =
𝑀&S
280
101.9𝐷!.!""
𝐻!.!"
2.18 + 𝐹! (𝐷. 4)
where the M&S is the Marshall and Swift index, or 1600 in modern day, D is the diameter of the
pressure vessel, H is the height or length of the pressure vessel, and 𝐹! is the correction factor,
defined as:
𝐹! = 𝐹! ∗ 𝐹! (𝐷. 5)
where 𝐹! is the factor that accounts for pressure and 𝐹! is the factor that accounts for material.
Specifically, 𝐹! is 1.00 for pressures up to 50 psi and 𝐹! is 1.00 for carbon steel.

31. 31
Gas Compressor
The gas compressor cost can be modeled as:
Installed Cost, $ =
𝑀&S
280
517.5 𝑏ℎ𝑝 !.!"
(2.11 + 𝐹!) (𝐷. 6)
Where ‘bhp’ is the brake horsepower and ‘FC’ is a correction factor which for this process have
values of 800 and 1, respectively.
Pump
Pumps are assumed to have negligible costs due to its several orders of magnitude cheaper than
all of the other equipment.
Separation System (Distillation Columns)
The distillation column initial cost can be modeled as the costs of two heat exchangers, which
are the reboiler and condenser, added the cost of the column shell and trays.
The column shell can be modeled as:
𝐶! = 𝐶!,!
𝑑
𝑑!
𝐻
𝐻!
∝!
(𝐷. 7)
where 𝑑 is the column diameter, 𝐻 is the column height, 𝑑! = 1 and 𝐻! = 6.1 for a calculation
in meters, ∝!= 0.82 for a shell constant, and 𝐶!,! is calculated as:
𝐶!,! =
𝑀&S
280
𝐹! 𝐹! − 1 + 𝐹! 𝐹! 𝑐!,! (𝐷. 8)
where 𝑀&𝑆 = 1600 for the Marshall and Swift index in modern day, 𝐹! = 1 for a carbon steel
tray material, 𝐹! = 1 for an operating pressure less than 4.5 bar, 𝐹! = 1.38 for indirect cost
factor, 𝐹! = 3.00 for direct cost factor, and 𝑐!,! = 5000 for the shell cost.
For columns operating at pressures greater than 4.5 bar, 𝐹! is determined by the following
relation:
𝐹! = 1 + 𝑡 1 + 𝑒
!!
! 𝑤ℎ𝑒𝑟𝑒 𝑡 =
𝑃 ∗ 𝑃!
𝑃!
(𝐷. 9)

32. 32
The tray cost can be modeled as:
𝐶! = 𝐶!,!
𝑑
𝑑!
∝!
𝐻
𝐻!
(𝐷. 10)
where ∝!= 1.8 for a tray constant, and 𝐶!,! is calculated as:
𝐶!,! =
𝑀&S
280
𝐹! + 𝐹! + 𝐹! 𝑐!,! (𝐷. 11)
where 𝐹! = 0 for sieve tray types. The total installed capital cost of the column is then:
𝐶!"# = 𝐶! + 𝐶! (𝐷. 12)
Vapor Recovery System
Total Capital and Operating Cost can be lumped into a single operating cost as determined by the
following equation:
𝐶 = 𝜆 𝜖 𝑊!"# (𝐷. 13)
Note: ‘C’ is the total annual operating cost, ‘λ’ is a scalar factor with a value of 6, and ‘WMin’ is
the minimum amount of work required for separation to occur as outlined in Appendix C.
Yearly Revenues and Costs
Revenues
Yearly revenues, or R, generated can be found by the sale value of all of the products.
𝑅 = 𝑃!"#$ ∗ 𝑣𝑎𝑙𝑢𝑒!"#$
!"#$%&'(
(𝐷. 14)
where the products is dimethyl carbonate, generating an annual revenue of $136 million dollars.
Costs

33. 33
Yearly costs, or C, can be calculated by:
𝐶 = 𝐶!"#$#"#%& + 𝐶!"#$%&"' + 𝐶!"#$%$&'()! (𝐷. 15)
Table D.1 Yearly Costs.
Profit Before Taxes
Profit Before Taxes = 𝑃𝐵𝑇 = 𝑅𝑒𝑣𝑒𝑛𝑢𝑒 − 𝐶𝑜𝑠𝑡 (𝐷. 16)
Fixed Capital
Fixed capital was calculated using the factored estimates approach, where
Fixed Capital = 𝐹𝐶 = 2.28 ∗ 𝐼𝑆𝐵𝐿 (𝐷. 17)
where ISBL is the sum of the installed costs of all the equipment, and the 2.28 comes from an
estimate of direct costs being the sum of the installed costs and offsite costs (~40% of the
installed costs), the indirect costs being ~30% of the direct costs, and there being a 25% c
ontingency on the direct costs.
Working Capital
Equipment Cap Cost, [$MM] Op Cost, [$MM/yr]
CSTR 0.13 0.66
Vap Recovery System 0.00 2.00
2 Flash Drums 0.12 0.00
Column 1 3.88 4.50
Column 2 2.67 8.95
Column 3 5.87 7.87
Gas Compressor 2.20 0.24
Coolers (A,B,C,D) 0.16 0.89
Heaters (A,B) 0.40 1.77
Process Heat Exhangers (A,B) 0.30 0.00
Methanol 53.8
Carbon Monoxide 17.0
Oxygen 20.6
Total 15.7 118

34. 34
For the purpose of this design, working capital was assumed to be worth approximately the cost
of two months of raw materials, or two months’ worth of ethyl-benzene.
Working Capital = 𝑊𝐶 = 1400 ∗ 𝐶!"#$%&"' (𝐷. 18)
Start-up Capital
For the purpose of this design, start-up capital was assumed to be worth approximately 10% of
the fixed capital.
Start Up Capital = 𝑆𝑈 = 0.1 ∗ 𝐹𝐶 (𝐷. 19)
Total Capital Investment
The total capital investment can be calculated as:
Total Capital Investment = 𝑇𝐶𝐼 = 𝑎! ∗ 𝐹𝐶 1 + 𝐶𝑅 !
!
+ 𝑆𝑈 + 𝑊𝐶 (𝐷. 20)
where j is construction years relative to plant start up, where j=0 at the finishing year of
construction, CR is the construction rate, and 𝑎! is the fractional allocation of the fixed capital
during construction years.
Profitability Measurement
The plant profitability can be determined by metrics of profit before taxes (PBT), net present
value (NPV), net present value at plant start-up (NPV0), discounted net present value (NPVproj),
and normalized net present value (NPV%). Variables, such as tax rate (TR), enterprise rate (ER),
total capital invested (TCI), finance rate (FR), greatly affect these metrics.
𝑁𝑃𝑉! = 1 − TR ∗ 𝑃𝐵𝑇 ∗ 𝑏! 1 + ER !!
!
!!!
− 1 − TR FR ∗ 𝑇𝐶𝐼 ∗ 𝑏! 1 + ER !!
!
!!!
+ 0.1
∗ TR FC + 1.1 𝑏! 1 + ER !!
!
!!!
+ 1 − TR 𝑊𝐶 + 𝑆𝑈 − 𝑇𝐶𝐼 1 + ER !!
(𝐷. 21)
where 𝑏! represents the fraction of profit received each year and m is the lifetime of plant
operation.

35. 35
𝑁𝑃𝑉!"#$ =
𝑁𝑃𝑉!
1 + 𝐸𝑅 !
(𝐷. 22)
where n is the number of construction years.
𝑁𝑃𝑉% =
𝑁𝑃𝑉!"#$
𝑚 + 𝑛 ∗ 𝑇𝐶𝐼
(𝐷. 23)
The return on investment before taxes is given by:
Return on Investment Before Taxes = 𝑅𝑂𝐼!" =
𝑃𝐵𝑇
𝐹𝐶 + 𝑊𝐶 + 𝑆𝑈
(𝐷. 24)
In particular, this design has a 𝑁𝑃𝑉% = 5.7% and 𝑅𝑂𝐼!" = 33.3%.
Table D.3 Economic parameters for Matlab and HYSYS designs
Internal Rate of Return
A root finding method on enterprise rate of the project calculates the internal rate of return when
there is no external financing. Excel spreadsheets allow for this value to be calculated. In
particular, the IRR is 9.8% at 25% tax rate, where it is 1.2% at 48% tax rate
Economic Parameter Matlab HYSYS
Total Capital Investment, [$MM] 53.8 53.8
Profit Before Taxes 21.5 17.9
Return on Investment Before Taxes, [%/yr] 39.8 33.3
Net Present Value, [$MM at 25% Tax Rate] 70.3 54.2
Net Present Value Percent, [25% Tax Rate] 8.65 5.7
Net Present Value, [$MM at 45% Tax Rate] 51.8 40.5
Net Present Value Percent, [45% Tax Rate] 6.34 4.2

36. 36
Figure D.1: Finance Sheet at tax rate 25%
All dollar amounts in table represent millions of dollars. Construction & operations period are in years.
Profit_BT = 18.0
Construction Rate 6.0% Tax Rate 25% Nconstruction 2
Finance Rate 4.0% Enterprise Rate 12.0% Noperations 10 Yields Tot.Cap.Inv. TI=FC+WC+SU
ROI_BT = 28.8%
Fixed Capital 36.0 a-3 0.00
alpha_Working Capital 20% a-2 0.00 b_1 0.80 TI=FC+WC+SU
alpha_Start-Up Capital 10% a-1 0.50 b_2 0.90
alpha_Salvage Value 3% a0 0.50 b_3 0.95
Capital In (+) Discount Discounted
Year DesignConstruction Period or Out (-) Factors Cash Flows
-3 Fixed Capital in Y-3 0.0 1.191 0.0
-2 Fixed Capital in Y-2 0.0 1.124 0.0
-1 Fixed Capital in Y-1 -18.0 1.060 -19.1
0 Fixed Capital in Y0 -18.0 1.000 -18.0
0 Working Capital -22.9 1.000 -22.9
0 Start-Up Capital -3.6 1.000 -3.6
0 Total of Capital Outlays -63.6
(=Sum of Constr. DCFs)
0 Total Capital Investment 63.6
(=Proceeds of Bond Issue)
Profit Bond Depreciation Profit Cash
Operations Period Before Taxes Financing Allowed After Taxes Flows
1 14.4 -2.5 -4.0 5.9 9.9 0.893 8.8
2 16.2 -2.5 -4.0 7.3 11.2 0.797 9.0
3 17.1 -2.5 -4.0 7.9 11.9 0.712 8.5
4 18.0 -2.5 -4.0 8.6 12.6 0.636 8.0
5 18.0 -2.5 -4.0 8.6 12.6 0.567 7.1
6 18.0 -2.5 -4.0 8.6 12.6 0.507 6.4
7 18.0 -2.5 -4.0 8.6 12.6 0.452 5.7
8 18.0 -2.5 -4.0 8.6 12.6 0.404 5.1
9 18.0 -2.5 -4.0 8.6 12.6 0.361 4.5
10 18.0 -2.5 -4.0 8.6 12.6 0.322 4.1
10 Working Capital 22.9 22.9 0.322 7.4
10 Salvage Value 1.1 0.6 0.6 0.322 0.2
10 Pay-Off TCI -63.6 -63.6 0.322 -20.5
WC & SV Total Profit Bond Interest Total Total Profit Total NPV(0) NPV-proj [NPV(0)
All Figures Represent Recovery Before Taxes Payments Depreciation After Taxes Cash Flow Discounted to EOY(-x)]
PV of Operations==> 7.6 96.4 -14.4 22.4 44.9 54.2 54.2 43.2
Bond Total Capital
Repayment Recovery
-20.5 29.7 Total Cash
Net Present Flow as Sum NPV(0) Avg. NPV_proj Avg.
Value of Bonds of NPVs Over y Years Over z Years
28.7 54.2 8.5% 5.7%
x = Nconstruction
y = Nooperations
z = Nconstruction + Noperations
NPV Increase per Year
normalized/annualized
CR and FR can be chosen independently. WC and SV are converted from inventory to Profit_BT in Year 10.
Table 3. Risk-Based Profitability Analysis: Net Present Value Via Discounted Cash Flow
Establishes the ROI_BT Based on TI (also TCI) and the Annual % Increase of NPV (Normalized by TCI)
Fixed Capital and Profit_BT are the two independent variables.
using Capitalization =

37. 37
Figure D.2: Finance sheet at tax rate 48%.
All dollar amounts in table represent millions of dollars. Construction & operations period are in years.
Profit_BT = 18.0
Construction Rate 6.0% Tax Rate 48% Nconstruction 2
Finance Rate 4.0% Enterprise Rate 12.0% Noperations 10 Yields Tot.Cap.Inv. TI=FC+WC+SU
ROI_BT = 28.8%
Fixed Capital 36.0 a-3 0.00
alpha_Working Capital 20% a-2 0.00 b_1 0.80 TI=FC+WC+SU
alpha_Start-Up Capital 10% a-1 0.50 b_2 0.90
alpha_Salvage Value 3% a0 0.50 b_3 0.95
Capital In (+) Discount Discounted
Year DesignConstruction Period or Out (-) Factors Cash Flows
-3 Fixed Capital in Y-3 0.0 1.191 0.0
-2 Fixed Capital in Y-2 0.0 1.124 0.0
-1 Fixed Capital in Y-1 -18.0 1.060 -19.1
0 Fixed Capital in Y0 -18.0 1.000 -18.0
0 Working Capital -22.9 1.000 -22.9
0 Start-Up Capital -3.6 1.000 -3.6
0 Total of Capital Outlays -63.6
(=Sum of Constr. DCFs)
0 Total Capital Investment 63.6
(=Proceeds of Bond Issue)
Profit Bond Depreciation Profit Cash
Operations Period Before Taxes Financing Allowed After Taxes Flows
1 14.4 -2.5 -4.0 4.1 8.1 0.893 7.2
2 16.2 -2.5 -4.0 5.0 9.0 0.797 7.2
3 17.1 -2.5 -4.0 5.5 9.5 0.712 6.7
4 18.0 -2.5 -4.0 6.0 9.9 0.636 6.3
5 18.0 -2.5 -4.0 6.0 9.9 0.567 5.6
6 18.0 -2.5 -4.0 6.0 9.9 0.507 5.0
7 18.0 -2.5 -4.0 6.0 9.9 0.452 4.5
8 18.0 -2.5 -4.0 6.0 9.9 0.404 4.0
9 18.0 -2.5 -4.0 6.0 9.9 0.361 3.6
10 18.0 -2.5 -4.0 6.0 9.9 0.322 3.2
10 Working Capital 22.9 22.9 0.322 7.4
10 Salvage Value 1.1 0.6 0.6 0.322 0.2
10 Pay-Off TCI -63.6 -63.6 0.322 -20.5
WC & SV Total Profit Bond Interest Total Total Profit Total NPV(0) NPV-proj [NPV(0)
All Figures Represent Recovery Before Taxes Payments Depreciation After Taxes Cash Flow Discounted to EOY(-x)]
PV of Operations==> 7.6 96.4 -14.4 22.4 31.2 40.5 40.5 32.3
Bond Total Capital
Repayment Recovery
-20.5 29.7 Total Cash
Net Present Flow as Sum NPV(0) Avg. NPV_proj Avg.
Value of Bonds of NPVs Over y Years Over z Years
28.7 40.5 6.4% 4.2%
x = Nconstruction
y = Nooperations
z = Nconstruction + Noperations
NPV Increase per Year
normalized/annualized
CR and FR can be chosen independently. WC and SV are converted from inventory to Profit_BT in Year 10.
Table 3. Risk-Based Profitability Analysis: Net Present Value Via Discounted Cash Flow
Establishes the ROI_BT Based on TI (also TCI) and the Annual % Increase of NPV (Normalized by TCI)
Fixed Capital and Profit_BT are the two independent variables.
using Capitalization =

38. 38
Figure D.3: IRR calculation sheet at tax rate 25%
All dollar amounts in table represent millions of dollars. Construction & operations period are in years.
Profit_BT = 18.0
Construction Rate 6.0% Tax Rate 25% Nconstruction 2
Finance Rate 0.0% Enterprise Rate 12.5% Noperations 10 Yields Tot.Cap.Inv. TI=FC+WC+SU
ROI_BT = 28.8%
Fixed Capital 36.0 a-3 0.00
alpha_Working Capital 20% a-2 0.00 b_1 0.80 TI=FC+WC+SU
alpha_Start-Up Capital 10% a-1 0.50 b_2 0.90
alpha_Salvage Value 3% a0 0.50 b_3 0.95
Capital In (+) Discount Discounted
Year DesignConstruction Period or Out (-) Factors Cash Flows
-3 Fixed Capital in Y-3 0.0 1.191 0.0
-2 Fixed Capital in Y-2 0.0 1.124 0.0
-1 Fixed Capital in Y-1 -18.0 1.060 -19.1
0 Fixed Capital in Y0 -18.0 1.000 -18.0
0 Working Capital -22.9 1.000 -22.9
0 Start-Up Capital -3.6 1.000 -3.6
0 Total of Capital Outlays -63.6
(=Sum of Constr. DCFs)
0 Total Capital Investment 63.6
(=Proceeds of Bond Issue)
Profit Bond Depreciation Profit Cash
Operations Period Before Taxes Financing Allowed After Taxes Flows
1 14.4 0.0 -4.0 7.8 11.8 0.889 10.5
2 16.2 0.0 -4.0 9.2 13.1 0.790 10.4
3 17.1 0.0 -4.0 9.9 13.8 0.703 9.7
4 18.0 0.0 -4.0 10.5 14.5 0.625 9.1
5 18.0 0.0 -4.0 10.5 14.5 0.556 8.1
6 18.0 0.0 -4.0 10.5 14.5 0.494 7.2
7 18.0 0.0 -4.0 10.5 14.5 0.439 6.4
8 18.0 0.0 -4.0 10.5 14.5 0.390 5.7
9 18.0 0.0 -4.0 10.5 14.5 0.347 5.0
10 18.0 0.0 -4.0 10.5 14.5 0.309 4.5
10 Working Capital 22.9 22.9 0.309 7.1
10 Salvage Value 1.1 0.6 0.6 0.309 0.2
10 Pay-Off TCI 0.0 -63.6 0.309 -19.6
WC & SV Total Profit Bond Interest Total Total Profit Total NPV(0) NPV-proj [NPV(0)
All Figures Represent Recovery Before Taxes Payments Depreciation After Taxes Cash Flow Discounted to EOY(-x)]
PV of Operations==> 7.2 94.5 0.0 21.9 54.6 64.0 0.4 0.3
Bond Total Capital
Repayment Recovery
-19.6 29.0 Total Cash
IRR Calculation Net Present Flow as Sum NPV(0) Avg. NPV_proj Avg.
-15% Value of Bonds of NPVs Over y Years Over z Years
with the IRR function 44.0 64.0 0.1% 0.0%
x = Nconstruction
y = Nooperations
z = Nconstruction + Noperations
NPV Increase per Year
normalized/annualized
CR and FR can be chosen independently. WC and SV are converted from inventory to Profit_BT in Year 10.
Table 3. Risk-Based Profitability Analysis: Net Present Value Via Discounted Cash Flow
Establishes the ROI_BT Based on TI (also TCI) and the Annual % Increase of NPV (Normalized by TCI)
Fixed Capital and Profit_BT are the two independent variables.
using Capitalization =

39. 39
Figure D.4: IRR calculation sheet at tax rate 48%
All dollar amounts in table represent millions of dollars. Construction & operations period are in years.
Profit_BT = 18.0
Construction Rate 6.0% Tax Rate 48% Nconstruction 2
Finance Rate 0.0% Enterprise Rate 2.5% Noperations 10 Yields Tot.Cap.Inv. TI=FC+WC+SU
ROI_BT = 28.8%
Fixed Capital 36.0 a-3 0.00
alpha_Working Capital 20% a-2 0.00 b_1 0.80 TI=FC+WC+SU
alpha_Start-Up Capital 10% a-1 0.50 b_2 0.90
alpha_Salvage Value 3% a0 0.50 b_3 0.95
Capital In (+) Discount Discounted
Year DesignConstruction Period or Out (-) Factors Cash Flows
-3 Fixed Capital in Y-3 0.0 1.191 0.0
-2 Fixed Capital in Y-2 0.0 1.124 0.0
-1 Fixed Capital in Y-1 -18.0 1.060 -19.1
0 Fixed Capital in Y0 -18.0 1.000 -18.0
0 Working Capital -22.9 1.000 -22.9
0 Start-Up Capital -3.6 1.000 -3.6
0 Total of Capital Outlays -63.6
(=Sum of Constr. DCFs)
0 Total Capital Investment 63.6
(=Proceeds of Bond Issue)
Profit Bond Depreciation Profit Cash
Operations Period Before Taxes Financing Allowed After Taxes Flows
1 14.4 0.0 -4.0 5.4 9.4 0.975 9.2
2 16.2 0.0 -4.0 6.4 10.3 0.951 9.8
3 17.1 0.0 -4.0 6.8 10.8 0.928 10.0
4 18.0 0.0 -4.0 7.3 11.3 0.905 10.2
5 18.0 0.0 -4.0 7.3 11.3 0.882 9.9
6 18.0 0.0 -4.0 7.3 11.3 0.860 9.7
7 18.0 0.0 -4.0 7.3 11.3 0.839 9.4
8 18.0 0.0 -4.0 7.3 11.3 0.818 9.2
9 18.0 0.0 -4.0 7.3 11.3 0.798 9.0
10 18.0 0.0 -4.0 7.3 11.3 0.778 8.8
10 Working Capital 22.9 22.9 0.778 17.8
10 Salvage Value 1.1 0.6 0.6 0.778 0.4
10 Pay-Off TCI 0.0 -63.6 0.778 -49.5
WC & SV Total Profit Bond Interest Total Total Profit Total NPV(0) NPV-proj [NPV(0)
All Figures Represent Recovery Before Taxes Payments Depreciation After Taxes Cash Flow Discounted to EOY(-x)]
PV of Operations==> 18.3 151.2 0.0 34.6 61.1 64.0 0.4 0.4
Bond Total Capital
Repayment Recovery
-49.5 52.4 Total Cash
IRR Calculation Net Present Flow as Sum NPV(0) Avg. NPV_proj Avg.
-12% Value of Bonds of NPVs Over y Years Over z Years
with the IRR function 14.1 64.0 0.1% 0.1%
x = Nconstruction
y = Nooperations
z = Nconstruction + Noperations
NPV Increase per Year
normalized/annualized
CR and FR can be chosen independently. WC and SV are converted from inventory to Profit_BT in Year 10.
Table 3. Risk-Based Profitability Analysis: Net Present Value Via Discounted Cash Flow
Establishes the ROI_BT Based on TI (also TCI) and the Annual % Increase of NPV (Normalized by TCI)
Fixed Capital and Profit_BT are the two independent variables.
using Capitalization =

40. 40
Appendix E: Sensitivity Analysis
Normalized Net Present Value (NPV%) is a metric that is integral to the financing of chemical
plants. This is a value that reflects the earning potential of an investment, and optimizations of
this value maximizes the annual rate of return and increase on the investment. This key value
determines the viability of a design against others.
Figure E.1: Sensitivity of NPV % compared to the Enterprise Rate. The NPV% of the
investment decreases with increasing enterprise rate because this endeavor becomes less fruitful
when the enterprise is doing well.
0 10 20 30 40 50
0
2
4
6
8
10
12
14
Enterprise Rate, [%]
NPV%

41. 41
Figure E.2: Sensitivity of NPV% compared to the tax rate.
Figure E.3: Sensitivity of NPV% compared to the finance rate.
0 10 20 30 40 50
5
5.5
6
6.5
7
7.5
8
8.5
9
Tax Rate, [%]
NPV%
0 10 20 30 40 50
-8
-6
-4
-2
0
2
4
6
8
10
Finance Rate, [%]
NPV%

42. 42
Figure E.4: Sensitivity of the proposed plant design with respect to fluctuations in the value of
all reactants and products at a tax rate of 25%.
-40 -30 -20 -10 0 10 20 30 40
-10
-5
0
5
10
15
20
Percent Change in Price, [%]
NPV%
Dimethyl Carbonate
Methanol

43. 43
Appendix F: Flow Diagrams
Figure F.1: Mass and Price Flowsheet
E-1
Fresh Oxygen
6.4e3 kg/h
$20.6 MM/yr
Fresh Carbon Monoxide
11.3e3 kg/h
$17 MM/yr
Vapor Recycle
129e3 kg/h
Compressor
$2.20 MM
$0.24 MM/yrFresh Methanol
12.6e3 kg/h
$53.8 MM/yr
P-11
P-15
CSTR
$0.13 MM
$0.66 MM/yr
Cooler A
$0.05MM
Negligible/yr
Flash Drum A
$0.08 MM
Heat Exchanger A
$0.10 MM
Flash Drum B
$0.04 MM
P-25
Heater A
$0.28 MM
$0.63 MM/yr
Wastewater
3.6e3 kg/h
Negligible/yr
Heat Exchanger B
$0.21 MM
Heater B
$0.12 MM
$1.10 MM/yr Cooler C
$0.02 MM
$0.48 MM/yr
99.8 wt% Dimethyl Carbonate
18.0e3 kg/h
$136 M/yr
P-48
Cooler B
$0.08 MM
$0.35 MM/yr
P-57
Vapor Recovery System
$2.00 MM/yr
Column 1 (44 stages)
(MeOH+DMC/Water)
$3.88 MM
$4.50 MM/yr
Carbon Dioxide Purge
8.9e3 kg/h
DMC/MeOH Recycle
14.6e3 kg/h
Column 2 (21 stages)
(MeOH + DMC/DMC)
$2.67 MM
$8.95 MM/yr
Column 3 (30 Stages)
(MeOH + DMC/MeOH)
$5.87 MM
$7.87 MM/yr
Methanol Recycle
41.3e3 kg/h
Cooler D
Negligible Cap
$0.06 MM/yr

44. 44
Figure F.2: Aspen HYSYS design

45. 45
Table F.1: Aspen HYSYS design flowsheet

46. 46
Figure F.3: Piping and instrumentation diagram displaying control loops.

47. 47
Appendix G: Matlab Code
% ChE 184B - Design Project 2
% Ramiro Ramirez
% Russell Wong
clear all;
close all;
clear;
clc;
% Target Production of Dimethyl Carbonate
% 150 MMkg/year
% 17857 kg/hr
% 4.96 kg/s
% 55.06 mol/s
% Components Data
% Molar Masses of Components, [kg/mol]
MMdmc = (90.08)/1000; % Dimethyl Carbonate
MMo2 = (32)/1000; % Oxygen
MMco = (28.01)/1000; % Carbon Monoxide
MMco2 = (44.009)/1000; % Carbon Dioxide
MMh2o = (18.01528)/1000; % Water
MMm = (32.04)/1000; % Methanol
% Values of Components
Val_dmc = 0.90; % [$/kg]
Val_o2 = 0.38; % [$/kg]
Val_co = 0.18; % [$/kg]
Val_m = 0.49; % [$/kg]
iteration = 1;
for MolarRm = [1/3]
%Initial parameters
Temp = 130 +273.15; % Kelvin, range, [80-130C]
Pres = 40; % atm, range, [10-40atm]
R = 1.987; % cal/molK
RR = 8.31446*10^-5; % (m^3*bar)/(K*mol)
%Design equations specifications
k1 = (1.4*10^11)*exp(-24000/(R*Temp));
k2 = (5.6*10^12)*exp(-22700/(R*Temp));
% Henry's Constant, [bar]
KH_o2 = 3179;
KH_co = 3107;
KH_co2 = 158;
% Calculating Partial Pressure of Methanol

48. 48
% A = 7.97, B = 1521.31, C = 234
ppm = 10^(5.15043 - (1549.48/(236.642 + 130))); % [bar]
%Pres = Pres - ppm;
iter = 1;
for tao = [25 35 50 75 100 120 145 165]
% Initial Flowrates into the reactor
MR = 25; % Molar ratio of CO to CO2
PPco = (MR/(MR+1))*Pres; % Partial Pressure of CO
PPo2 = (1/(1+MR))*Pres; % Partial Pressure of O2
xco = PPco/KH_co;
xo2 = PPo2/KH_o2;
% Calculating Molar Ratio of Methanol to Vapor (Oxy/CO)
%MolarRm = 1/3;
c_a0 = 20.32; % At specified conditions
c_b0 = (PPo2/(0.0821*Temp))*(PPo2/KH_o2);
c_c0 = (PPco/(0.0821*Temp))*(PPco/KH_co);
q = 30; % Volumetric Flowrate, [L/s]
Output = 112;
Target = 55.06;
while Output >= 111;
c_a = fsolve(@(c_a)(tao*(k1*(c_a.^2)*(c_b0.^0.5)*(q))-55.06),1); % Formation of dmc
P_dmc = 55.06;
Fm0 = c_a0*q; % mol/L
F_m = c_a*q;
Output = (Fm0 - F_m);
q = q - 0.1;
end
% Determining vapor rate from molar ratio
Fvap = Fm0*(1/MolarRm);
Fco0 = Fvap*(MR/(1+MR));
Fo20 = Fvap*(1/(1+MR));
P_co2 = tao*(k2*(c_b0^0.5)*(q));
P_h2o = tao*(k1*(c_a^2)*(c_b0^0.5)*(q));
F_co = Fco0 - tao*(k1*(c_a^2)*(c_b0^0.5)*(q) + k2*(c_b0^0.5)*(q));
F_o2 = (Fo20 - tao*(0.5*k1*(c_a^2)*(c_b0^0.5)*(q) + 0.5*k2*(c_b0^0.5)*(q)));
% Defining Volumetric Flowrate of gas phase
qv = (Fco0 + Fo20)*((0.0821*Temp)/Pres); % [L/s]
% Defining Selectivity, Conversion and Yield
Conversion_m = (Fm0-F_m)/Fm0;
Conversion_o2 = (Fo20-F_o2)/(Fo20);
Selectivity_m = P_dmc/(Fm0-F_m);
Selectivity_o2 = P_dmc/(Fo20-F_o2);

49. 49
Yield_o2 = Selectivity_o2.*Conversion_o2;
% Converting to mass flowrates
MF_m = F_m * MMm; % [kg/s]
MF_o2 = F_o2 * MMo2; % [kg/s]
MP_dmc = P_dmc * MMdmc; % [kg/s]
MP_co2 = P_co2 * MMco2; % [kg/s]
MF_co = F_co * MMco; % [kg/s]
MP_h2o = P_h2o * MMh2o; % [kg/s]
% Initial
MFm0 = Fm0*MMm;
MFo20 = Fo20*MMo2;
MFco0 = Fco0*MMco;
% Determining Recycle Rates
R_m = F_m;
R_o2 = F_o2;
R_co = F_co;
MR_m = R_m*MMm;
MR_o2 = R_o2*MMo2;
MR_co = R_co*MMco;
% Determining what you have to buy
Fm_fed = Fm0 - R_m;
Fo2_fed = Fo20 - R_o2;
Fco_fed = Fco0 - R_co;
MFm_fed = Fm_fed*MMm;
MFo2_fed = Fo2_fed*MMo2;
MFco_fed = Fco_fed*MMco;
% Checking to see if all mass adds up
Min = MFm0 + MFo20 + MFco0; % [kg/s]
Mout = MP_h2o + MP_dmc + MP_co2+ MF_m + MF_co + MF_o2; % [kg/s]
MBal = Min-Mout; % [kg/s]
% Measuting Water Build Up
Water_frac = MP_h2o/(MP_dmc + MF_m + MP_h2o);
% Determining Mole Fraction into Sep System
MolesOut = R_m + R_o2 + R_co + P_dmc + P_co2 + P_h2o;
Frac_m(iter) = R_m/MolesOut;
Frac_o2(iter) = R_o2/MolesOut;
Frac_co(iter) = R_co/MolesOut;
Frac_dmc(iter) = P_dmc/MolesOut;
Frac_co2(iter) = P_co2/MolesOut;
Frac_h2o(iter) = P_h2o/MolesOut;
% Determining Mass Fractions entering the seperation system
MFrac_m = MR_m/Mout;
MFrac_o2 = MR_o2/Mout;
MFrac_co = MR_co/Mout;
MFrac_dmc = MP_dmc/Mout;
MFrac_co2 = MP_co2/Mout;
MFrac_h2o = MP_h2o/Mout;
% Determining the volume of the reactor

50. 50
V = (tao*q); %+ (tao*qv); % [L]
% Measuring Economic Potential
Cost_Raw = (MFm_fed*Val_m*8400*3600) + (MFco_fed*Val_co*8400*3600) +
(MFo2_fed*Val_o2*3600*8400);
Profit_Raw = MP_dmc*Val_dmc*8400*3600;
EP = Profit_Raw - Cost_Raw;
Conversion_meoh(iter) = Conversion_m;
Conversion_oxy(iter) = Conversion_o2;
Selectivity_meoh(iter) = Selectivity_m;
Selectivity_oxy(iter) = Selectivity_o2;
RVol(iter) = V; % Reactor Volume, [L]
Econ_P(iter) = EP; % Economic Potential, [$/yr]
Balance(iter) = MBal; % Mass Balance
Water_per(iter) = Water_frac*100; % Mass percent of water
Methanol(iter) = Fm0-F_m;
DMC(iter) = P_dmc;
Oxygen(iter) = Fo20 - F_o2;
CarbonMonoxide(iter) = Fco0 - F_co;
CarbonDioxide(iter) = P_co2;
H2O(iter) = P_h2o;
Recycle_m(iter) = MR_m;
Recycle_o2(iter) = MR_o2;
Recycle_co(iter) = MR_co;
Flow_in(iter) = Min;
Fresh_m(iter) = MFm_fed;
Fresh_co(iter) = MFco_fed;
Fresh_o2(iter) = MFo2_fed;
% Mass Balance Check
if sum(abs(Balance)) > 0.1
disp('Mass Balance is Off');
end
%OxygenConversion(iteration,:) = Conversion_oxy;
%OxygenSelectivity(iteration,:) = Selectivity_oxy;
%iteration = iteration + 1;
%end
%%%%% Design and Analysis of Separation System
% Required Thermodynamic Data
% Heats of Vaporization
Hvap_m = 35278; % [J/mol]
Hvap_dmc = 37700;
Hvap_h2o = 40000;
% Saturated Liquid Densities
rho_m = 791.80;

51. 51
rho_h2o = 999.97;
rho_dmc = 1000;
% Design of the Vapor Recovery System
ff = (P_co2)/(P_co2 + R_o2 + R_co); % Fraction of feed that exists in the enriched stream
zco2 = (P_co2)/(P_co2 + R_o2 + R_co);
epss = ff/zco2;
Wmin = (1-zco2)*log(1/(1-zco2*epss))-epss*zco2*log(zco2); %+ zco2*(1-epss)*log((1-epss)/(1-zco2*epss))
Wmin = Wmin*(P_co2 + R_o2 + R_co)*Temp*8.314; % Minimum energy required for separation, [W]
OpCost_Vap = Wmin*(1/1000)*6*(0.06)*8400; % Operating cost, [$/yr]
% Design of Flash Separators
% Flash Drum #1
flowin1 = 265; % [m^3/hr]
fd1_vol = flowin1*(1/6); % [m^3]
fd1_dia = (fd1_vol/(8*pi))^(1/3); % Diameter to height ratio 1:3
COS_fd1 = (1600/280)*(1*1 - 1 + 1.38*6)*5000;
COT_fd1= (1600/280)*(1.0+0+0)*500;
CS_fd1 = COS_fd1*(fd1_dia/1)*(fd1_dia/6.1)^0.82;
CT_fd1 = COT_fd1*((fd1_dia/1)^1.8)*(fd1_dia/6.1);
fd1_cost = CS_fd1+CT_fd1; % Flash Drum 1: Cap Cost
% Flash Drum #2
flowin2 = 98; % [m^3/hr]
fd2_vol = flowin2*(1/6); % [m^3]
fd2_dia = (fd2_vol/(8*pi))^(1/3); % Diameter to height ratio 1:3
COS_fd2 = (1600/280)*(1*1 - 1 + 1.38*6)*5000;
COT_fd2= (1600/280)*(1.0+0+0)*500;
CS_fd2 = COS_fd2*(fd2_dia/1)*(fd2_dia/6.1)^0.82;
CT_fd2 = COT_fd2*((fd2_dia/1)^1.8)*(fd2_dia/6.1);
fd2_cost = CS_fd2+CT_fd2; % Flash Drum 2: Cap Cost
% Flash Drum #3
%flowin3 = 294.9; % [m^3/hr]
%fd3_vol = flowin3*(1/6); % [m^3]
%fd3_dia = (fd3_vol/(8*pi))^(1/3); % Diameter to height ratio 1:3
%COS_fd3 = (1600/280)*(1*1 - 1 + 1.38*6)*5000;
%COT_fd3= (1600/280)*(1.0+0+0)*500;
%CS_fd3 = COS_fd3*(fd3_dia/1)*(fd3_dia/6.1)^0.82;
%CT_fd3 = COT_fd3*((fd3_dia/1)^1.8)*(fd3_dia/6.1);
%fd3_cost = CS_fd3+CT_fd3; % Flash Drum 3: Cap Cost
fd3_cost = 0;
% Distillation Column #1 (H2O / Methanol,DMC)
F1 = F_m + P_dmc + P_h2o; % molar flowrate of A,B,C,D [mol/s]
q1 = 1; % Sat liquid
%Tcol1 = 167.39322; % Celcius, Bubble Point of mixture
Tcol1 = 90;
Pcol1 = 1; % bar

52. 52
% Feed Composition, molar
z1_m = F_m/F1;
z1_h2o = P_h2o/F1;
z1_dmc = P_dmc/F1;
% Distillate and Bottoms flowrate assuming 100% recovery of light key in
% the distillate, Molar Flowrates
D1 = F1*z1_m + F1*z1_dmc; % mol/s
B1 = F1*z1_h2o; % mol/s
% Establishing the reflux ration (HYSYS)
R1 = 4.0;
% Establising the value of S (HYSYS)
S1 = 30;
% Establishing the number of stages (HYSYS)
N1_real = 44;
% Calculating Vapor Rate inside column
VB1 = S1*B1;
VT1 = (R1 + 1) * D1;
% Calculating Heat Load
% Latent heats of vaporiztion, 100% Benzene in the distillate
% Saturated liquid products
% Using a weighted average for the bottoms
lambdaD1 = Hvap_m*(z1_m/(z1_m+z1_dmc)) + Hvap_dmc*(z1_dmc/(z1_dmc+z1_m));
lambdaB1 = (Hvap_h2o);
% Calculating the heat loads
Qc1 = 0.91*lambdaD1*VT1; % Watts
Qr1 = 1.5*lambdaB1*VB1; % Watts
% Column Sizing
Phi_flood = 0.6;
Frac_flow = 0.8;
c0 = 252; % Assuming 24 inch tray spacing, Table 6.1
Ht = 0.31; % Tray spacing in meters, (12 inches)
% Area of the column, Eq. 6.12
% Molecular weight of vapor, Mv
% Weighted average
Mv1 = ((z1_m*F1*MMm)/D1) + ((z1_dmc*F1*MMdmc)/D1); % kg/mol
Mv1 = Mv1 * 1000; % g/mol
% Calculating Weighted densities of liquid and vapor
% Liquid Density
rho_l1 = ((z1_h2o*F1*rho_h2o)/B1); % g/L
% Vapor Density
%rho_v1 = ((z1_t*F1*rho_t)/D1) + ((z1_b*F1*rho_b)/D1); % g/L

53. 53
rho_v1 = (2*Mv1)/(0.0821*120);
% Calculation of the area
Area_col1 = (Mv1/(sqrt(rho_l1*rho_v1)))*(1/(Phi_flood*c0))*(1/Frac_flow)*VT1*(1/1000)*3600;
% Defining minimum height
Ht_min = 3 * Ht;
% Calculating column diameter
Dia_col1 = 2*sqrt(Area_col1/pi); % Meters
% Calculating column height
Height_col1 = Ht_min + N1_real*Ht; % Meters
%Height_col1_hys = Ht_min + 44*Ht;
% Calculating Heat Exchanger Areas
% Using Table 6.2 to determine heat xfer coefficients
U1_c = 800; % Condensor [W/m^2*K]
U1_r = 800; % Reboiler [w/m^2*K]
TwIN = 30; % Cooling Water inlet, [C]
TwOUT = 50; % Cooling Water outlet, [C]
LMtemp1_con =((Tcol1-TwOUT) - (Tcol1-TwIN)) / ((log(Tcol1-TwOUT)) - log(Tcol1-TwIN)); % Log mean
temperature
LMtemp1_reb = 99.65 - Tcol1;
% HYSYS values
Qr1_hys = (7.55*10^4)*1000;
Qc1_hys = (7.70*10^4)*1000;
% Calculating condensor and reboiler surface area
Area_reb1 = Qr1_hys/(U1_r*LMtemp1_reb); % [m^2]
Area_con1 = Qc1_hys/(U1_c*LMtemp1_con); % [m^2]
% For HYSYS calculations
%%%Qr1 = Qr1_hys;
%%%Qc1 = Qc1_hys;
Area_reb1_hys = Qr1_hys/(U1_r*LMtemp1_reb); % [m^2]
Area_con1_hys = Qc1_hys/(U1_c*LMtemp1_con); % [m^2]
%Heat Exchanger Costs
HExFC = (.8+0)*1.00;
HExIC_reb1=(1600/280)*(101.3*Area_reb1_hys^(0.65)*(2.29+HExFC));
HExIC_con1=(1600/280)*(101.3*Area_con1_hys^(0.65)*(2.29+HExFC));
% Column Cost
COS1 = (1600/280)*(1*1 - 1 + 1.38*6)*5000;
COT1 = (1600/280)*(1.0+0+0)*500;
CS1 = COS1*(Dia_col1/1)*(Height_col1/6.1)^0.82;
CT1 = COT1*((Dia_col1/1)^1.8)*(Height_col1/6.1);
Col1_cost = CS1+CT1 + HExIC_reb1 + HExIC_con1; % Total cost
% Column Operating Cost, (Cost of Utilities)
steamreb1 = ((Qr1_hys/1000) * (1/2213)) * 8400 * 3600; % kg/yr steam
steamreb1_cost = (steamreb1/1000)*2.20; % Op. cost, [$/yr]
cwatercon1 = ((Qc1_hys/1000) * (1/(4.2*20)))*8400*3600; % kg/yr cooling water

54. 54
cwatercon1_cost = (cwatercon1/1000) * 0.08; % Op. cost, [$,yr]
ww1_cost = (steamreb1/1000)*0.06 + (cwatercon1/1000) * 0.06; % Waste Water Cost, [$/yr]
% Calculating total operating cost for column #1
col1_opcost = cwatercon1_cost + steamreb1_cost; % + ww1_cost; % [$/yr]
% Distillation Column #2 (DMC / DMC-Methanol)
F2 = P_dmc + F_m; % molar flowrate of A,B,C,D [mol/s]
q2 = 1; % Sat liquid
Tcol2 = 200;
Pcol2 = 40; % bar
% Feed Composition, molar
z2_m = F_m/F2;
z2_dmc = P_dmc/F2;
% Distillate and Bottoms flowrate assuming 100% recovery of light key in
% the distillate, Molar Flowrates
D2 = 0.75*F2; % mol/s
B2 = 0.25*F2; % mol/s
% Establishing the reflux ration (HYSYS)
R2 = 3.5;
% Establising the value of S (HYSYS)
S2 = 105;
% Establishing the number of stages (HYSYS)
N2_real = 21;
% Calculating Vapor Rate inside column
VB2 = S2*B2;
VT2 = (R2 + 1) * D2;
% Calculating Heat Load
% Using a weighted average for the bottoms
lambdaD2 = Hvap_m;
lambdaB2 = (Hvap_dmc);
% Calculating the heat loads
Qc2 = 3*lambdaD2*VT2; % Watts
Qr2 = 0.25*lambdaB2*VB2; % Watts
% Column Sizing
Phi_flood = 0.6;
Frac_flow = 0.8;
c0 = 252; % Assuming 12 inch tray spacing, Table 6.1
Ht = 0.31; % Tray spacing in meters, (12 inches)
% Area of the column, Eq. 6.12
% Molecular weight of vapor, Mv

55. 55
% Weighted average
Mv2 = ((0.80*F2*MMm)/D2) + ((0.2*F2*MMdmc)/D2); % kg/mol
Mv2 = Mv2 * 1000; % g/mol
% Calculating Weighted densities of liquid and vapor
% Liquid Density
rho_l2 = ((z1_dmc*F2*rho_dmc)/B2); % g/L
% Vapor Density
%rho_v1 = ((z1_t*F1*rho_t)/D1) + ((z1_b*F1*rho_b)/D1); % g/L
rho_v2 = (2*Mv2)/(0.0821*120);
% Calculation of the area
Area_col2 = (Mv2/(sqrt(rho_l2*rho_v2)))*(1/(Phi_flood*c0))*(1/Frac_flow)*VT2*(1/1000)*3600;
% Defining minimum height
Ht_min = 3 * Ht;
% Calculating column diameter
Dia_col2 = 2*sqrt(Area_col2/pi); % Meters
% Calculating column height
Height_col2 = Ht_min + N2_real*Ht; % Meters
%Height_col1_hys = Ht_min + 44*Ht;
% Calculating Heat Exchanger Areas
% Using Table 6.2 to determine heat xfer coefficients
U1_c = 800; % Condensor [W/m^2*K]
U1_r = 800; % Reboiler [w/m^2*K]
TwIN = 30; % Cooling Water inlet, [C]
TwOUT = 50; % Cooling Water outlet, [C]
LMtemp2_con =((Tcol2-TwOUT) - (Tcol2-TwIN)) / ((log(Tcol2-TwOUT)) - log(Tcol2-TwIN)); % Log mean
temperature
LMtemp2_reb = 221 - Tcol2;
% Calculating condensor and reboiler surface area
Area_reb2 = Qr2/(U1_r*LMtemp2_reb); % [m^2]
Area_con2 = Qc2/(U1_c*LMtemp2_con); % [m^2]
% HYSYS values
Qr2_hys = (1.6*10^5)*1000;
Qc2_hys = (1.4*10^5)*1000;
% For HYSYS calculations
%%%Qr1 = Qr1_hys;
%%%Qc1 = Qc1_hys;
Area_reb2_hys = Qr2_hys/(U1_r*LMtemp2_reb); % [m^2]
Area_con2_hys = Qc2_hys/(U1_c*LMtemp2_con); % [m^2]
%Heat Exchanger Costs
HExFC = (.8+0)*1.00;
HExIC_reb2=(1600/280)*(101.3*Area_reb2_hys^(0.65)*(2.29+HExFC));
HExIC_con2=(1600/280)*(101.3*Area_con2_hys^(0.65)*(2.29+HExFC));
% Column Cost
COS2 = (1600/280)*(1*1 - 1 + 1.38*6)*5000;
COT2 = (1600/280)*(1.0+0+0)*500;

56. 56
CS2 = COS2*(Dia_col2/1)*(Height_col2/6.1)^0.82;
CT2 = COT2*((Dia_col2/1)^1.8)*(Height_col2/6.1);
Col2_cost = CS2+CT2 + HExIC_reb2 + HExIC_con2; % Total cost
% Column Operating Cost, (Cost of Utilities)
steamreb2 = ((Qr2_hys/1000) * (1/2213)) * 8400 * 3600; % kg/yr steam
steamreb2_cost = (steamreb2/1000)*2.25; % Op. cost, [$/yr]
cwatercon2 = ((Qc2_hys/1000) * (1/(4.2*20)))*8400*3600; % kg/yr cooling water
cwatercon2_cost = (cwatercon2/1000) * 0.08; % Op. cost, [$,yr]
ww2_cost = (steamreb2/1000)*0.06 + (cwatercon2/1000) * 0.06; % Waste Water Cost, [$/yr]
% Calculating total operating cost for column #2
col2_opcost = cwatercon2_cost + steamreb2_cost + ww1_cost; % [$/yr]
% Distillation Column #3 (DMC / DMC-Methanol)
F3 = 0.1*P_dmc + F_m; % molar flowrate of A,B,C,D [mol/s]
q3 = 1; % Sat liquid
%Tcol1 = 167.39322; % Celcius, Bubble Point of mixture
Tcol3 = 63;
Pcol3 = 1; % bar
% Feed Composition, molar
z3_m = F_m/F3;
z3_dmc = 0.1*P_dmc/F3;
% Distillate and Bottoms flowrate assuming 100% recovery of light key in
% the distillate, Molar Flowrates
D3 = z3_dmc*F3 + 0.5*z3_m*F3; % mol/s
B3 = 0.5 * z2_m*F3; % mol/s
% Establishing the reflux ration (HYSYS)
R3 = 3;
% Establising the value of S (HYSYS)
S3 = 10;
% Establishing the number of stages (HYSYS)
N3_real = 30;
% Calculating Vapor Rate inside column
VB3 = S3*B3;
VT3 = (R3 + 1) * D3;
% Calculating Heat Load
% Using a weighted average for the bottoms
lambdaD3 = Hvap_m;
lambdaB3 = (Hvap_m);
% Calculating the heat loads
Qc3 = 5.5*lambdaD3*VT3; % Watts
Qr3 = 2*lambdaB3*VB3; % Watts

57. 57
% Column Sizing
Phi_flood = 0.6;
Frac_flow = 0.8;
c0 = 252; % Assuming 12 inch tray spacing, Table 6.1
Ht = 0.31; % Tray spacing in meters, (12 inches)
% Area of the column, Eq. 6.12
% Molecular weight of vapor, Mv
% Weighted average
Mv3 = ((0.80*F3*MMm)/D3) + ((0.2*F3*MMdmc)/D3); % kg/mol
Mv3 = Mv3 * 1000; % g/mol
% Calculating Weighted densities of liquid and vapor
% Liquid Density
rho_l3 = ((z1_m*F3*rho_m)/B3); % g/L
% Vapor Density
%rho_v1 = ((z1_t*F1*rho_t)/D1) + ((z1_b*F1*rho_b)/D1); % g/L
rho_v3 = (2*Mv3)/(0.0821*120);
% Calculation of the area
Area_col3 = (Mv3/(sqrt(rho_l3*rho_v3)))*(1/(Phi_flood*c0))*(1/Frac_flow)*VT3*(1/1000)*3600;
% Defining minimum height
Ht_min = 3 * Ht;
% Calculating column diameter
Dia_col3 = 2*sqrt(Area_col3/pi); % Meters
% Calculating column height
Height_col3 = Ht_min + N3_real*Ht; % Meters
%Height_col1_hys = Ht_min + 44*Ht;
% Calculating Heat Exchanger Areas
% Using Table 6.2 to determine heat xfer coefficients
U1_c = 800; % Condensor [W/m^2*K]
U1_r = 800; % Reboiler [w/m^2*K]
TwIN = 30; % Cooling Water inlet, [C]
TwOUT = 50; % Cooling Water outlet, [C]
LMtemp3_con =((Tcol3-TwOUT) - (Tcol3-TwIN)) / ((log(Tcol3-TwOUT)) - log(Tcol3-TwIN)); % Log mean
temperature
LMtemp3_reb = 64.2 - Tcol3;
% Calculating condensor and reboiler surface area
Area_reb3 = Qr3/(U1_r*LMtemp3_reb); % [m^2]
Area_con3 = Qc3/(U1_c*LMtemp3_con); % [m^2]
% HYSYS values
Qr3_hys = (1.24*10^5)*1000;
Qc3_hys = (1.60*10^5)*1000;
% For HYSYS calculations
%%%Qr1 = Qr1_hys;
%%%Qc1 = Qc1_hys;

58. 58
Area_reb3_hys = Qr3_hys/(U1_r*LMtemp3_reb); % [m^2]
Area_con3_hys = Qc3_hys/(U1_c*LMtemp3_con); % [m^2]
%Heat Exchanger Costs
HExFC = (.8+0)*1.00;
HExIC_reb3=(1600/280)*(101.3*Area_reb3_hys^(0.65)*(2.29+HExFC));
HExIC_con3=(1600/280)*(101.3*Area_con3_hys^(0.65)*(2.29+HExFC));
% Column Cost
COS3 = (1600/280)*(1*1 - 1 + 1.38*6)*5000;
COT3 = (1600/280)*(1.0+0+0)*500;
CS3 = COS3*(Dia_col3/1)*(Height_col3/6.1)^0.82;
CT3 = COT3*((Dia_col3/1)^1.8)*(Height_col3/6.1);
Col3_cost = CS3+CT3 + HExIC_reb3 + HExIC_con3; % Total cost
% Column Operating Cost, (Cost of Utilities)
steamreb3 = ((Qr3_hys/1000) * (1/2213)) * 8400 * 3600; % kg/yr steam
steamreb3_cost = (steamreb3/1000)*2.20; % Op. cost, [$/yr]
cwatercon3 = ((Qc3_hys/1000) * (1/(4.2*20)))*0.9*8400*3600; % kg/yr cooling water
cwatercon3_cost = (cwatercon3/1000) * 0.08; % Op. cost, [$,yr]
ww3_cost = (steamreb3/1000)*0.06 + (cwatercon3/1000) * 0.06; % Waste Water Cost, [$/yr]
% Calculating total operating cost for column #1
col3_opcost = cwatercon3_cost + steamreb3_cost + ww1_cost; % [$/yr]
%%% Costing Plant Heaters and Coolers
% Costing Heater 1
% Calculating Heat Exchanger Areas
% Using Table 6.2 to determine heat xfer coefficients
U1_c = 800; % Condensor [W/m^2*K]
U1_r = 800; % Reboiler [w/m^2*K]
TwIN = 30; % Cooling Water inlet, [C]
TwOUT = 50; % Cooling Water outlet, [C]
LMtemp_cooler1 =((120-TwOUT) - (120-TwIN)) / ((log(120-TwOUT)) - log(120-TwIN)); % Log mean
temperature
LMtemp_cooler2 =((97-TwOUT) - (97-TwIN)) / ((log(97-TwOUT)) - log(97-TwIN));
LMtemp_cooler3 =((89-TwOUT) - (89-TwIN)) / ((log(89-TwOUT)) - log(89-TwIN));
LMtemp_cooler4 =((168-TwOUT) - (168-TwIN)) / ((log(168-TwOUT)) - log(168-TwIN));
LMtemp_heater1 = 11;
LMtemp_heater2 = 70;
LMtemp_cross1 = ((100-17) - (130-17)) / ((log(100-17)) - log(130-17));
LMtemp_cross2 = ((165-97) - (97-61)) / ((log(165-97)) - log(97-61));
Q_heater1 = (2.084*10^4)*1000;
Q_heater2 = (3.8*10^4)*1000;
Q_cooler1 = (1.206*10^4)*1000;
Q_cooler2 = (1.659*10^4)*1000;
Q_cooler3 = (2142)*1000;
Q_cooler4 = (168.3)*1000;

59. 59
Q_cross1 = (3.64*10^4)*1000;
Q_cross2 = (6.01*10^4)*1000;
Area_heater1 = Q_heater1/(U1_r*LMtemp_heater1); % [m^2]
Area_heater2 = Q_heater2/(U1_c*LMtemp_heater2); % [m^2]
Area_cooler1 = Q_cooler1/(U1_c*LMtemp_cooler1);
Area_cooler2 = Q_cooler2/(U1_c*LMtemp_cooler2);
Area_cooler3 = Q_cooler3/(U1_c*LMtemp_cooler3);
Area_cooler4 = Q_cooler4/(U1_c*LMtemp_cooler4);
Area_cross1 = Q_cross1/(U1_c*LMtemp_cross1);
Area_cross2 = Q_cross2/(U1_c*LMtemp_cross2);
%Heat Exchanger Costs
HExFC = (.8+0)*1.00;
HExIC_heater1 = (1600/280)*(101.3*Area_heater1^(0.65)*(2.29+HExFC));
HExIC_heater2 = (1600/280)*(101.3*Area_heater2^(0.65)*(2.29+HExFC));
HExIC_cooler1 = (1600/280)*(101.3*Area_cooler1^(0.65)*(2.29+HExFC));
HExIC_cooler2 = (1600/280)*(101.3*Area_cooler2^(0.65)*(2.29+HExFC));
HExIC_cooler3 = (1600/280)*(101.3*Area_cooler3^(0.65)*(2.29+HExFC));
HExIC_cooler4 = (1600/280)*(101.3*Area_cooler4^(0.65)*(2.29+HExFC));
HExIC_cross1 = (1600/280)*(101.3*Area_cross1^(0.65)*(2.29+HExFC));
HExIC_cross2 = (1600/280)*(101.3*Area_cross2^(0.65)*(2.29+HExFC));
TotalCoolerCost = HExIC_cooler1 + HExIC_cooler2 + HExIC_cooler3 + HExIC_cooler4;
TotalHeaterCost = HExIC_heater1 + HExIC_heater2;
TotalCrossCost = HExIC_cross1 + HExIC_cross2;
% Operating Costs of Coolers and Heaters
% Total Cooling Heat Load
cooltotal = (138.3) + (1.206*10^4) + (1.659*10^4) + 2142; % [kw]
cwatercoolers = cooltotal*(1/(4.2*20))*8400*3600;
cwatercoolers_cost = (cwatercoolers/1000) * 0.08; % Op Cost, [$/yr]
% Total Heater Heat Load
heattotal = (3.8*10^4) +(2.084*10^4) ;
steamheattotal = heattotal * (1/2213) * 8400 * 3600;
steamheattotal_cost = (steamheattotal/1000)*2.2; % Op Cost, [$/yr]
%%% Costing Gas Compressor
GC_power = 478.4*1000; % [Watts]
GC_hpower = GC_power/746; % [hp]
GC_opcost = (GC_power/1000)*8400*0.06; % Annual Operating Cost
GC_capcost = (1600/280)*(517.5)*((GC_hpower/0.8)^0.82)*(2.11 + 1); % Compressor Capital Cost
%%% Costing the CSTR reactor
cstr_d = 2*((3/(2*pi))^1/3); % Reactor diameter, [m]
cstr_h = cstr_d; % Reactor height, [m]

60. 60
PRFC = 1;
HExArea = 57.45; % Area of Heat Exchanger in m^2
HExFC = (.8+0)*1.00;
HExPC=(1600/280)*(101.3*HExArea^(0.65)*HExFC);
HExIC=(1600/280)*(101.3*HExArea^(0.65)*(2.29+HExFC));
% CSTR Operating Cost
Qcstr = (2.3*10^4)*1000;
cwater_cstr = ((Qcstr/1000) * (1/(4.20*20)))*8400*3600; % kg/yr cooling water
cwatercstr_cost = (cwater_cstr/1000) * 0.08; % Op. cost, [$,yr]
wwcstr_cost = (cwater_cstr/1000) * 0.06; % Waste Water Cost, [$/yr]
% Calculating total operating/capital cost for column #1
cstr_opcost = cwatercstr_cost; % [$/yr]
cstr_capcost = 500*(1600/280)*(101.9*cstr_d^(1.066)*cstr_h^(0.82)*(2.18+PRFC)) + HExIC; % Cap cost of
reactor, [$]
%%%%% Economic Analysis
SepCapCost = Col1_cost + Col2_cost + Col3_cost; % Capital cost of separations system
SepOpCost = col1_opcost + col2_opcost + col3_opcost; % Operating cost of sepratation
system
FS_capcost = fd1_cost + fd2_cost + fd3_cost; % Cap cost of flash separators
%OpCost_Vap % Vapor sep
%Pbt = Profit_Raw - Cost_Raw - SepOpCost - cstr_opcost - OpCost_Vap - GC_opcost - steamheattotal_cost -
cwatercoolers_cost; % Profit before taxes, [$/yr]
Pbt = (5.0*Val_dmc*3600*8400) - ( (3.63*Val_m*3600*8400) + (1.78*Val_o2*8400*3600) +
(3.12*Val_co*8400*3600))-SepOpCost - cstr_opcost - OpCost_Vap - GC_opcost - steamheattotal_cost -
cwatercoolers_cost;
WC = (((MFm_fed * Val_m * 8400 * 3600)./12).*2) + (((MFco_fed * Val_co * 8400 * 3600)./12).*2) +
(((MFo2_fed * Val_o2 * 8400 * 3600)./12).*2); % Working Cap, [$] Represents two months prod and feed
ISBL = (SepCapCost + cstr_capcost + FS_capcost + TotalCrossCost + TotalCoolerCost + TotalHeaterCost +
GC_capcost);
FCI= 2.28.*(ISBL); % Fixed capital Investment, [$]
SU = FCI.*0.1; % Start Up capital, [$]
TI = FCI + SU + WC; % Total Investment, [$]
TCI(iter) = (2.5.*ISBL) + WC; % Guthrie's Correlations
ROI_BT(iter) = (Pbt./TI).*100; % Return on Inv., [%/yr]
Dep = (FCI+SU).*(1/10); % Depreciation, [$]
POT = (TI - WC)./((1-0.48).*Pbt + 0.48.*Dep); % Pay out time, [yrs]
%%%% Include NPV calculations
FR=0.04; %finance rate
TR=0.25; %Tax rate
ER=0.12; %Enterprise Rate
CashFlow=zeros(10,1);
CashFlow(1,:)=(1-TR)*(Pbt.*0.8)+(1-TR)*(-TI.*FR)+TR*0.1*(1.1.*FCI);
CashFlow(2,:)=(1-TR)*(Pbt.*0.9)+(1-TR)*(-TI.*FR)+TR*0.1*(1.1.*FCI);
CashFlow(3,:)=(1-TR)*(Pbt.*0.95)+(1-TR)*(-TI.*FR)+TR*0.1*(1.1.*FCI);
for n=4:10
CashFlow(n,:)=(1-TR)*(Pbt)+(1-TR)*(-TI.*FR)+TR*0.1*(1.1.*FCI);
end
NCashFlow=CashFlow(:,1);
NPVb=0;
for i=1:10
NPVa(i)=(1+ER)^(-i)*(NCashFlow(i));
NPVb=NPVb+NPVa(i);

61. 61
end
%NPVa is before the WC, SV, and TCI and NPVb is the sum of the values
NPV(iter)=NPVb + (1+ER)^(-10)*(WC+0.03.*FCI-TI);
NPVper(iter)=100*(NPV(iter)/(1+ER)^(2))/(12*TI);
iter = iter + 1;
end
OxygenConversion(iteration,:) = Conversion_oxy;
OxygenSelectivity(iteration,:) = Selectivity_oxy;
NPVVV(iteration,:) = NPV;
TCIII(iteration,:) = TCI;
ROIBTTT(iteration,:) = ROI_BT;
NPVPERRR(iteration,:) = NPVper;
iteration = iteration + 1;
end
% Plotting all relevant figures
% Plotting S v X at diff molar ratios
figure(20)
xx1 = 0:0.01:1;
y2 = spline(OxygenConversion(1,:),OxygenSelectivity(1,:),xx1);
y3 = spline(OxygenConversion(2,:),OxygenSelectivity(2,:),xx1);
y4 = spline(OxygenConversion(3,:),OxygenSelectivity(3,:),xx1);
plot(xx1,y2,xx1,y3,xx1,y4,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Selectivity','FontSize',14,'FontName','Times New Roman');
legend('MR = 1/3','MR = 1','MR = 3');
axis([0 1 0 2]);
% Conversion vs Selectivity - Methanol
figure(1)
plot(Conversion_meoh, Selectivity_meoh);
xlabel('Conversion','FontSize',14,'FontName','Times New Roman');
ylabel('Selectivity','FontSize',14,'FontName','Times New Roman');
title('Selectivity VS Conversion, Methanol');
% Conversion vs Selectivity - Oxygen
figure(2)
xx1 = 0:0.01:1;
y1 = spline(Conversion_oxy,Selectivity_oxy,xx1);
plot(xx1,y1,'Linewidth',1.5);
xlabel('Conversion','FontSize',14,'FontName','Times New Roman');
ylabel('Selectivity','FontSize',14,'FontName','Times New Roman');
title('Selectivity VS Conversion, Oxygen');
axis([0 1 0 2]);
% Conversion vs Reactor Volume - Methanol
figure(3)
y_vol = spline(Conversion_oxy,RVol./1000,xx1);
plot(xx1,y_vol,'Linewidth',1.5);

62. 62
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Reactor Volume, [m^3]','FontSize',14,'FontName','Times New Roman');
% Conversion vs Economic Potential
figure(4)
plot(Conversion_oxy,Econ_P);
xlabel('Conversion, Oxygen','FontSize',14,'FontName','Times New Roman');
ylabel('Economic Potential, [$/yr]','FontSize',14,'FontName','Times New Roman');
% Mass percentage of Water in Solution
figure(5)
ywater = spline(Conversion_oxy,Water_per,xx1);
plot(xx1,ywater,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Mass Percent Water in Reactor','FontSize',14,'FontName','Times New Roman');
axis([0 1 0 7]);
% Mole Fraction of each component entering separation system
figure(6)
mfm = spline(Conversion_oxy,Frac_m,xx1);
mfh2o = spline(Conversion_oxy,Frac_h2o,xx1);
mfco = spline(Conversion_oxy,Frac_co,xx1);
mfco2 = spline(Conversion_oxy,Frac_co2,xx1);
mfo2 = spline(Conversion_oxy,Frac_o2,xx1);
mfdmc = spline(Conversion_oxy,Frac_dmc,xx1);
plot(xx1,mfm,xx1,mfh2o,xx1,mfco,xx1,mfco2,xx1,mfo2,xx1,mfdmc,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Mole Fraction','FontSize',14,'FontName','Times New Roman');
legend('Methanol','Water','Carbon Monoxide','Carbon Dioxide','Oxygen','DMC');
axis([0 1 0 0.3]);
% Fresh Feed Rates vs Conversion
figure(7)
yfm = spline(Conversion_oxy,Fresh_m,xx1);
yfo2 = spline(Conversion_oxy,Fresh_o2,xx1);
yfco = spline(Conversion_oxy,Fresh_co,xx1);
plot(xx1,yfm,xx1,yfo2,xx1,yfco,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Fresh Feed Rate, [kg/s]','FontSize',14,'FontName','Times New Roman');
legend('Methanol','Oxygen','Carbon Monoxide');
axis([0 1 1 5]);
% Recycle flow rate vs Conversion
figure(8)
yrm = spline(Conversion_oxy,Recycle_m,xx1);
yrco = spline(Conversion_oxy,Recycle_co,xx1);
yro2 = spline(Conversion_oxy,Recycle_o2,xx1);
plot(xx1,yrm,xx1,yro2,xx1,yrco,'Linewidth',1.5);
xlabel('Conversion, (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Recycle Flowrate, [kg/s]','FontSize',14,'FontName','Times New Roman')
legend('Methanol','Oxygen','Carbon Monoxide');
axis([0 1 -5 250]);
% Total Flowrate into the reactor vs Reactor Conversion
figure(9)
yflowin = spline(Conversion_oxy,Flow_in,xx1);

63. 63
plot(xx1,yflowin,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Flowrate Into Reactor, [kg/s]','FontSize',14,'FontName','Times New Roman');
axis([0 1 25 400]);
% Recycle of Methanol v Conversion
figure(10)
y_rec = spline(Conversion_oxy,Recycle_m,xx1);
plot(xx1,y_rec,'Linewidth',1.5)
xlabel('Oxygen Conversion');
ylabel('Recycle Flowrate, [mol/s]');
% TCI vs Conversion
figure(11)
ytci = spline(Conversion_oxy,TCI,xx1);
plot(xx1,ytci);
xlabel('Conversion (O_2)');
ylabel('Total Capitol Investment, [$]');
% NPV vs Conversion
figure(12)
ynpv = spline(Conversion_oxy,NPV,xx1);
plot(xx1,ynpv);
xlabel('Conversion (O_2)');
ylabel('Net Present Value');
% NPV percent vs Conversion
figure(13)
ynpvper = spline(Conversion_oxy,NPVper,xx1);
plot(xx1,ynpvper);
xlabel('Conversion (O_2');
ylabel('Net Present Value Percent');
% ROIBT vs Conversion
figure(14)
yroibt = spline(Conversion_oxy,ROI_BT,xx1);
plot(xx1,yroibt,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Return on Investment Before Taxes, [%/yr]','FontSize',14,'FontName','Times New Roman');
axis([0 1 -80 80]);
% Total Capital Investemt vs Conversion
figure(15)
y15 = spline(OxygenConversion(1,:),TCIII(1,:),xx1);
y16 = spline(OxygenConversion(2,:),TCIII(2,:),xx1);
y17 = spline(OxygenConversion(3,:),TCIII(3,:),xx1);
plot(xx1,y15./(10^6),xx1,y16./(10^6),xx1,y17./(10^6),'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('Total Capital Investment, [$MM]','FontSize',14,'FontName','Times New Roman');
legend('MR = 1/3','MR = 1','MR = 3');
% NPV vs Conversion
figure(16)
y18 = spline(OxygenConversion(1,:),NPVVV(1,:),xx1);
y19 = spline(OxygenConversion(2,:),NPVVV(2,:),xx1);

64. 64
y20 = spline(OxygenConversion(3,:),NPVVV(3,:),xx1);
plot(xx1,y18./(10^6),xx1,y19./(10^6),xx1,y20./(10^6),'Linewidth',1.5);
xlabel('Oxy Conversion','FontSize',14,'FontName','Times New Roman');
ylabel('NPV_P_r_o_j, [$MM]','FontSize',14,'FontName','Times New Roman');
legend('MR = 1/3','MR = 1','MR = 3');
axis([0 1 -600 400]);
% ROIBT vs Conversion
figure(17)
y21 = spline(OxygenConversion(1,:),ROIBTTT(1,:),xx1);
y22 = spline(OxygenConversion(2,:),ROIBTTT(2,:),xx1);
y23 = spline(OxygenConversion(3,:),ROIBTTT(3,:),xx1);
plot(xx1,y21,xx1,y22,xx1,y23);
xlabel('ROIBT');
ylabel('Conversion (O_2)');
legend('MR = 1/6','MR = 1/3','MR = 1','MR = 3');
% NPVper vs Conversion
figure(18)
y24 = spline(OxygenConversion(1,:),NPVPERRR(1,:),xx1);
y25 = spline(OxygenConversion(2,:),NPVPERRR(2,:),xx1);
y26 = spline(OxygenConversion(3,:),NPVPERRR(3,:),xx1);
plot(xx1,y24,xx1,y25,xx1,y26,'Linewidth',1.5);
xlabel('Conversion (O_2)','FontSize',14,'FontName','Times New Roman');
ylabel('NPV_%','FontSize',14,'FontName','Times New Roman');
legend('MR = 1/3','MR = 1','MR = 3');
axis([0 1 -25 25]);
% Sensitivity Analysis
NPVper_sens_m = [6.9 2.83 11.20];
NPVper_sens_o2 = [6.9 5.34 8.512];
NPVper_sens_co = [6.9 4.31 9.59];
NPVper_sens_dmc = [6.9 17.0 -3.18];
m_perc = [0.0 15 -15];
o2_perc = [0.0 15 -15];
co_perc = [0.0 30 -30];
dmc_perc = [0.0 15 -15];
% Price Fluctuation vs NPV%
figure(1)
xx2 = -40:1:40;
ysensdmc = spline(dmc_perc,NPVper_sens_dmc,xx2);
ysensm = spline(m_perc,NPVper_sens_m,xx2);
ysenso2 = spline(o2_perc,NPVper_sens_o2,xx2);
ysensco = spline(co_perc,NPVper_sens_co,xx2);
plot(xx2,ysensdmc,xx2,ysensm,'Linewidth',1.5);
xlabel('Percent Change in Price, [%]','FontSize',14,'FontName','Times New Roman');
ylabel('NPV_%','FontSize',14,'FontName','Times New Roman');
legend('Dimethyl Carbonate','Methanol');
axis([-40 40 -10 20]);
grid on
NPVer = [6.9 9.96 13.14 2.53 1.5065];
ERR = [12 5 0 35 50];
NPVfr = [6.9 2.41 -1.82];

65. 65
FRR = [4 20 35];
NPVtr = [6.9 8.05 5.17];
TRR = [25 10 48];
% NPV% vs Enterprise Rate
figure(2)
xx3 = 0:1:50;
yer = spline(ERR,NPVer,xx3);
plot(xx3,yer,'Linewidth',1.5);
xlabel('Enterprise Rate, [%]','FontSize',14,'FontName','Times New Roman');
ylabel('NPV_%','FontSize',14,'FontName','Times New Roman');
% NPV% vs Finance Rate
figure(3)
yfr = spline(FRR,NPVfr,xx3);
plot(xx3,yfr,'Linewidth',1.5);
xlabel('Finance Rate, [%]','FontSize',14,'FontName','Times New Roman');
ylabel('NPV_%','FontSize',14,'FontName','Times New Roman');
% NPV% vs Tax Rate
figure(4)
ytr = spline(TRR,NPVtr,xx3);
plot(xx3,ytr,'Linewidth',1.5);
xlabel('Tax Rate, [%]','FontSize',14,'FontName','Times New Roman');
ylabel('NPV_%','FontSize',14,'FontName','Times New Roman');

66. 66
Team
Member
Work
Statement
My
Contributions
to
this
report
were:
- CSTR
design
equations
and
Matlab
Code
- Costing
of
all
major
pieces
of
equipment.
- Economic
Analysis
- Generation
of
graphs
of
plant
of
economic
parameters
as
a
function
of
key
design
variables.
- Exec.
Summary/Production
Chemistry/Sensitivity
Analysi
Print
Name
and
Sign:
_________________________________
Date:
_________
Agreed:
_________________________________
Date:
_________
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Name
and
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Agreed:
_________________________________
Date:
_________
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Name
and
Sign