G10 Final

1
BICC Dimethyl Carbonate Plant Design and
Profitability Analysis
Charmaine Bennett
Sarah Maxel
Ye Yuan
Team #10
Date Submitted: June 3rd
, 2015
Executive Summary
Using phosgene to produce polycarbonate is dangerous and environmentally irresponsible. BICC
is one of the world’s leading manufacturers of polycarbonate resins, and our business continues
to grow by 10% per year; however, our current processes rely on phosgene to provide the
carbonate group during the synthesis of polycarbonate. A much less harmful chemical that can
also supply the carbonate group is dimethyl carbonate (DMC). To aid BICC in its transition to
greener chemical processes without losing its position in the global polycarbonate market, we
propose that it would be profitable to construct a plant capable of producing 150 MM kg of
DMC per year. DMC would be produced via the oxidation of carbon monoxide with oxygen and
methanol catalyzed by CuCl. The reaction occurs in a continuously stirred slurry reactor with a
liquid volume of 1.1 m3
, an operating temperature of 130℃, and a constant total pressure of 30
atm. The profitability analysis indicates that the total capitalized investment (TCI) required to
finance this undertaking will be $40 MM, with an expected net present value (NPVproj) of $75
MM. Given a positive NPVproj and a rate of return before taxes (ROIBT) of 53%, we conclude
that this DMC plant will be a profitable investment. The calculated the NPV% is 18.6%, and the
internal rate of return (IRR) is 37%. This NPV% calculated is well above the minimum 10%
annual return required to justify an investment, and the high IRR indicates an ideal business
venture.

2
Table of Contents
1. Introduction:..........................................................................................................................................3
2. Plant Design:.........................................................................................................................................3
2.1 Design Overview: ...............................................................................................................................3
2.2 Reactor Design:...................................................................................................................................4
2.3 Process Flow Diagram ........................................................................................................................6
2.4 HYSYS Simulation Flow-sheet ..........................................................................................................7
2.4 Piping and Instrumentation Diagram ..................................................................................................8
2.5 Aspen HYSYS Simulation..................................................................................................................9
2.6 Separation System Design ................................................................................................................10
2.7 Plant Process Control........................................................................................................................11
3. Economic Analysis .............................................................................................................................13
4. Health, Safety & Environmental (HSE) Considerations.....................................................................15
4.1 Materials of Construction .................................................................................................................17
5. Process Alternatives............................................................................................................................17
6. Conclusions:........................................................................................................................................17
Appendix A. Additional Figures.................................................................................................................18
Appendix B. Reactor Modelling and S vs X analysis.................................................................................21
Appendix C. Distillation Design Summary ................................................................................................22
Appendix C-1. Aspen User Interface / Aspen HYSYS Comparison Table............................................22
Appendix C-2. Ternary Phase Diagram for First Distillation Column ...................................................24
Appendix C-3. Ternary Phase Diagram for Second Distillation Column...............................................25
Appendix D. Economic Analysis................................................................................................................26
Appendix D-1. Economic Calculations ..................................................................................................26
Appendix D-2. Economic Spreadsheet ...................................................................................................31
Appendix D-3. IRR Spreadsheet.............................................................................................................32
Appendix D-4. Breakeven Analysis........................................................................................................33
Appendix E. Chemical Properties...............................................................................................................34
Appendix F. Level 2 Molar Flow Rates......................................................................................................36
Appendix G. Level 3 Mole Balances..........................................................................................................39
Appendix H. MATLAB Script....................................................................................................................40
Appendix H-1. Reactor Conceptual design and optimization.................................................................40

3
1. Introduction:
As a versatile polymer product, polycarbonate resin is an important commodity chemical
in modern society. It is a vital raw material for [1]1many manufacturing industries including
electronic, household utilities, and automotive industries. Currently, BICC relies on highly toxic
phosgene to provide the carbonate building blocks necessary for the polycarbonate resin
production [3]. Besides the hazardous nature of the reactants, such chemical processes
undermine environmental health by producing unsafe byproducts such as HCl. As environmental
sustainability continues be a major concern in the chemical industry, it is a necessity for BICC to
pursue cleaner, safer, and economical routes to create carbonate monomers as it expands its
production [2]. The generation of polycarbonate resins via the synthesis of dimethyl carbonate
(DMC) is a clean technology with low-cost raw materials and high production capacity; this
route is also capable of meeting the demands of BICC. We believe this process will be a highly
beneficial venture to ensure BICC’s dominant position in the global polycarbonate market for the
next decade.
2. Plant Design:
2.1 Design Overview:
Dimethyl carbonate (DMC) will be produced via the oxidation of carbon monoxide (CO)
with oxygen (O2) and methanol (MeOH) in a gas-liquid-solid slurry reactor [Rxn. 1]. Due to its
large vapor liquid interfacial area, this reactor type is the ideal choice for a liquid phase reaction
with a solid catalyst and two predominately gas-phase reactants. Two reactions will be occurring
in the liquid phase of the reaction solution: the desired reaction, formation of DMC [Rxn. 1], and
an undesired reaction, formation of carbon dioxide (CO2) [Rxn. 2]. Due to the exothermic nature
of both reactions (ΔHrxn1 = -73 kcal/mol, ΔHrxn2 = -68 kcal/mol), the reactor will operate
isothermally to prevent run-away reactions. Cuprous chloride (CuCl) as well as small quantities
of propriety additives will serve as catalyst for these reactions. The catalyst is insoluble in the
reaction mixture, and must be filtered out and recycled back to the reactor for reuse. The catalyst
is susceptible to poising if exposed to large quantities of water; therefore, a large molar excess of
methanol is necessary to maintain low water concentrations in the reactor.
2𝐶𝐻3 𝑂𝐻 +
1
2
𝑂2 + 𝐶𝑂 → 𝐶𝐻3 − 𝑂 − (𝐶𝑂) − 𝑂 − 𝐶𝐻3 + 𝐻2 𝑂 [Rxn. 1]

4
1
2
𝑂2 + 𝐶𝑂 → 𝐶𝑂2 [Rxn. 2]
The reactor inlet stream includes a feed of MeOH, CO, and O2. The O2 fed into the
reactor is expected to reach 99% conversion while the unreacted components, CO and MeOH,
will be separated and recycled back into the reactor feed to conserve raw material and minimize
production costs. Separated DMC will leave the plant at 35 ℃ and 1 atm for storage; separated
water will be treated as a waste product, and a waste gas stream of CO and CO2 will be released
into the atmosphere.
A minimum selectivity of 14% is needed to ‘break-even’ between the revenues from
selling DMC and the cost of raw reactants [Appendix D-4]. We advise operating at 46%
selectivity where the maximum economic potential is approximately $40 MM assuming that the
prices of MeOH, CO, O2, and DMC remain at $0.49 /kg, $0.18 /kg, $0.38 /kg, and $0.90 /kg
respectively.
2.2 Reactor Design:
In the conceptual design, the gas-liquid-solid slurry reactor is modelled as a
continuously-stirred tank reactor (CSTR). Both reactions occur in the liquid phase, and the
amount of CO, O2, and CO2 in the liquid are approximated with Henry’s law. The reactor is
characterized by four design variables: operation temperature (Treact), operation pressure (P), inlet
MeOH to O2 molar ratio (MR1), and inlet CO to O2 molar ratio (MR2). Allowable operation
temperature and pressure ranges are 80-130℃ and 10-40 atm respectively. A minimum MR2 of
24 ensures the O2 concentration in vapor phase remains below the explosion limit of 4 mol%.
Rxn. 1 has a higher activation energy and lower pre-exponential factor compared to Rxn. 2;
while the heat of reaction for Rxn. 1 has a larger absolute value. This suggests that Rxn. 1 is
thermodynamically favored, and Rxn. 2 is kinetically favored. Therefore, in order to increase the
selectivity of Rxn. 1, one must increase the operation temperature. This prediction is
corroborated by our selectivity (S) vs conversion (X) analysis in MATLAB. It indicates that with
a given conversion, higher selectivities can be reached with a higher operating temperature.
Although P has minimal effect on the S vs X relationship, increased pressures can reduce
the reactor volume, which is economically favorable. The effect MR1 has on the S vs X
relationship is non-monotonic, while for MR2 the effect is minimal; however, large MR values
lead to high flow rates and over-sized equipment. Given these considerations, the reactor will

5
operate at 130℃ and a total pressure of 30 atm. As a compromise between high selectivity and
low flow rates, a MR1 of 15 is advised. Since MR2 has little effect on S vs X, it is chosen to be
26, a value above the lower safety limit. Furthermore, adiabatic temperature rise calculation
reveals that with a feed inlet temperature of 130℃, the outlet temperature can reach 450℃ [Fig. 2
Appendix A]. Therefore, an isothermal reactor is advised.
The S vs X relationship at various reactor conditions are below illustrated below. Since
pressure and MR2 have minimal impact on the S vs X relationship, they are illustrated as Figure
11 and 12 in Appendix A.
Figure 1: Selectivity (S) vs conversion (X) at P = 30
atm, MR1 = 15, MR2 = 26, and various temperatures.
For a certain conversion, the higher the temperature, the
higher the selectivity
Figure 2: S vs X at T = 130℃, P = 30 atm, MR2 = 26,
and various MR1. For a given conversion, the
selectivities increase with MR1 until MR1 ≈ 20. Then S
decreases with increasing MR1. At X=0.99 maximum S
occurs when 10 < MR1 < 20

6
2.3 Process Flow Diagram
6 A
E-100:
Reactor Feed
Preheater
1: Fresh Feed CO
14.6 MT/h
$22.0 MM/y
4
E-101:
Vap. Effluent &
Vapor Recycle
Interchanger
8
E-1200:
Coln. Feed
Preheater
12
T-1200:
Water Extraction
Distillation Column
7
14
15: Waste Water
3.5 MT/h
$1 780 /y
17
P-1200/1200X
E-1400:
Vap. Recovery
Cooler
19
QH -1200
5
2: Fresh Feed O2
14.6 MT/h
$22.2 MM/y
7
14
11
K-1000/
K-1000X
M
CW
A-100/A-100X
DMC Slurry Reactor
Isobaric & Isothermal
Total Volume: 1.80 m3
Liquid Volume: 1.08 m3
Temp.: 130 °C
Pressure: 30 atm
6 B
10 A
E-1300:
Col. 1300
Feed Preheater
T-1300:
High Pressure
DMC
Distillation Column
15
19
P-1300/1300XA-900:
Vap. Effluent
Flash DrumE-102:
Vap. Effluent
Cooler
9
10 B
23
25
32: Waste Gas
12.8 MT/h
27
27
28
E-103:
Liq. Effluent &
Vapor Recycle
Interchanger
13 A
CW
Sat. Steam
X-1200:
Coln. Feed
Deaerator
13 B
24
31
30
XL-1400
Vapor Recovery
System
E-1301:
Product DMC
& Waste Water
Interchanger
QC -1200
18
20 A
QC -1300
QH -1300
21
20 B
E-1302: DMC
Product Cooler
20 C: DMC 99.9 w/w%
17.7 MT/h
22: Liq.
Recycle
Sat. Steam
E-2000:
Recycle CO
Heater
29
Sat. Steam
26
16: Col. T-1200 Vent
3: Fresh Feed MeOH
14.6 MT/h
$53.0 MM/y
Figure 3: Detailed process flow sheet with the properties of key variables and process streams

7
2.4 HYSYS Simulation Flow-sheet
Figure 4: HYSYS simulation flow sheet with properties of key process streams

8
2.4 Piping and Instrumentation Diagram
6 A
E-100
1: Fresh Feed CO
4
3: Fresh Feed
MeOH
E-101
8
E-1200
12
T-1200
7
14
15: Waste Water
17
P-1200/1200X
QH -1200
V-8
5
2: Fresh Feed O2
7
V-5 14
11
K-1000/
K-1000X
M
CW
A-100/A-100X
DMC Slurry Reactor
6 B
10 A
E-1300
T-1300
15
19 P-1300/1300X
A-900
E-102
9
10 B
23
32:
Waste Gas
27
28
E-103
V-14
13 A
CW
Sat. Steam
X-1200
13 B
24
31
30
XL-1400
E-1301
QC -1200
20 A
QC -1300
QH -1300
21
20 B
E-1302
22
Sat.
Steam
E-2000
29
Sat. Steam
26
16: Purge
V-4
V-13
V-7
V-13
V-22
V-21
V-1
V-2
V-3
RC
1
FT
2
FT
1
RC
3
FT
3
AC
3
FC
4
AC
1
FC
5
FC
8
LC
8
FC
7
TT
7
FC
6
TT
6
TT
8
V-9
FC
9
LC
9
V-10
FC
10
TT
10
PT
4
V-11
FC
11
PT
11
FC
14
LC
14
V-12
FC
12
SC
12
PT
12
FC
22
AC
22
Sat.
Steam
TT
19
FC
23
V-18
FC
18
PT
18
FC
15
PC
15
V-16
V-26
FC
26
LT
26
TT
4
TT
5
V-6
V-15
V-19
FC
19
19
V-23
20 C:
DMC
99.9 w/w%
V-24
FC
24
TT
24
TT
22
E-1400
CW
V-17
FC
17
25
TT
17
FC
16
LC
16
V-20
FC
20
LC
20
PT
21
TT
21
FC
21
S-112
FC
13
TT
13
V-25
FC
25
PC
25
Figure 5: Piping and instrumentation diagram for DMC plant

9
2.5 Aspen HYSYS Simulation
Taking the non-idealities of the mixtures into account, the design variables obtained from
MATLAB calculations are fine tuned in Aspen HYSYS. Table 1 below compares the key design
parameters associated with the reactor generated in MATLAB to those produced in HYSYS
simulation.
Table 1: Comparison of variables calculated on MATLAB verses the values obtained from the
Aspen HYSYS simulation.
Design Parameter MATLAB HYSYS Units
Fresh Feed Flowrates
Mass Flowrate of MeOH 16.2 12.9 MT/h
Mass Flowrate of CO 8.9 14.6 MT/h
Mass Flowrate of O2 6.7 6.9 MT/h
Reactor Specifications
Reactor Type Slurry reactor, isothermal, isobaric
Reactor Volume (liquid) 1.0 1.1 m3
Temperature 130 130 ℃
Total Pressure 30 30 atm
Single Pass Conversion 0.91 0.99 -
Single Pass Selectivity 0.66 0.46 -
Reactor Inlet Conditions
Pressure 30 30 atm
Total Flowrate 260.5 312.2 MT/h
Molar Ratio (MeOH to O2) 15.0 15.3 -
Molar Ratio (CO to O2) 26.0 26.6 -
Reactor Outlet Conditions
Pressure 30 30 atm
Total Flowrate 260.5 312.2 MT/h
Mass Flowrate of MeOH 84.7 94.0 MT/h
Mass Flowrate of CO 142.2 149.7 MT/h
Mass Flowrate of O2 0.6 0.1 MT/h
Mass Flowrate of DMC 22.8 54.3 MT/h
Mass Flowrate of H2O 4.6 3.9 MT/h
Mass Flowrate of CO2 5.7 10.1 MT/h

10
2.6 Separation System Design
There are two streams exiting our reactor due to the two-phases present in the reactor.
The liquid stream is comprised of methanol, water and DMC in addition to dissolved carbon
oxides, and a vapor phase predominately comprised of carbon oxides (COx). The presence of
two minimum boiling azeotropes (water-DMC and methanol-DMC) in the reactor effluent
streams necessitated the use of pressure swing distillation to obtain DMC at the specified 99.8
w/w% purity [4]. The first column operates at 1 atm and removes the water, while the second
column yields DMC. High pressure flash drums and a package vapor recovery system were used
to extract and recycle CO to the reactor.
The major constraint regarding the separation of the liquid reactor outlet stream is the
presence of two low-boiling point azeotropes [5]. The liquid effluent is first saturated and then
fed to the first column in the distillation train. This column produces a waste water bottoms
stream, and a near azeoptropic methanol-DMC stream as the distillate at atmospheric pressure. It
should be noted that COx needed to be removed from the liquid reactor effluent using a short
separating column, prior to distillation. Represented by vessel X-1200, this deaerator produces a
saturated feed stream to the first column with negligible concentrations of COx. The vapor
effluent is directed to the vapor recovery system for further purification. X-1200 also assists in
the removal of residual oxygen, and thereby reduced the rate of corrosion in downstream pipes
and equipment.
A small vent stream leaving the first column purges the carbon oxides dissolved in the
stream and is then directed to the vapor recovery system. Atmospheric pressure was used in the
initial column because higher column pressures result in unreasonably high separation
temperatures and operations costs. Pressure swing was utilized to facilitate the separation of the
MeOH/DMC azeotrope by crossing the distillation boundary. The distillate is pumped to the
second column at 11.5 atm, in which a high purity stream of DMC is produced as the bottoms
and a distillate near the corresponding azeotropic composition, at this pressure. This stream is
then pumped and recycled to the reactor to recover unreacted MeOH. The presence of DMC in
the recycle stream improves the ease of separation by shifting the compositions of subsequent
streams in the distillation system closer to the distillation boundary as seen on the residue curves.
The main goals of the vapor recovery system are to minimize of fresh feed of carbon
monoxide and purge CO2 from the plant. The CO and CO2 mixture is separated into a high purity

11
CO stream which is recycled to the reactor, a waste gas stream of carbon dioxide, and a liquid
stream which merges with the distillate of the first column. A possible design for the vapor
recovery system consists of an AMDEA absorber, and a heated flash for CO2 removal [6].
2.7 Plant Process Control
To ensure the plant operates within the design specifications and maintain safety, control
systems are implemented throughout the process. Such key control variables includes the molar
ratios of MeOH to O2, and CO to O2, as they affect the reactor efficiency by altering the vapor to
liquid ratio. In addition, in order to avoid explosion, control systems must ensure the mole
fraction of O2 remains below 4% in all vapor streams. Furthermore, pressure relief valves should
be installed to prevent over-pressurization. The control system must also meet the conversion
and the productions rate constraints despite the small fluctuations in the flow rate an purities of
the reactor feeds. Therefore, the nominal reactor pressure and temperature are specified and
maintained. Constraints associated with the process are tabulated below.
Table 2: Process control constraints
1 O2 mole fraction in vapor phase must be less than or equal to 4%
2 MeOH/O2 molar ratio at the reactor inlet must be equal to 15, within 3% tolerance
3 CO/O2 molar ratio at the reactor inlet must be equal to 26, within 3% tolerance
4 Mass flow rate of DMC in the product stream must be equal to 150MM kg/yr, within 3%
tolerance
5 Mass fraction of DMC in the product stream must be no less than 99.8 w/w%
6 Temperature at the reactor inlet must be equal to 130 ℃
7 Pressure into and out of the reactor must be equal to 30 atm
8 All streams entering pumps must be pure liquid
9 All streams entering compressors must be pure vapor
10 Liquid level in reactor must be maintained above operational threshold and below flooding
limit
11 Liquid levels in both distillation columns must stay below the flooding limit
Controlled and manipulated variables and their corresponding mechanismes for plant-
wide process control are tabulated below.

12
Table 3: Proposed control system scheme and associated control loops for the DMC plant.
Loop
Number
Controller
Type
Controlled Variable
CV
Manipulated Variable
MV
Valve
1 Cascade
Molar Ratio of CO/O2
MR 1
Flowrate of Fresh CO
V1
Feedback Composition of Reactor Feed Molar Ratio of CO/O2 V1
2 Feedback
MR 1
Flowrate of Fresh O2 V2
3 Cascade
Molar Ratio of MeOH/O2
MR 2
Flowrate of Fresh MeOH V3
Feedback Composition of Reactor Feed
Molar Ratio of MeOH/O2
MR 2
V3
4 Feedback Temp. of Reactor Feed
Flowrate of saturated steam to E-100 (Reactor
Feed Preheater)
V4
Feedback Reactor (A-100) P.
Flowrate of saturated steam to E-100 (Reactor
Feed Preheater)
V4
5 Feedback Reactor (A-100) Temp.
Flowrate of cooling water steam to A-100X
(Reactor Heat Exchanger)
V5
6 Feedback Temp. of Stream 27 (Vap. Recycle) Flowrate of Reactor Vap. Effluent V6
7 Feedback
Temp. of Stream 9 (A-900: Feed to Vap.
Effluent Flash Drum)
Flowrate of cooling water to E-102 (Reactor
Vap. Effluent Cooler)
V7
8 Feedback Reactor Liq. level Flowrate of 6 A: Liq. Reactor Effluent V8
Feedback Temp. of Stream 6 B Flowrate of 6 A: Liq. Reactor Effluent V8
9 Feedback Reactor Liq. level
Flowrate of 10 A: Liq. Effluent of A-900: Vap.
Effluent Flash Drum
V9
10 Feedback
Temp. of Stream 12 (Feed to Pre-
Distillation Deaerator)
Flowrate of saturated steam to E-1200 (Column
T-1200 Feed Preheater)
V10
11 Feedback
P. of A-900: Feed to Vap. Effluent Flash
Drum
Vap. flowrate exiting A-900 V11
12 Feedback
P. of Stream 23
(Recovered Vap. from Deaerator)
Flowrate of stream exiting compressor K-1000
Speed of K-1000X (Compressor Motor)
V12
13 Feedback Temp. of Recycle CO
Flowrate of saturated steam to E-2000 (Recycle
CO Heater)
V13
14 Feedback
Liq. level in X-1200: Pre-distillation
Deaerator
Flowrate of 13 A: Liq. Effluent of X-1200 V14
15 Feedback
P. of Stage 1 in T-1200: Water Extraction
Distillation Coln.
Flowrate of cooling water to T-1200 Condenser V15
16 Feedback Liq. level in T-1200 Condenser
Flowrate of 13 A: Liq. Effluent of T-1200
Condenser
V16
17 Feedback
Temp. of Feed to XL-1400: Vap.
Recovery system
Flowrate of cooling water to E-1400: Vap.
Recovery Pre-Cooler
V17
18 Feedback P. of Stream 18: Feed to Coln. T-1300 Flowrate of Stream 19: Feed to Coln. T-1300 V18
19 Feedback
Temp. of Stream 19 (Feed to Feed to
Coln. T-1300)
Flowrate of saturated steam to E-1300 (Column
T-1300 Feed Preheater)
V19
20 Feedback Liq. level in T-1300 Condenser Flowrate of 13 A: Liq. Reflux of T-1300 V20
21 Feedback
Temp. & P. of Stage 1 in T-1300: HP
DMC Distillation Coln
Flowrate of cooling water to T-1300 Condenser V21
22 Cascade
Temp. of Stream 20 A: Bottoms Stream
of Coln. T-1300
Flowrate of saturated steam to Column T-1300
reboiler
V22
Composition of 20 A: Bottoms Stream of
Coln. T-1300
Flowrate of saturated steam to Column T-1300
reboiler
V22
23 Cascade Flowrate of Product DMC
& of MeOH/ O2
V23
24 Feedback
Temp. of Waste Gas Stream to the
environment
Flowrate of Vap. exiting XL-1400 Vaoor
recovery system
V24
25 Feedback P. of Stream 22: Liq. Recycle to Reactor Flowrate of Liq. exiting P-1300 V25
26 Feedback Liq. level in last tray in T-1200 Flowrate of T-1200 bottoms stream V26

13
3. Economic Analysis
A cost analysis was conducted to determine the profitability and economic feasibility of
our conceptual design. For the current reactor conditions, conversion, and selectivity, our annual
profit before taxes (PBT) is $21 MM. The revenue generated from producing DMC is $135 MM
per year [Appendix D-1]. Our plant has an economically feasible ROIBT of 53% and our NPV% is
18.6%. The total capital investment (TCI) required is $40 MM. The yearly cost of fresh feed
reactants, $95 MM, is our largest expense. The following tables detail the total investment and
the operating costs required to finance our DMC plant.
Table 4: Installed costs of each piece of equipment are listed.
Equipment Cost [$MM]
Reactor 0.015
Separations System 4.3
Compressors 4.0
Flash Drum 0.23
Heat Exchangers 0.64
Total Equipment Costs 13
Fixed Capital 21
Total Investment 39
Table 5: Operating costs of the DMC plant are listed.
Operating Costs Cost [$MM/year]
Separations Unit 7.8
Fresh Feed of Oxygen 22
Fresh Feed of Methanol 53
Fresh Feed of Carbon Monoxide 22
Steam for Heat Exchangers 0.77
Coolant for Heat Exchangers 0.04
Chilled Water for Heat Exchangers 4.2
Electricity for Compressors 0.8
Vapor Recovery 2.0
Waste Water 0.0018
Yearly Operating Cost 113
Working Capital 16
Several assumptions were made in order to produce the economic analysis for our design.
The total investment (TI) was calculated as the sum of the start-up capital (SU), the working
capital (WC) and the fixed capital (FC). The start-up (SU) cost was estimated as 10% of FC. The
working capital (WC) was the cost of two months’ worth of raw feed materials (oxygen,
methanol, and carbon monoxide). The fixed capital investment (FC) is normally calculated as the
sum of the direct and indirect costs; for our analysis the direct cost was calculated as the sum of

14
the ISBL and the OSBL and the indirect cost was estimated as 30% of the direct costs. Both the
ISBL and the OSBL were estimated with a contingency of 25% and the OSBL was estimated as
40% more than the ISBL costs. ISBL was determined by summing the installed cost of each
piece of equipment [Appendix D-1]. These simplifications allowed us to calculate TI with the
following expression:
𝑇𝐼 = 2.5 × 𝐼𝑆𝐵𝐿 + 𝑊𝐶 (1)
The remainder of our profitability parameters was calculated using a Discounted Cash Flow
Analysis, which is detailed in Appendix D-2.
Aside from the cost of fresh feeds, the largest expense for both operating and installment
costs is the separations units; therefore minimizing costs associated with the distillation columns
is critical to maximizing profit. Note that rough estimations of profitability (Figure 6) from
HYSYS at lower pressures suggest that 11 atm is not the optimal pressure; a more detailed
analysis is needed to balance the various factors in determining optimal conditions.
Figure 6: Operating pressure in the DMC distillation column has a significant impact on profitability. Above
pressures of 11.5 atm in the second column, all economic parameters (NPV%, ROIBT,%, NPVproject, and TCI) start to
indicate lower profit.
We recommend that lower operating pressures for the DMC distillation column be
explored in order to reduce operation and installment costs associated with the reboiler and
condenser. The available operation pressures are limited by the azeotropic composition of DMC
and MeOH at each respective pressure. However, due to a high recycle flow of DMC, the current
design is sufficiently far away from the distillation boundary.

15
Figure 7: Sensitivity analysis of the net present value
versus conversion. Tax rate varied from 25% to 48%.
Figure 8: Sensitivity analysis of the net present value
versus conversion. Price of DMC varied to reflect
change in investment risk.
According to the sensitivity analysis illustrated in Figures 7 and 8, the tax rate has a
significant impact on the risk associated with this investment. However, even with a maximum
tax rate of 48%, this design is still profitable. As shown in Figure 8, a 10% drop in the selling
price of DMC would cause the NPV% to drop to approximately 10%, which is still economically
justifiable. This indicates the economic robustness of this design. If a competitor entered the
DMC market or the global demand of polycarbonate resins decreased, the profitability of our
plant would continue to be positive.
4. Health, Safety & Environmental (HSE) Considerations
One of the major process concerns of the proposed plant is the high flammability,
explosion hazards, and toxicity of the products in the plant. Dimethyl carbonate, carbon
monoxide, oxygen and methanol are extremely flammable and their can form explosive mixtures
with air. As a result, regular maintenance should be scheduled to checks for leaks, wear and
corrosion of the fittings in the plant. In addition, the exothermicity of the both reactions may give
rise to explosions in the reactor in the event of uncontrolled temperature increases. It is therefore
imperative for the cooling system for the reactor to be fully operational at all times, and
multiple/redundant control schemes are advised [6].
DMC and methanol are incompatible with oxidizing agents such as acids, chlorates,
nitrates and peroxides, therefore the ingress of oxygen into high purity streams of these species
should be avoided [6]. This hazard also required that concentrations of oxygen in all streams

16
across the plant are kept below 4 mol%. Oxygen was also selected as the limiting reagent to
restrict its presence to the front-end of the plant, and to minimize its concentrations downstream
the reactor. The presence of dissolved oxygen significantly increases the rate of corrosion which
would lead to increased maintenance costs and specialty materials needed for all pipes and
vessels. It is necessary that the CO to O2 ratio entering the reactor be greater than 24 to avoid the
formation of explosive mixtures. Due to the high risks to employee and environmental safety
within and around the plant and BICC’s industrial complex, tight control loops must be put in
place to stop the process if the molar ratio drops below the minimum for safe operation.
The catalyst used for the desired reaction chemistry is comprised of insoluble, finely-
dispersed cuprous chloride solid particle with proprietary additives in minute concentrations. The
chloride species generated in the process has the potentially to rapidly increase the rate of
corrosion within the plant, exacerbated but the presence of water [6]. While losses in catalytic
activity are inevitable over the lifetime of the plant, significant fouling will require expensive
and lengthy solutions to continue operations. As a result, the concentration of liquid water
entering the reactor was minimized using methanol. However, a compromise was reached
between the liquid molar ratio and the flammability of methanol. In the event of catalyst wetting,
and the activity has declined to critical levels, the plant must be shut down to replace the catalyst.
Solid particle traps and filters can be used to minimize catalyst lost and wear via friction. With
respect to the temperature profile in the reactor, the liquid and vapor inlet streams run counter
currently from the top and bottom of the vessel to minimize stagnant volumes and recirculation
eddies that would promote the formation of hot-zones.
With regards to the health of the employees at the plant, stringent precautions should be
enforced as many of the chemicals on the plant are either highly toxic or flammable. The storage
areas should be well ventilated, and the tanks should be grounded and fitted with deluge systems.
Methanol, oxygen and carbon monoxide leaks/ releases should require immediate
evacuation due to the asphyxiant hazard [7]. All work areas designated to have recognized
chemical carcinogens should have appropriate hazard warnings, and higher levels of personal
protective equipment should be enforced. These areas include the areas around the storage tanks
of the raw material and final products, and around the associated piping. In addition, the waste
water should be considered a hazardous material.

17
4.1 Materials of Construction
Stainless steels and hastelloy are used in almost all areas of the plant, with the exception
of the fresh slurry reactor [9]. The corrosiveness of liquids and solid catalyst particles would
cause stainless steel pipes to corrode at a rate of < 500 𝜇m per year [Appendix E]. Due to the
criticality of this vessel to plant operations, a molybdenum alloy should be used, as it provides
greater protections against rate corrosion and because it greater thermal stability at the reactor
conditions [7] [8].
5. Process Alternatives
Theoretically, a cascade of multiple CSTRs behaves like a plug flow reactor, and should
provide higher conversion with smaller volume. As an alternative process, the reactor was
modelled a series of 3 CSTRs with identical volumes. It was observed that with the same total
volume, three separate CSTRs produced nearly identical S vs X relationships as a single CSTR.
Despite the fact that the three CSTRs in series produced higher conversion than that from a
single CSTR at a given residence time, this advantage quickly diminishes above 20s of residence
time. This advantage can hardly be justified with the increased cost associated with additional
reactors. For graphical comparisons of the designs, please see Figure 9 and 10 in Appendix A.
To break the binary azeotropes between methanol and DMC in our system, pressure
swing distillation was utilized. However, an alternative separation design can incorporate
entrainers that alter the behavior of the azeotropic species. Suitable entrainers for this system
include aromatic hydroxy compounds, alkyl aryl ethers, and dialkyl carbonates [12].
6. Conclusions:
An ROIBT of 53% and an NPV of $75 MM indicates a profitable economic outlook for
our conceptual design. The NPV% of 18.6% is an acceptable value given the amount of risk
incurred for investing in a commodity chemical. A high IRR of 37% also suggests that DMC
production will be valuable investment. At this time we recommend a further assessment of the
design profitability because of its strong investment potential and green chemistry.

18
Appendix A. Additional Figures
Figure 1: Selectivity (S) versus reactor conversion (X)
at reactor temperature of 130 ℃, total pressure of 30
atm, MR1 of 15, MR2 of 26. Selectivity decreases with
increasing conversion, within a small window (65%<
S<71%)
Figure 2: Adiabatic temperature rise analysis: Reactor
inlet temperature (X-axis) versus reactor outlet
temperature (Y-axis). Due to the exothermic nature of
the reactions, adiabatic condition results in dangerously
high outlet temperature if the inlet temperature is 130 ℃
Figure 3: Flow rate of fresh O2, recycle flow rate of CO,
and total moles in and out of the reactor at reactor
temperature of 130 ℃, total pressure of 30 atm, MR1 of
15, MR2 of 26.
Figure 4: Mole fraction of each component entering the
separation system versus the reactor conversion at
reactor temperature of 130 ℃, total pressure of 30 atm,
MR1 of 15, MR2 of 26. (MeOH and CO are on the left
y-axis)
According to Figure 3, the total flow rate of CO in the recycle stream, as well as total flows out
of the react decrease slightly as conversion increases. The fresh feed rate of oxygen required to
produce the required amount of DMC increases as conversion increases. As seen in Figure 4, the
amount of DMC exiting the reactor maximizes at maximum conversion. Therefore, operating the
reactor at conversions very close to 100% is a reasonable choice.

19
Figure 5: Volume versus conversion at various reactor
temperatures. P=30atm, MR1=15, MR2=26
Figure 6: Volume versus conversion at various
pressures. T=130 ℃, MR1=15, MR2=26
Figure 7: Volume versus conversion at various MR1
values. T=130 ℃, P=30atm, MR2=26
Figure 8: Volume versus conversion at various MR2
values. T=130 ℃, P=30atm, MR1=15
Modelling the reactor as three equal volume CSTRs in series has indicates that Although multi-
CSTR design offers slightly faster approach to 100% conversion, and higher selectivity at given
conversion, the advantage is negligible at reasonably longer time scales
Figure 9: Conversion versus residence time for a single CSTR
and 3 CSTRs with identical volume in series. Higher conversion
at same residence time, the advantage is only visible at very
short time scales.
Figure 10: Selectivity versus conversion for the single CSTR
and multi-CSTR design.

20
Figure 11: S vs X at T = 130℃, MR1 = 15, MR2 = 26,
and various pressures. Pressure has minimal effect on S
vs X.
Figure 12: S vs X at T = 130℃, P = 30 atm, MR1 = 15,
and various MR2. Molar ratio of CO to O2 at the inlet
has minimal effect on S vs X.

21
Appendix B. Reactor Modelling and S vs X analysis
Equations 1-4 below show the reaction rate and the corresponding rate constants:
𝑟1 = 𝑘1 𝐶 𝑀𝑒
2
𝐶 𝑂2
0.5
Eqn. 1
𝑟2 = 𝑘2 𝐶 𝑂2
0.5
Eqn. 2
𝑘1 = 1.4 × 1011
exp[−
24000𝑐𝑎𝑙/𝑔𝑚𝑜𝑙𝑒
𝑅𝑇
] Eqn. 3
𝑘2 = 5.6 × 1012
exp[−
22700𝑐𝑎𝑙/𝑔𝑚𝑜𝑙𝑒
𝑅𝑇
] Eqn. 4
Where 𝑟1 and 𝑟2 are reaction rates for Rxn. 1 and Rxn. 2, respectively. The unites for rates are mol/(L s)
The reactor S vs X analysis is performed with the following algorithm:
1. Fix reactor conditions: T, Ptot, MR1, and MR2
2. Evaluate rate constants k1 and k2 at the selected T
3. Fix arbitrary O2 flow rate at the inlet at n0
4. Calculate the initial flow rates of MeOH, CO, and O2 based on MR1 and MR2
5. Pick residence time 𝜏
6. Solve reactor design equation a-f:
a. 𝐹 𝑀𝑒𝑂𝐻 = 𝐹 𝑀𝑒𝑂𝐻0 − 2𝑞𝜏𝑘1 (
𝜌𝐹 𝑀𝑒𝑂𝐻
𝐹 𝑀𝑒𝑂𝐻+𝐹 𝐷𝑀𝐶+𝐹 𝐻2𝑂
)
2
(
𝑃𝑡𝑜𝑡 𝜌𝐹 𝑂2
𝐾 𝐻𝑂2(𝐹 𝑂2+𝐹𝐶𝑂+𝐹𝐶𝑂2)
)
1
2
b. 𝐹𝐶𝑂 = 𝐹𝐶𝑂0 − 𝑞𝜏(𝑘1 (
)
2
(
)
1
2
+ 𝑘2 (
)
1
2
)
c. 𝐹𝑂2 = 𝐹𝑂20 −
1
2
𝑞𝜏(𝑘1 (
)
2
(
)
1
2
+ 𝑘2 (
)
1
2
)
d. 𝐹 𝐷𝑀𝐶 = 𝑞𝜏𝑘1 (
)
2
(
)
1
2
e. 𝐹 𝐻2𝑂 = 𝑞𝜏𝑘1 (
)
2
(
)
1
2
f. 𝐹𝐶𝑂2 = 𝑞𝜏𝑘2 (
)
1
2
Where 𝜌 =
𝜌 𝑀𝑒𝑂𝐻 𝐹 𝑀𝑒𝑂𝐻+𝜌 𝐷𝑀𝐶 𝐹 𝐷𝑀𝐶+𝜌 𝐻2𝑂 𝐹 𝐻2𝑂
and 𝜌𝑖 are the molar densities of the liquid species (mol/L);
𝑞 is the volumetric flow rate of the liquid species, and 𝑞 =
𝜌
;
and 𝐾 𝐻𝑂2 is Henry’s constant for O2
7. Calculate selectivity (S) and conversion (X)
8. Calculate the liquid volumetric flow rate at the inlet via the following equation:
𝑞𝑖𝑛𝑙𝑒𝑡 =
𝑃𝐷𝑀𝐶
0.5(𝑆 × 𝑋)(𝑀𝑅1 + 1)𝜌
Where 𝑃𝐷𝑀𝐶 is the target DMC production rate in mol/s
9. Compute reactor volume:
𝑉 = 𝜏 𝑞𝑖𝑛𝑙𝑒𝑡
10. Increment residence time 𝜏 repeat from step 5 until X is sufficiently close to 1

22
Appendix C. Distillation Design Summary
Appendix C-1. Aspen User Interface / Aspen HYSYS Comparison Table
Table 1: Comparison of distillation column variables calculated from conceptual designs on
Aspen Plus with the values obtained from the Aspen HYSYS simulation.
Design Parameter Aspen Plus HYSYS Units
T-1200: Water Extraction Column
Number of Theoretical Stages 14.3 14 -
Reflux Ratio 1.0 1.0 -
Reboil Ratio 25.7 23.3 -
V/F -
Distillate Temperature 64.4 63.7 °C
Distillate Composition
- CO
- O2
- CO2
- H2O
- DMC
- MeOH
0.0
0.0
0.0
0.0
0.17
0.82
0.0
0.0
0.0
0.6
17.0
82.5
Mol %
Bottoms Temperature 88.6 99.1 °C
Bottoms Composition
- CO
- O2
- CO2
- H2O
- DMC
- MeOH
0.0
0.0
0.0
93.0
1.0
7.0
0.0
0.0
0.0
99.3
0.0
0.7
Mol %
Vent Temperature 63.7 °C
Vent Composition
- CO
- O2
- CO2
- H2O
- DMC
- MeOH
1.6
0.0
0.1
0.2
15.5
82.5
Mol %
Condenser Duty 0.25 GJ/h
Reboiler Duty 0.19 GJ/h

23
T-1300: High Pressure DMC Column
Number of Theoretical Stages 19 19 -
Reflux Ratio 1.4 1.4 -
Reboil Ratio 34.2 45.4 -
V/F 2.25 -
Distillate Temperature 142.1 140.8 °C
Distillate Composition
- CO
- O2
- CO2
- H2O
- DMC
- MeOH
0.0
0.0
0.0
0.6
12.0
87.4
0.0
0.0
0.0
0.6
12.1
87.3
Mol %
Bottoms Temperature 190.7 189.0 °C
Bottoms Composition
- CO
- O2
- CO2
- H2O
- DMC
- MeOH
0.0
0.0
0.0
0.03
99.8
0.1
0.0
0.0
0.0
0.0
100.0
0.0
Mol %
Condenser Duty 0.63 0.24 GJ/h
Reboiler Duty 0.54 0.23 GJ/h

24
Appendix C-2. Ternary Phase Diagram for First Distillation Column
Figure 1: Ternary phase diagram for the Water Extraction Distillation Column with distillation boundary and
distillation profiles.

25
Appendix C-3. Ternary Phase Diagram for Second Distillation Column
Figure 2: Ternary phase diagram for the High Pressure Distillation Column with distillation boundary and
distillation profiles.

26
Appendix D. Economic Analysis
Appendix D-1. Economic Calculations
Revenue (R) = $135 MM
Dimethyl Carbonate (DMC) Production:
150 𝑀𝑀
𝑘𝑔 𝑜𝑓 𝐷𝑀𝐶
𝑦𝑒𝑎𝑟
×
$0.90
𝑘𝑔 𝑜𝑓 𝐷𝑀𝐶
= $135 𝑀𝑀/𝑦𝑒𝑎𝑟
Cost (C) = $ 137 MM
Raw Oxygen Flow Rate:
7000 𝑘𝑔 𝑜𝑓𝑂2
ℎ𝑟
×
8400 ℎ𝑟
1 𝑦𝑒𝑎𝑟
×
$0.38
𝑘𝑔 𝑜𝑓 𝑂2
= $22 𝑀𝑀
Raw Carbon Monoxide Flow Rate:
1.4 × 104
𝑘𝑔 𝑜𝑓𝐶𝑂
ℎ𝑟
×
8400 ℎ𝑟
1 𝑦𝑒𝑎𝑟
×
$0.18
𝑘𝑔 𝑜𝑓 𝐶𝑂
= $22 𝑀𝑀
Raw Methanol Flow Rate:
1.3 × 104
𝑘𝑔 𝑜𝑓𝑂2
ℎ𝑟
×
8400 ℎ𝑟
1 𝑦𝑒𝑎𝑟
×
$0.49
𝑘𝑔 𝑜𝑓 𝑀𝑒𝑂𝐻
= $55 𝑀𝑀
𝑆𝑒𝑙𝑒𝑐𝑡𝑖𝑣𝑖𝑡𝑦 = 0.46
𝐶𝑜𝑛𝑣𝑒𝑟𝑠𝑖𝑜𝑛 = 98.9%
𝑀𝑅1 = 14.8
𝑀𝑅2 = 26.6
Total Coolant Flow Rate:

27
𝐹𝑐𝑜𝑜𝑙 =
𝑄 𝐶
𝑐 𝑝 × ∆𝑇𝑐
Table 1: Coolant flowrates calculated.
Cooler 𝑄 𝐶 [𝐾𝑊] 𝑐 𝑝 [
𝐾𝐽
𝑘𝑔℃
] ∆𝑇𝑐[℃] 𝐹𝑐𝑜𝑜𝑙 [
𝑘𝑔
ℎ𝑟
]
E-1302 1467 4.179 20 6.3 × 104
Condenser 1 6.7 × 104
4.179 20 2.8 × 106
Condenser 2 6.4 × 104
4.179 20 2.8 × 106
∑ 𝐹𝑐𝑜𝑜𝑙 𝑖
𝟕
𝒊
× 𝐶𝑜𝑜𝑙𝑎𝑛𝑡 𝑃𝑟𝑖𝑐𝑒 = 5.7 × 106
𝑘𝑔 𝑜𝑓 𝑐𝑜𝑜𝑙𝑎𝑛𝑡
𝑦𝑒𝑎𝑟
(
$0.08
1000 𝑘𝑔 𝑜𝑓 𝑐𝑜𝑜𝑙𝑎𝑛𝑡
)
= $3.8 𝑀𝑀/𝑦𝑒𝑎𝑟
Total Steam Flow Rate for Heat:
𝐹ℎ𝑒𝑎𝑡 =
𝑄 𝐻
∆𝐻𝑠𝑡𝑒𝑎𝑚
Table 2: Steam flow rates for heating calculated.
Heater 𝑄 𝐻 [𝐾𝑊] ∆𝐻𝑠𝑡𝑚 [
𝐾𝐽
𝑘𝑔
] 𝐹𝑠𝑡𝑒𝑎𝑚 [
𝑘𝑔
ℎ𝑟
]
E-100 2064 1960 3.8 × 103
E-1200 1.5 × 104
1960 2.8 × 104
E-2000 2690 1960 4.9 × 103
E-1400 522 1960 9.6 × 102
E-1300 9935 1960 1.8 × 104
Reactor 2.3 × 104
1960 4.2 × 104
Reboiler 1 5.2 × 104
1960 9.6 × 104
Reboiler 2 6.5 × 104
1960 1.2 × 105
∑ 𝐹ℎ𝑒𝑎𝑡 𝑖
𝟓
𝒊
× 𝑆𝑡𝑒𝑎𝑚 𝑃𝑟𝑖𝑐𝑒 = 3.1 × 105
𝑘𝑔 𝑜𝑓 𝑠𝑡𝑒𝑎𝑚
𝑦𝑒𝑎𝑟
(
$0.08
1000 𝑘𝑔 𝑜𝑓 𝑠𝑡𝑒𝑎𝑚
) = $5.8 𝑀𝑀/𝑦𝑒𝑎𝑟
Total Chilled Water:
ℎ𝑒𝑎𝑡 𝑙𝑜𝑎𝑑 = 1.0 × 108
𝑘𝐽
ℎ𝑟

28
1.0 × 108
𝑘𝐽
ℎ𝑟
×
8400 ℎ𝑟
8400
× (
$4.9
𝐺𝐽
) = $4.2
𝑀𝑀
𝑦𝑒𝑎𝑟
Waste Water:
3.5 × 103
𝑘𝑔 𝑜𝑓 𝑤𝑎𝑠𝑡𝑒
𝑦𝑒𝑎𝑟
×
$0.06
1000 𝑘𝑔 𝑜𝑓 𝑐𝑜𝑜𝑙𝑎𝑛𝑡
= $1.7 × 103
/𝑦𝑒𝑎𝑟
Profit Before Taxes = $ 21 MM
𝑃𝐵𝑇 = 𝑅 − 𝐶
Fixed Capital (FC) = $21 MM
𝐹𝐶 = 2.28 × 𝐼𝑆𝐵𝐿
ISBL(Installed Cost)= $9 MM
Installed Cost Separator Columns:
$ 9.2 × 105
Table 3: Values listed were used to cost distillation columns.
𝐶𝑜𝑙𝑢𝑚𝑛 𝑐0 [
𝑚
ℎ𝑟
] 𝑉 [
𝑘𝑚𝑜𝑙
ℎ
] 𝑁 𝐻0 [𝑚] 𝐷 𝑇 [𝑚] 𝐻 [𝑚] 𝐶𝑜𝑠𝑡 [$]
1 329 3500 14 0.46 6.6 13 1.9 × 105
2 329 300 19 0.46 4.1 18 1.4 × 105
𝐴𝑟𝑒𝑎 =
𝑀𝜐
√ 𝜌𝜄 𝜌 𝜈
1
𝜙 𝑓𝑙𝑜𝑜𝑑 𝑐0
(
𝐴
𝐴 𝑛
)
0.8
𝑉
𝐷 𝑇 = 2 (
𝐴
𝜋
)
0.5
𝐻 = 2𝑁 × 𝐻0 + 𝐻0
𝐼𝐶 = (
𝑀&𝑆
280
) 101.9 𝐷 𝑇
1.066
𝐻0.82
+ (
𝑀&𝑆
280
) 4.7 𝐷 𝑇
1.55
𝐻 𝐹𝑐
𝑀&𝑆 = 1600
𝑀𝜐 = 42
𝑔
𝑚𝑜𝑙

29
𝜌 𝜈 = 1.5 𝑘𝑔/𝑚3
𝜌𝜄 = 829 𝑘𝑔/𝑚3
𝜙 𝑓𝑙𝑜𝑜𝑑 = 0.6
𝐴
𝐴 𝑛
= 0.8
Heat Exchangers:
Area of heat exchangers determined with the following two equations:
𝐴 =
𝑄
𝑈∆𝑇𝑙𝑜𝑔
∆𝑇𝑙𝑜𝑔 =
(𝑇𝑖𝑛
ℎ𝑜𝑡
− 𝑇𝑜𝑢𝑡
𝑐𝑜𝑙𝑑
) − (𝑇𝑜𝑢𝑡
ℎ𝑜𝑡
𝑐𝑜𝑙𝑑
)
ln (
𝑇𝑖𝑛
ℎ𝑜𝑡
𝑐𝑜𝑙𝑑
𝑇𝑜𝑢𝑡
ℎ𝑜𝑡
𝑐𝑜𝑙𝑑)
Table 4: Values listed were used to calculate surface area and costs for heat exchangers.
𝐻𝑒𝑎𝑡
𝐸𝑥𝑐ℎ𝑎𝑛𝑔𝑒𝑟
𝑈 [
𝑘𝑊
𝑚2 𝐾
] 𝑄[𝑘𝑊] ∆𝑇𝑙𝑜𝑔[℃] 𝑈𝐴 [
𝑘𝐽
℃
] 𝐴 [𝑚2] 𝐶𝑜𝑠𝑡 [$]
1 E-100 0.82 2.0 × 103 81 - 31 1.8E+04
2 E-102 0.485 2.8 × 104 50 - 1163 1.8E+05
3 E-1200 0.82 1.6 × 104 60 - 316 8.0E+04
4 E-2000 0.53 2.7 × 103 60 - 85 3.4E+04
5 E-1400 0.11 5.2 × 102 159 - 30 1.7E+04
6 E-1300 0.82 9.9 × 103 86 - 140 4.7E+04
7 Reactor 0.53 2.3 × 104 64 - 681 1.3E+05
8 E-101 0.095 - - 28 292
9 E-103 0.1 - - 28 278
10 E-1301 0.39 - - 14 36
11 E-1302 0.8 2.0 × 103 32 - 58
1 Condenser 1 0.105 2.0 × 103 22 - 2800 1.6E+06
1 Reboiler 1 0.82 2.0 × 103 94 - 672 1.3E+05
2 Condenser 2 0.3 2.0 × 103 99 - 2170 1.1E +06
2 Reboiler 2 0.82 2.0 × 103 5 - 15670 1.0E+06
𝐹𝑚 = 1 𝑓𝑜𝑟 𝐶𝑎𝑟𝑏𝑜𝑛 𝑆𝑡𝑒𝑒𝑙
𝐹𝑝 = 0 𝑓𝑜𝑟 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒𝑠 𝑢𝑝 𝑡𝑜 150 𝑝𝑠𝑖𝑎
𝐹𝑑 = 1 𝑓𝑜𝑟 𝑓𝑖𝑥𝑒𝑑 − 𝑡𝑢𝑏𝑒 𝑠ℎ𝑒𝑒𝑡

30
Reactor=Pressure Vessel + Heat Exchanger (A-101-X): $ 0.136 MM
𝐼𝑛𝑠𝑡𝑎𝑙𝑙𝑎𝑡𝑒𝑑 𝐶𝑜𝑠𝑡 = 𝐼𝐶
Pressure Vessel:
𝐼𝐶 𝑝𝑟𝑒𝑠𝑠𝑢𝑟𝑒 𝑣𝑒𝑠𝑠𝑒𝑙 = (
1600
280
) (101.9 × 𝐷1.066)(𝐻0.82)(2.18 + 𝐹𝑐) = $4000
𝐹𝑐 = 𝐹𝑚 𝐹𝑝
𝐹𝑚 = 1
𝐹𝑝 = 2.5 𝑓𝑜𝑟 𝑃𝑟𝑒𝑠𝑠𝑢𝑟𝑒 𝑜𝑣𝑒𝑟 𝑡𝑜 1000 𝑝𝑠𝑖𝑎
𝑀&𝑆2015 = 1600
Table 5: Values listed were used to calculate surface area and costs for decanter and flash drum.
Pressure Vessel Residence
Time [h]
Flow Rate
[m3
/h]
Height [m] Diameter
[m]
Cost [$]
Flash 0.25 16 2 1.6 5400
Total Investment (TI) =$ 39 MM
𝑇𝐼 = 2.5 (𝐼𝑆𝐵𝐿) + 𝑊𝐶
Working Capital (WC) = $ 16 MM
Two Month’s Cost of Raw Oxygen Flow Rate:
$22
𝑀𝑀
𝑦𝑒𝑎𝑟
×
1 𝑦𝑒𝑎𝑟
~6 𝑚𝑜𝑛𝑡ℎ𝑠
= $3.7 𝑀𝑀
Two Month’s Cost of Raw Carbon Monoxide Flow Rate:
$22
𝑀𝑀
𝑦𝑒𝑎𝑟
×
1 𝑦𝑒𝑎𝑟
= $8.8 𝑀𝑀
Two Month’s Cost of Raw Methanol Flow Rate:

31
$22
𝑀𝑀
𝑦𝑒𝑎𝑟
×
1 𝑦𝑒𝑎𝑟
= $3.7 𝑀𝑀
Appendix D-2. Economic Spreadsheet

32
Appendix D-3. IRR Spreadsheet

33
Appendix D-4. Breakeven Analysis

34
Appendix E. Chemical Properties
Name
Chemical
Formula
Molecular
Weight
Physical
Properties
Boiling
Point (°C)
Freezing
Point
(°C)
Flash
Point
(°C)
Flammabiliy
Limits (in Air) /
Explosion Limits
AutoIgnition
Temperature (
°C)
Specific
Gravity
Price
Corrosiveness (µm per year)
& Materials of Manufacture
Dimethyl
Carbonate
(DMC)
OC(OCH3)2 90.1
Clear Liquid,
Mildly Sweet
Odor
89.9 4 16.7
Lower: 3.1% (V)
Lower
Flammability Limit
Temp. = 8.9 °C
Upper: 20.5%(V)
Upper
Flammability Limit
Temp. = 46.9 °C
458 1.066
$0.90
/kg
Stainless Steel: < 50 µm/y
Hasalloy: < 50 µm/y
Carbon
Dioxide
CO2 44.0
Colorless,
Odorless gas
-
Sublimes
-78.5 (at 1
atm)
-56.57 - - -
1.52
(Vap. at
21 °C)
-
Stainless Steel: < 50 µm/y
Hasalloy: < 50 µm/y
Carbon
Monoxide
CO2 28.0
Colorless,
Odorless gas
-191.45 -205 -
Lower: 12.5% (V)
Upper:74.0%(V)
605
1.25
(Liq.)
$0.18
/kg
Carbon Steel: < 50 µm/y
Stainless Steel: µm/y
Oxygen O2 32.0
Colorless,
Odorless gas
-183.0
(at 1 atm)
-218.8
(at 1 atm)
- - - 1.105
$0.38
/kg
Stainless Steel: >1.27
mm/yrHasalloy: <0.05 mm/yr
Copper (I)
Chloride
CuCl 99.0
White
powderSlightly
green from
oxidized
impurities
1490 429.85 -
Lower: 1.1%
(V)Upper: 7.1%(V)
- 4.14 -
Stainless Steel: Unsuitable
Hasalloy: < 500 µm/yr
Methanol CH3OH 32.1
Clear, colorless
liquid. Alcohol
odor
64.6 -98 11
Lower: 6.0 % (V)
Upper: 36.0 %(V)
385 0.79
$0.49/kg
Stainless Steel: >0.18
µm/yrHasalloy: < 50
µm/yExtremely corrosive to
aluminum and iron Corrosion
exacerbated by the presence of
Oxygen

35
Name Toxicology Special Precautions NFPA Rating
Dimethyl
Carbonate
(DMC)
Can cause eye and skin irritation upon contact
Inhalation of vapors can cause anesthetic effect leading to
death
May be irritating to digestive tract if ingested
Not a known carcinogen
Highly Flammable
Keep container in a well-ventilated place and away from sources of
ignition
Avoid extreme temperatures
Incompatible with strong oxidizers and acids
Flammability: 3
Health:
Reactivity: 1
Other:
Carbon
Dioxide
Can cause eye and skin irritation upon contact
Inhalation of vapors can cause anesthetic effect leading to
death
May be irritating to digestive tract if ingested
Not a known carcinogen
Ventillation should ensure oxygen concentration remains above
19.5% and C02 concentration does not exceed 5000 ppm
Flammability: 0
Health: 1
Reactivity: 0
Other: -
Carbon
Monoxide
Toxic
Asphyxiant
Moderate concentrations may cause headache, drowsiness,
dizziness, and unconsciousness
Prolonged exposure leads to death
Extremely Flammable Gas
May form explosive mixtures with air and oxidizing agents
Ventilate area or move containers to well-ventilated areas
If venting or leaking gas catches fire, do not extinguish flames.
Flammable vapors may spread from leak, creating an explosive
reignition hazard
Flammability: 4
Health: 3
Reactivity: 0
Other: -
Oxygen
Breathing >80% at atmospheric pressure may cause ,
cough, sore throat, chest pain and breathing
difficulty.Retinal damage may occur after exposure for
extended periods
May cause or intensify fire (oxidizer)Extremely flammable in the
presence of reducingmaterials, combustible materials and organic
materials.Protect from sunlight when ambient temp. exceeds
52°CStore in a well-ventilated area
Flammability: 0
Health: 0 (gas) 3 (liq.)
Reactivity: 0
Other:Oxidizer
Copper (I)
Chloride
Very hazardous in case of ingestion
Eye contact can result in corneal damage or blindness.
Skin contact can produce inflammation and blistering
Over-exposure can produce lung damage, choking,
unconsciousness or death
Store in a separate safety storage cabinet or room
Prevent ingress of water, ingestion, dust inhalation and contact with
eyes
Wear suitable protective clothing in case of insufficient ventilation
Flammability: 0
Health: 3
Reactivity: 0
Other: -
Methanol
Highly Toxic May cause blindness if ingested Toxic via
inhalation - May cause headache, convulsions, and
eventually death. May be absorbed through the skin in
harmful amounts. Irritating to body tissues.Prolonged and/or
repeated contact may cause defatting of the skin and
dermatitis.Chronic exposure may cause reproductive
disorders and teratogenic effects (i.e. may cause birth
defects in the child or halt pregnancy
Extremely Flammable Gas When heated to decomposition, emits
acrid smoke and irritating fumesMay form explosive mixtures with
air and oxidizing agentsVentilate area or move containers to well-
ventilated areasIf venting or leaking gas catches fire, do not
extinguish flames.Flammable vapors may spread from leak, creating
an explosive reignition hazard
Flammability: 3
Health: 1
Reactivity: 0
Other: -

36
Appendix F. Level 2 Molar Flow Rates
Figure 1: Molar flowrates entering and leaving the DMC plant.
Given Reactions:
2𝐶𝐻3 𝑂𝐻 +
1
2
𝑂2 + 𝐶𝑂 → 𝐶𝐻3 − 𝑂 − (𝐶𝑂) − 𝑂 − 𝐶𝐻3 + 𝐻2 𝑂 (1)
1
2
𝑂2 + 𝐶𝑂 → 𝐶𝑂2 (2)
Initial mole balance matrix:
𝐹𝐷𝑀𝐶 𝑃𝐷𝑀𝐶 1 0 0
𝐹𝐶𝑂2
𝑃𝐶𝑂2 0 1 0
𝐹𝑂2
- 𝑃𝑂2
+ −1/2 −1/2 𝜉1 = 0
𝐹𝐶𝑂 𝑃𝐶𝑂 −1 −1 𝜉2 0
𝐹 𝑀𝑒𝑂𝐻 𝑃 𝑀𝑒𝑂𝐻 −2 0 0
𝐹 𝐻2 𝑂 𝑃 𝐻2 𝑂 1 0 0
𝐼𝑛𝑖𝑡𝑖𝑎𝑙 𝑀𝑜𝑙𝑒 𝐵𝑎𝑙𝑎𝑛𝑐𝑒𝑠:
𝐷𝑖𝑚𝑒𝑡ℎ𝑦𝑙 𝐶𝑎𝑟𝑏𝑜𝑛𝑎𝑡𝑒: 𝐹𝐷𝑀𝐶 − 𝑃𝐷𝑀𝐶 + 𝜉1 = 0
𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒: 𝐹𝐶𝑂2
− 𝑃𝐶𝑂2
+ 𝜉2 = 0
𝑂𝑥𝑦𝑔𝑒𝑛: 𝐹𝑂2
− 𝑃𝑂2
− 1/2 𝜉1 − 1/2 𝜉2 = 0
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝐹𝐶𝑂 − 𝑃𝐶𝑂 − 𝜉1 − 𝜉2 = 0
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝐹 𝑀𝑒𝑂𝐻 − 𝑃 𝑀𝑒𝑂𝐻 + 2𝜉1 = 0
𝑊𝑎𝑡𝑒𝑟: 𝐹 𝐻2 𝑂 − 𝑃 𝐻2 𝑂 + 𝜉1 = 0

37
Assuming none of the reactants leave the plant, flows set to zero where appropriate:
0 𝑃𝐷𝑀𝐶 1 0 0
0 𝑃𝐶𝑂2 0 1 0
𝐹𝑂2
- 0 + −1/2 −1/2 𝜉1 = 0
𝐹𝐶𝑂 0 −1 −1 𝜉2 0
𝐹 𝑀𝑒𝑂𝐻 0 −2 0 0
0 𝑃 𝐻2 𝑂 1 0 0
𝑆𝑢𝑏𝑠𝑒𝑞𝑢𝑒𝑛𝑡 𝑀𝑜𝑙𝑒 𝐵𝑎𝑙𝑎𝑛𝑐𝑒𝑠:
𝐷𝑖𝑚𝑒𝑡ℎ𝑦𝑙 𝐶𝑎𝑟𝑏𝑜𝑛𝑎𝑡𝑒: − 𝑃𝐷𝑀𝐶 + 𝜉1 = 0
𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒: −𝑃𝐶𝑂2
+ 𝜉2 = 0
− 1/2 𝜉1 − 1/2 𝜉2 = 0
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝐹𝐶𝑂 − 𝜉1 − 𝜉2 = 0
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝐹 𝑀𝑒𝑂𝐻 + 2𝜉1 = 0
𝑊𝑎𝑡𝑒𝑟: − 𝑃 𝐻2 𝑂 + 𝜉1 = 0
Reference species DMC and CO2 used to solve for extents:
𝜈 𝑟𝑒𝑓 = [
1 0
0 1
]
𝜈 𝑟𝑒𝑓
−1
= [
1 0
0 1
]
[
𝜉1
𝜉2
] = [
1 0
0 1
] [
𝑃𝐷𝑀𝐶
𝑃𝐶𝑂2
]
[
𝜉1
𝜉2
] = [
𝑃𝐷𝑀𝐶
𝑃𝐶𝑂2
]
0 𝑃𝐷𝑀𝐶 1 0 0
0 𝑃𝐶𝑂2 0 1 0
𝐹𝑂2
- 0 + −1/2 −1/2 𝑃𝐷𝑀𝐶 = 0
𝐹𝐶𝑂 0 −1 −1 𝑃𝐶𝑂2 0
𝐹 𝑀𝑒𝑂𝐻 0 −2 0 0
0 𝑃 𝐻2 𝑂 1 0 0
− 1/2 𝑃𝐷𝑀𝐶 − 1/2𝑃𝐶𝑂2
= 0

38
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝐹𝐶𝑂 − 𝑃𝐷𝑀𝐶 − 𝑃𝐶𝑂2
= 0
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝐹 𝑀𝑒𝑂𝐻 + 2𝑃𝐷𝑀𝐶 = 0
𝑊𝑎𝑡𝑒𝑟: − 𝑃 𝐻2 𝑂 + 𝑃𝐷𝑀𝐶 = 0
𝑆𝑝𝑒𝑐𝑖𝑓𝑖𝑐𝑎𝑡𝑖𝑜𝑛𝑠:
𝑃𝐷𝑀𝐶 = 150,000
𝑘𝑔
𝑦𝑟
𝑆𝑒𝑙𝑒𝑐𝑡𝑖𝑣𝑖𝑡𝑦 = 𝑆 =
𝑃𝐷𝑀𝐶
2 𝐹𝑂2
→ 𝐹𝑂2
=
𝑃𝐷𝑀𝐶
2 𝑆
𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒: 𝑃𝐶𝑂2
= 𝑃𝐷𝑀𝐶
(1 − 𝑆)
𝑆
=
𝑃𝐷𝑀𝐶
2 𝑆
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝐹𝐶𝑂 =
𝑃𝐷𝑀𝐶
𝑆
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝐹 𝑀𝑒𝑂𝐻 = 2𝑃𝐷𝑀𝐶
𝑊𝑎𝑡𝑒𝑟: 𝑃 𝐻2 𝑂 = 𝑃𝐷𝑀𝐶

39
Appendix G. Level 3 Mole Balances
Figure 2: Level 3 recycle material balance.
𝑆𝑝𝑒𝑐𝑖𝑓𝑖𝑐𝑎𝑡𝑖𝑜𝑛𝑠:
𝑃𝐷𝑀𝐶 = 150,000
𝑘𝑔
𝑦𝑟
𝑆𝑒𝑙𝑒𝑐𝑡𝑖𝑣𝑖𝑡𝑦 = 𝑆 =
𝑃𝐷𝑀𝐶
2 𝐹𝑂2
→ 𝐹𝑂2
=
𝑃𝐷𝑀𝐶
2 𝑆
𝐸𝑥𝑖𝑡 𝐹𝑙𝑜𝑤𝑠:
𝑃𝐷𝑀𝐶 = 150,000
𝑘𝑔
𝑦𝑟
𝑊𝑎𝑡𝑒𝑟: 𝑃 𝐻2 𝑂 = 𝑃𝐷𝑀𝐶
𝐶𝑎𝑟𝑏𝑜𝑛 𝐷𝑖𝑜𝑥𝑖𝑑𝑒: 𝑃𝐶𝑂2
= 𝑃𝐷𝑀𝐶
(1 − 𝑆)
𝑆
𝐹𝑟𝑒𝑠ℎ 𝐹𝑒𝑒𝑑:
𝑂𝑥𝑦𝑔𝑒𝑛: 𝐹𝐹,𝑂2
=
𝑃𝐷𝑀𝐶
2 𝑆
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝐹𝐹,𝑀𝑒𝑂𝐻 = 2𝑃𝐷𝑀𝐶
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝐹𝐹,𝐶𝑂 =
𝑃𝐷𝑀𝐶
𝑆
𝐹𝑒𝑒𝑑 𝑖𝑛𝑡𝑜 𝑅𝑒𝑎𝑐𝑡𝑜𝑟:
=
𝑃𝐷𝑀𝐶
2 𝑆𝑋
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝐹 𝑀𝑒𝑂𝐻 = 𝑀𝑅1 𝐹𝑂2
= 𝑀𝑅1
𝑃𝐷𝑀𝐶
2 𝑆𝑋
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝐹𝐶𝑂 = 𝑀𝑅2 𝐹𝑂2
= 𝑀𝑅2
𝑃𝐷𝑀𝐶
2 𝑆𝑋
𝑅𝑒𝑐𝑦𝑐𝑙𝑒 𝐹𝑙𝑜𝑤𝑠:
𝑅𝑒𝑐𝑦𝑙𝑒 𝐹𝑒𝑒𝑑 = 𝐹𝑒𝑒𝑑 𝑖𝑛𝑡𝑜 𝑅𝑒𝑎𝑐𝑡𝑜𝑟 − 𝐹𝑟𝑒𝑠ℎ 𝐹𝑒𝑒𝑑
𝑂𝑥𝑦𝑔𝑒𝑛: 𝑅 𝑂2
=
𝑃𝐷𝑀𝐶
2 𝑆
(
1 − 𝑋
𝑋
)
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙: 𝑅 𝑀𝑒𝑂𝐻 = 𝑀𝑅1
𝑃𝐷𝑀𝐶
2 𝑆
− 2𝑃𝐷𝑀𝐶 = 𝑃𝐷𝑀𝐶 (
𝑀𝑅1
2 𝑆𝑋
− 2)
𝐶𝑎𝑟𝑏𝑜𝑛 𝑀𝑜𝑛𝑜𝑥𝑖𝑑𝑒: 𝑅 𝐶𝑂 =
𝑃𝐷𝑀𝐶
𝑆
(
𝑀𝑅2
2𝑋
− 1)

40
Appendix H. MATLAB Script
Appendix H-1. Reactor Conceptual design and optimization
clc, clear, close all
%single condition run
Mme=32.04;%molar mass of methanol in g/mol
Mco=28.01;%molar mass of CO in g/mol
Mdmc=90.08;%molar mass of DMC in g/mol
Mh2o=18.01;%moalr mass of water in g/mol
Mo2=15.9994*2;%molar mass of O2 in g/mol
Mco2=44.01;%molar mass of CO2 in g/mol
Pdmc=150*10.^6*1000/Mdmc/8400/3600;%target flow rate of DMC in mol/s
Kho2=3179;%O2 Henry's const in bar
Khco=3107;%CO Henry's const in bar
Khco2=158;%CO2 Henry's const in bar
Ame=5.31301;%Antoine for methanol
Bme=1676.569;
Cme=-21.728;%Antoine range (80-210C)
Ah2o=4.6543;
Bh2o=1435.264;
Ch2o=-64.848;
Admc=4.77616;
Bdmc=1721.904;
Cdmc=-37.959;
Ptot0=30;%paressure in atm at inlet
R=1.9872041;%IG const in cal K-1 mol-1
RR=.082057;%IG const in L atm K-1 mol-1
tau=logspace(-3,2.5,60);
MR1=15;%MR between MeOH to O2
MR2=26;%MR between CO to O2 (Explosion Limit is less than 24)
for i=1:length(tau)
n0=58.3333;%inlet oxygen molar flow rate
Treact=130+273.15;%temperature in K
%rho=Ptot0/(1);%mol/L
Ptot0bar=Ptot0*1.01325;%total inlet pressure in bar
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));
xo20=(Ptot0bar/(MR2+1))/(Kho2);%initial O2 mol frac in liquid
xco0=(MR2*Ptot0bar)/(1+MR2)/(Khco);%initial CO mol frac in liquid
xme0=1-xo20-xco0;%initial MeOH mol frac in liquid
Ftot0liq=MR1*n0/xme0;%initial total molar flow rate in liq
Fme0=Ftot0liq*xme0;%initial MeOH molar flow in liq
Fo20=n0;%O2 initial flow rate in vap
Fco0=MR2*n0;%CO initial flow rate in vap
Fo20liq=Ftot0liq*xo20;%initial O2 molar flow in liq
Fco0liq=Ftot0liq*xco0;%initial CO molar flow in liq
Cme=22.2;%molar volume of MeOH in mol/L at 105C and 10-40 bar
rholiq=Cme;

41
Ch2o=52.5;%molar volume of water in mol/L
Cdmc=11.8;%molar volume of DMC in mol/L
q=Fme0/Cme;
%q=((Y(1)+Y(4)+Y(5))/(Cme*(Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4
)+Y(5)))))
%assume MeOH takes over the liquid phase property
%mol ratio times 22.2 mol/L gives me the concentrations, which is usable
%for the rate laws
time=tau(i);% Picking a tau [s]
a=[Fme0 Fo20 Fco0 0 0 0 Cme]; %initial condition for 9 equations to solve
error=1;
count=0;
while error>10^(-4);
CSTR_func=@(Y)([-Y(1)+Fme0-
((Y(1)+Y(4)+Y(5))/(Cme*(Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y
(5))))).*2.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(Y(1
)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^2.5);
-Y(2)+Fco0-
(5))))).*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(Y(1)/(
Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^2.5)-
(5))))).*time.*k2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
(5))))).*.5.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(Y(
1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^2.5)-
(5))))).*.5.*time.*k2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-
Y(4)+((Y(1)+Y(4)+Y(5))/(Cme*(Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y
(4)+Y(5))))).*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(
Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^2.5);
-
-
(4)+Y(5))))).*time.*k2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*(Cme*(Y(1)/(Y(1)+Y(4)+Y(5)))+Ch2o*
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
options=optimset('fsolve');
options.MaxFunEvals = 50000;
options.MaxIter = 2000;
options.TolFun=1e-28;
roots=real(fsolve(CSTR_func,a,options));%everything in result is in mol/s
result(i,:)=real(roots);
error=max(abs(a-result(i,:)));

42
a=result(i,:);%redifining initial guess
count=count+1;
result=abs(result);
end
molfrac(i,:)=[result(i,1)/sum(result(i,:));result(i,2)/sum(result(i,:));result(i,3)/sum(result(i,:));result(i,4)/sum(result(i,:
));result(i,5)/sum(result(i,:));result(i,6)/sum(result(i,:))];
conv=(Fo20-result(i,3))/Fo20;
sel=result(i,4)*.5/(Fo20-result(i,3));
flow_tot_inlet=Pdmc/(sel*conv)*.5*(1+MR1)/((result(i,1)/sum(result(i,1:6)))*Cme+(result(i,4)/sum(result(i,1:6)))*
Cdmc+(result(i,5)/sum(result(i,1:6)))*Ch2o);
volCSTR(i)=time*flow_tot_inlet;%CSTR liquid volume in L
convCSTR(i)=conv;
selCSTR(i)=sel;
Fresho2(i)=Pdmc/(sel)*.5;%fresh feed O2 mol/s
n0=Fresho2(i);
Freshco=n0*MR2*3600/10^6*2;
tot_out_mol(i)=sum(result(i,:));%total inlet mass flowrate
Rco(i)=-(Freshco-result(i,2));
tot_in_mol(i)=(1+MR1+MR2)*Fresho2(i)+Rco(i);
end
%mass_out=[result(i,1)*Mme*3600/10^(6);result(i,2)*Mco*3600/10^(6);result(i,3)*Mo2*3600/10^(6);result(i,4)*
Mdmc*3600/10^(6);result(i,5)*Mh2o*3600/10^(6);result(i,6)*Mco2*3600/10^(6)]%outlet mass flow rates in MT/h
Freshme=n0*selCSTR(i)*convCSTR(i)*Mme*3600/10^6*4;%fresh feed MeOH MT/h
tot_in_mass=(n0*Mo2+MR1*n0*Mme+n0*MR2*Mco)*3600/10^6;%total inlet mass flowrate
figure(1)
plot(tau,convCSTR)
axis([0,time,0,1]);
hold all
xlabel('Tau (sec)')
ylabel('X')
figure(2)
plot(convCSTR,selCSTR)
hold all
xlabel('X')
ylabel('S')
axis([0,1,0,1])
figure(3)
plot(convCSTR,volCSTR)
xlabel('X')
ylabel('Volume (m^3)')
%axis([0,1,0,2500]);
volCSTR(i)%design volume in L
result(i,4)*Mdmc/1000*3600*8400*10^(-6);%DMC flow rate in MMkg/yr
figure(4)
%plot(conv,Fresho2,conv,Reb,conv,Ftot_in,conv,Ftotal)

43
primary_axis=zeros(length(Fresho2),2);
primary_axis(:,1)=Rco;
primary_axis(:,2)=tot_in_mol;
primary_axis(:,3)=tot_out_mol;
secondary_axis=zeros(i,2);
secondary_axis(:,1)=Fresho2;
xlabel('X')
ylabel('Flow Rate (mol/s)')
[ax,h1,h2]=plotyy(convCSTR,primary_axis,convCSTR,secondary_axis);
set(ax(1),'YLim',[1000, 4000])
set(ax(1),'YTick',linspace(1000, 4000,5))
set(ax(2),'YLim',[35, 45])
set(ax(2),'YTick',linspace(35,45,5))
set(h2,'color','red');
set(h1(2),'LineStyle','--');
set(h1(3),'LineStyle','-.');
set(h1,'color','blue');
set(ax,{'ycolor'},{'blue';'red'});
ylabel('Flow Rate (mol/s)')
set(get(ax(2),'Ylabel'),'string','Flow Rate (mol/s)')
xlabel('Conversion')
legend('Recycle CO','Total_i_n','Total_o_u_t','Fresh O_2')
figure(5)
%plot(conv,Fresho2,conv,Reb,conv,Ftot_in,conv,Ftotal)
primary_axis1(:,1)=molfrac(:,1);
primary_axis1(:,2)=molfrac(:,2);
secondary_axis1(:,1)=molfrac(:,3);
xlabel('X')
ylabel('Mole Fraction')
[ax,h1,h2]=plotyy(convCSTR,primary_axis1,convCSTR,secondary_axis1);
set(ax(1),'YLim',[.3,.8])
set(ax(1),'YTick',linspace(.3,.8,5))
set(ax(2),'YLim',[0,.04])
set(ax(2),'YTick',linspace(0,.04,5))
set(ax,{'ycolor'},{'blue';'red'});

44
ylabel('Mole Fraction')
set(get(ax(2),'Ylabel'),'string','Flow Rate (mol/s)')
xlabel('Conversion')
legend('MeOH','CO','O_2','DMC/H_2O','CO_2')
Hfme=-57.02*1000;%heat of formation for liquid MeOH ethylbenzene in cal/mol
Hfco=-26.416*1000;%heat of formation for gas phase CO in cal/mol
Hfo2=0;%heat of formation for gas phase O2 in cal/mol
Hfdmc=-145.312*1000;%heat of formation for liquid DMC in cal/mol
Hfh2o=-68.317*1000;%heat of formation for liquid H2O in cal/mol
Hfco2=-94.052*1000;%heaet of formation for gas phase CO2 in cal/mol
Hrxn1=Hfdmc+Hfh2o-(2*Hfme+Hfco+.5*Hfo2);%heat of rxn1 in cal/mol
Hrxn2=Hfco2-(Hfco+.5*Hfo2);%heat of rxn2 in cal/mol
%Adiabatic Temp Rise
Tin=linspace(80,130,20);
Tout=zeros(1,length(Tin));
tau=logspace(-3,2,60);
for j=1:length(Tin)
for i=1:length(tau)
Treact=Tin(j)+273.15;%temperature in K
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));
rholiq=Cme;
q=Fme0/Cme;
error=1;
count=0;

45
-Y(2)+Fco0-
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
count=count+1;
result=abs(result);
end
flow_tot_inlet=Pdmc/(sel*conv)*.5*(1+MR1+MR2)/((result(i,1)/sum(result(i,1:6)))*Cme+(result(i,4)/sum(result(i,1
:6)))*Cdmc+(result(i,5)/sum(result(i,1:6)))*Ch2o);
volCSTR(i)=time*flow_tot_inlet/1000;%CSTR volume in m^3
convCSTR(i)=conv;
selCSTR(i)=sel;
Heat1=Hrxn1*roots(4);%heat generated from rxn1 in cal
Heat2=Hrxn2*roots(6);%heat generated from rxn2 in cal

46
end
Cpme80=22.685; %MeOH Cp at 80C cal/molK
Cpme130=27.046;%MeOH Cp at 130C cal/mol/K
cpme=@(tmp)((tmp-(273.15+80))*(Cpme130-Cpme80)/(130-80)+Cpme80);
Cpme=cpme(Treact);%heat capacity of methane in cal/mol K
Qr=Heat1+Heat2;
Cptot_out=Fme0*Cpme;
solveTout=@(To)((Treact-To)*Cptot_out-Qr);
To=fzero(solveTout,Treact)-273.15;
Tout(j)=To;
end
figure(6)
plot(Tin,Tout)
xlabel('T_i_n(C)')
ylabel('T_o_u_t(C)')
Temps=linspace(80,130,5);
%reactor optimization
n0=.5*Pdmc;%inlet oxygen molar flow rate
%Vary Temps
for j=1:length(Temps)
for i=1:length(tau)
Treact=Temps(j)+273.15;%temperature in K
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));

47
rholiq=Cme;
q=Fme0/Cme;
error=1;
count=0;
%{
q.*2.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^2.5);
-Y(2)+Fco0-
q.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^2.5)-
q.*time.*k2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^.5;
-Y(3)+Fo20-
q.*.5.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^2.5)-
q.*.5.*time.*k2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^.5;
-
Y(4)+q.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^2.5);
-
Y(5)+q.*time.*(k1.*(Y(1)./(Y(1)+Y(4)+Y(5))).^2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^2.5);
-Y(6)+q.*time.*k2.*(Y(3).*Ptot0bar./(Kho2.*(Y(2)+Y(3)+Y(6)))).^.5.*rholiq.^.5]);
%}
%below takes into account average variable liquid phase mole fractions
-Y(2)+Fco0-
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-

48
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
%xco_instant=Ptot0bar*Y(2)/(Y(2)+Y(3)+Y(6))/Khco
%xo2_instant=Ptot0bar*Y(3)/(Y(2)+Y(3)+Y(6))/Kho2
%xco2_instant=Ptot0bar*Y(6)/(Y(2)+Y(3)+Y(6))/Khco2
%Ftot_instant=(Y(1)+Y(4)+Y(5))/(1-Ptot0bar*(Y(2)/(Y(2)+Y(3)+Y(6))/Khco-
Y(3)/(Y(2)+Y(3)+Y(6))/Kho2-Y(6)/(Y(2)+Y(3)+Y(6))/Khco2))
count=count+1;
result=abs(result);
end
convCSTR(i)=conv;
selCSTR(i)=sel;
end
figure(7)
plot(tau,convCSTR)
axis([0,time,0,1]);
hold all
xlabel('Tau (sec)')
ylabel('X')
legend('T=80.0C','T=92.5C','T=105.0C','T=117.5C','T=130.0C')
figure(8)
hold all
xlabel('X')
ylabel('S')
axis([0,1,0,1])
figure(9)
hold all
xlabel('X')

49
%axis([0,1,0,2500]);
box on
end
%Vary Pressure
Pressures=linspace(10,40,5);
for j=1:length(Pressures)
for i=1:length(tau)
Ptot0bar=Pressures(j)*1.01325;%total inlet pressure in bar
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));
rholiq=Cme;
q=Fme0/Cme;
error=1;
count=0;
-Y(2)+Fco0-
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-

50
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
count=count+1;
result=abs(result);
end
convCSTR(i)=conv;
selCSTR(i)=sel;
end
figure(10)
plot(tau,convCSTR)
axis([0,time,0,1]);
hold all
xlabel('Tau (sec)')
ylabel('X')
legend('P=10.0atm','P=17.5atm','P=25.0atm','P=32.5atm','P=40.0atm')
figure(11)
hold all
xlabel('X')
ylabel('S')
axis([0,1,0,1])
figure(12)
hold all
xlabel('X')

51
%axis([0,1,0,250]);
box on
end
MR1s=[4,6,10,20,60];
%Vary MR1
for i=1:length(tau)
if j==5;
MR1=5;
else
MR1=MR1s(j);
end
Ptot0bar=30*1.01325;%total inlet pressure in bar
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));
rholiq=Cme;
q=Fme0/Cme;
)+Y(5)))))
%for the rate laws
error=1;
count=0;
-Y(2)+Fco0-

52
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
count=count+1;
result=abs(result);
end
convCSTR(i)=conv;
selCSTR(i)=sel;
end
figure(13)
plot(tau,convCSTR)
axis([0,time,0,1]);
hold all
xlabel('Tau (sec)')
ylabel('X')
legend('MR1=4','MR1=6','MR1=10','MR1=20','MR1=60')

53
figure(14)
hold all
xlabel('X')
ylabel('S')
axis([0,1,0,1])
figure(15)
hold all
xlabel('X')
% axis([0,1,0,250]);
box on
end
MR1=15;
MR2s=logspace(1.380211242,2,5);
%Vary MR2
for i=1:length(tau)
MR2=MR2s(j);
Ptot0bar=40*1.01325;%total inlet pressure in bar
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));
rholiq=Cme;
q=Fme0/Cme;
error=1;
count=0;

54
-Y(2)+Fco0-
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
count=count+1;
result=abs(result);
end
convCSTR(i)=conv;
selCSTR(i)=sel;
end
figure(16)
plot(tau,convCSTR)
axis([0,time,0,1]);
hold all

55
xlabel('Tau (sec)')
ylabel('X')
figure(17)
hold all
xlabel('X')
ylabel('S')
axis([0,1,0,1])
figure(18)
hold all
xlabel('X')
%axis([0,1,0,250]);
box on
end
%single vs multi CSTR
Mme=32.04;%molar mass of methanol in g/mol
Mco=28.01;%molar mass of CO in g/mol
Mdmc=90.08;%molar mass of DMC in g/mol
Mh2o=18.01;%moalr mass of water in g/mol
Mo2=15.9994*2;%molar mass of O2 in g/mol
Mco2=44.01;%molar mass of CO2 in g/mol
Pdmc=150*10.^6*1000/Mdmc/8400/3600;%target flow rate of DMC in mol/s
Kho2=3179;%O2 Henry's const in bar
Khco=3107;%CO Henry's const in bar
Khco2=158;%CO2 Henry's const in bar
Ame=5.31301;%Antoine for methanol
Bme=1676.569;
Cme=-21.728;%Antoine range (80-210C)
Ah2o=4.6543;
Bh2o=1435.264;
Ch2o=-64.848;
Admc=4.77616;
Bdmc=1721.904;
Cdmc=-37.959;
% Psatme=10.^(Ame-(Bme/(Treact+Cme)));%Inputs T in K and outputs Psat in bar
% Psath2o=10.^(Ah2o-(Bh2o/(Treact+Ch2o)));%Inputs T in K and outputs Psat in bar
% Psatdmc=10.^(Admc-(Bdmc/(Treact+Cdmc)));%Inputs T in K and outputs Psat in bar
tau=logspace(-3,1.0887,60);
n0=58.3333;%inlet oxygen molar flow rate

56
for i=1:length(tau)
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));
rholiq=Cme;
q=Fme0/Cme;
)+Y(5)))))
%for the rate laws
error=1;
count=0;
-Y(2)+Fco0-
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;

57
-
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);
count=count+1;
result=abs(result);
end
volCSTR(i)=time*flow_tot_inlet;%CSTR liquid volume in L
convCSTR(i)=conv;
selCSTR(i)=sel;
end
convCSTRsingle=convCSTR;
selCSTRsingle=selCSTR;
volCSTRsingle=volCSTR;
tausingle=tau;
Fo20_old=Fo20;
%dividing it into 3 CSTRs of identical residence time
% tau(1:20)=linspace(0.001,45,20);
% tau(21:40)=linspace(46,90,20);
% tau(41:60)=linspace(91,10^2.1306,20);
for i=1:54
k1=1.4*10.^(11)*exp(-24000/(R*Treact));
k2=5.6*10.^(12)*exp(-22700/(R*Treact));

58
rholiq=Cme;
q=Fme0/Cme;
error=1;
count=0;
%below takes into account average variable liquid phase mole fractions
-Y(2)+Fco0-
/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-Y(3)+Fo20-
Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5;
-
-
-
(Y(5)/(Y(1)+Y(4)+Y(5)))+Cdmc*(Y(4)/(Y(1)+Y(4)+Y(5)))).^.5]);

G10 Final

Recommended

Recommended

More Related Content

What's hot

What's hot (20)

Similar to G10 Final

Similar to G10 Final (20)

G10 Final