Simulation of Vinyl Acetate Production

Chemical & Process Engineering
Faculty of Engineering and Physical Sciences
Process Modelling and Simulation
(ENGM214)
Simulation of Vinyl Acetate Production
Coursework 2
Report Authors
Mehdi Aissani
Ogorchukwu Chimsunum
Ammar Grewal
Afnan Shareef
Date of Investigation
1st
December 2014 – 16th
December 2014
Date of Submission
16th
December 2014

1
Contents
List of Figures ........................................................................................................................................2
Introduction.............................................................................................................................................1
1.0 Mathematical Model of the System ............................................................................................3
1.1 Preheater .................................................................................................................................3
1.1.1 Mass Balance for Preheater.............................................................................................3
1.1.2 Energy Balance for Preheater..........................................................................................4
1.2 Reactor....................................................................................................................................6
1.2.1 Mass Balance for Reactor ...............................................................................................6
1.2.2 Energy Balance for Reactor ............................................................................................7
1.3 Separator .................................................................................................................................8
1.3.1 Mass Balance for Separator ............................................................................................8
1.3.2 Energy Balance for Separator .......................................................................................10
1.4 CO2 Separator.......................................................................................................................11
1.4.1 Mass Balance for CO2-Separator .................................................................................11
1.4.2 Energy Balance for CO2-Separator ..............................................................................13
1.5 Distillation Column...............................................................................................................14
1.5.1 Mass Balance for Distillation Column..........................................................................14
1.5.2 Energy Balance for Distillation Column.......................................................................15
1.6 Flash Drum............................................................................................................................17
1.6.1 Mass Balance for Flash Drum.......................................................................................17
1.6.2 Energy Balance for Flash Drum....................................................................................18
1.7 Decanter................................................................................................................................19
1.7.1 Mass Balance for Decanter ...........................................................................................19
1.7.2 Energy Balance for Decanter ........................................................................................21
1.8 Overall Process .....................................................................................................................23
1.8.1 Overall Process Material Balance.................................................................................23
1.8.2 Overall Process Energy Balance...................................................................................24
2.0 Simulating the System ..............................................................................................................24
2.1 Simulation Input file .............................................................................................................30
2.2 Simulation Report file...........................................................................................................35
2.3 Simulation Heat and Mass balance table ..............................................................................63
3.0 Sensitivity Analysis ..................................................................................................................64
4.0 Verification of Results ..............................................................................................................75

2
List of Figures
Figure 1: The mixer with the three main feed streams going in ...........................................................25
Figure 2: Reactor with pre-heated feed stream, and exit stream passing through cooler......................25
Figure 3: Reaction 1 inside the reactor, with 25% conversion of ethylene...........................................26
Figure 4: Reaction 2 inside the reactor, with 5% conversion of ethylene.............................................26
Figure 5: Liquid phase from Separator going in to the DSTWU..........................................................27
Figure 6: Results of the DSTWU..........................................................................................................27
Figure 7: Products streams from Flash Drum going in to Decanter and Absorber...............................28
Figure 8: Aspen Plus V8.4 simulation of a vinyl acetate monomer process.........................................29

1
Introduction
The production of Vinyl Acetate from ethylene is an 11-step process which utilises the process of a
stoichiometric reactor, distillation column, flash phase separators and a decanter. Firstly combining
their separate feeds in a mixer to create a mixture creates a mixture of oxygen, ethylene and acetic
acid. This mixture’s temperature is then raised to 148o
C in a preheater to facilitate the requirements
for the reaction process.
Ethylene, oxygen, and acetic acid, are converted into vinyl acetate. Water and carbon dioxide are
formed as resulting by-products. The reaction process of combining ethylene, oxygen, and acetic acid
to form vinyl acetate occurs in the stoichiometric reactor and is defined by the following chemical
process.
The reactions are irreversible and the reaction rates have an Arrhenius-type dependence on
temperature. The exit stream of the reactor then passes through a process-to-
process heat exchanger. During this process, the reactor effluent is cooled with
cooling water and the vapour, containing oxygen, ethylene, carbon dioxide; is separated from the
liquid containing vinyl acetate, water, acetic acid. From the heat exchanger the mixture passes through
a phase separator which separates compounds in gas phase from those in liquid phase. The vapour
stream containing carbon dioxide, oxygen and ethylene; flows from the separator to a compressor
separator where the carbon dioxide and unreacted ethylene are recycled to the beginning process. The
separator liquid efflux containing vinyl acetate, acetic acid and water enters a distillation column. The
azeotropic distillation column exploits the low boiling point of vinyl acetate and water; converting
their phase to a gas, where they exit the top of the column. The residing acetic acid exits the
distillation column via the bottom and is recycled into the process at the mixing stage.
The distillate then enters another separator where any residing CO2 oxygen and ethylene are removed
leaving only water and vinyl acetate. This final stream enters a decanter completing the final
separation of the process.
Material Balances calculations:
The general equation used for the material balances for the process is:
𝐴𝑐𝑐𝑢𝑚𝑢𝑙𝑎𝑡𝑖𝑜𝑛 = 𝑖𝑛 − 𝑜𝑢𝑡 + 𝑔𝑒𝑛𝑒𝑟𝑎𝑡𝑖𝑜𝑛 − 𝑎𝑐𝑐𝑢𝑚𝑢𝑙𝑎𝑡𝑖𝑜𝑛 (1)
Note: we assume the process is operating as steady state, therefore the accumulation is equal to zero.
The molar flowrates of the feeds for the process were given as kmol/day, this was converted to
perform the material balances.

2
Energy Balances calculations:
Equation (1) is the fundamental equation used for the energy balances, in order to perform the energy
balances specific heat capacities values were obtained for literature, for example to obtain the specific
heat capacities for Acetic acid and Vinyl acetate, the following correlation were used from Perry’s
chemical engineering handbook:
Acetic Acid:
𝐶𝑝 = 139600 − 320.8(𝑇) + 0.8985(𝑇2) =
𝐽
𝑚𝑜𝑙𝑒
. 𝑘 (2)
Vinyl Acetate:
𝐶𝑝 = 136306 − 106.17(𝑇) + 0.75175(𝑇2) =
𝐽
𝑚𝑜𝑙𝑒
. 𝑘 (3)
These values were converted to KJ/Kg.K using the molecular weight of each component, the Cp
values for oxygen, water, carbon dioxide and ethylene were obtained from engineering tool box
website using interpolation.

3
1.0 Mathematical Model of the System
1.1 Preheater
1.1.1 Mass Balance for Preheater
Material balance For the Preheater:
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
− 𝐹𝑜𝑢𝑡 ∗ 𝑋𝑖
𝑜𝑢𝑡
+ Ri = 0 (4)
Fin
= Flow rate of feed in (kmol/h)
Xi
in
= molar concentration of component in feed in
Fout
= Flow rate of feed out (kmol/h)
Xi
in
= molar concentration of component in feed out
Ri = Product made
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡 ∗ (𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
+ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐻2𝑂
𝑜𝑢𝑡
+ 𝑋 𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐸𝑇𝐻
𝑜𝑢𝑡
+ 𝑋𝐴𝐶𝐸
𝑜𝑢𝑡
+ 𝑅𝑖 (5)
𝐹 𝑖𝑛
= 𝐹 𝑜𝑢𝑡
= 𝐹
𝑅𝑖 = 0
The flow in is equal to the flow out as there is only one stream in to the pre-heater and one stream out
of the pre-heater.
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
𝑜𝑢𝑡
+ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐻2𝑂
𝑜𝑢𝑡
+ 𝑋 𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐸𝑇𝐻
𝑜𝑢𝑡
+ 𝑋𝐴𝐶𝐸
𝑜𝑢𝑡
) (6)
In the above equation we can see that the overall amount of material present passing through the
preheater is defined by the rate of the flow in to the preheater multiplied by the fraction of each
component present in the inlet feed. This fraction is calculated by dividing the molar flow of feed for
that component in by the molar flow of the total feed flow rate. The feed in is subject to the overall
molar flow rate exiting the preheater too. As no reaction is taking place thus no new products are
formed the amount of product in is equal to the amount of product leaving the preheater. Due to the
process only serving to raise the temperature of the feed to a certain value the parameters for the mass
of each component in the stream will remain the same and are given below.
Preheater
FEED REACT-IN

4
Table 1. Material balance around Preheater
Component Input(kmol/h) Output(kmol/h)
Acetic Acid 855.9280188 855.9280188
Ethylene 1747.621118 1747.621118
Oxygen 961.4858647 961.4858647
VAC 10.19419971 10.19419971
Water 68.1870525 68.1870525
Carbon Dioxide 1.246252032 1.246252032
Total 3644.662505 3644.662505
1.1.2 Energy Balance for Preheater
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
− 𝐹𝑜𝑢𝑡 ∗ 𝐶𝑃𝑥
𝑜𝑢𝑡
+ 𝑄 (7)
Fin
= Flow rate of feed in (Kmol/h)
CPx
in
= Specific heat capacity of component in feed in (Kg/Kj.K)
Fout
= Flow rate of feed out (Kmol/h)
CPx
out
= Specific heat capacity of component in feed out (Kg/Kj.K)
Q = Heat Required by process
𝐶𝑝𝑥 𝑚𝑖𝑥
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
+ 𝐶𝑝 𝐶𝑂2
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝐶𝑝 𝐸𝑇𝐻
𝑖𝑛
∗ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝐶𝑝 𝐴𝐶𝐸
𝑖𝑛
∗ 𝑋𝐴𝐶𝐸
𝑖𝑛
+ 𝐶𝑝 𝐻𝑂2
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
+ 𝐶𝑝 𝑉𝐴𝐶
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (8)
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (9)
𝑑𝑈
𝑑𝑡
= 𝐹 ∗ (𝐶𝑃𝑥,𝑚𝑖𝑥
𝑖𝑛
− 𝐶𝑃𝑥,𝑚𝑖𝑥
𝑜𝑢𝑡
) + 𝑄 (10)

5
𝐹 𝑖𝑛
= 𝐹 𝑜𝑢𝑡
= 𝐹
The flow in is equal to the flow out as there is only one stream in to the pre-heater and one stream out
of the pre-heater
The overall energy balance of the components passing through the preheater is defined by: the rate of
the flow in to the preheater multiplied by the specific heat capacity of the fraction of each component
present in the inlet feed. The feed in is subject to the overall specific heat capacity of the components
exiting the preheater as a mixture. As no reaction is taking place thus no new products are formed the
amount of product in is equal to the amount of product leaving the preheater. As the flow rate also
does not change, the flow rates of both feeds are equal and can be represented by one F variable as
mentioned above. The Q value is dependent on the product of the mass of each component, the
specific heat capacity and the temperature of the stream, thus defining the heat energy required by the
process. The table below show the results obtained for the streams in and out for the pre-heater.
Table 2 Energy balance around Preheater
UNIT BLOCK INPUT (KJ/H) OUTPUT (KJ/H)
Preheater 68,846,916 114,347,154
Note: The difference in the input and output energy of the system is the heat provided by the preheater
to raise the temperature of the feed to the required value. The enthalpy of each stream is calculated in
Kj/h and then converted to cal/sec using the following correlation:
0.0663919048222 × 𝑄 𝑘𝐽/ℎ = 𝑄𝑐𝑎𝑙/𝑠𝑒𝑐 (11)
This correlation was obtained from a power conversion table (www.aqua-calc.com)
Preheater

6
1.2 Reactor
1.2.1 Mass Balance for Reactor
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
𝑜𝑢𝑡
+ 𝑅𝑖 = 0 (11)
Fin
= Flow rate of feed in (Kmol/h)
Xi
in
Fout
= Flow rate of feed out (Kmol/h)
Xi
in
Ri = Reaction of ethylene
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
𝑜𝑢𝑡
+ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐻2𝑂
𝑜𝑢𝑡
+ 𝑋 𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐸𝑇𝐻
𝑜𝑢𝑡
+
𝑋𝐴𝐶𝐸
𝑜𝑢𝑡
) (12)
The reactor is the block process where we convert Ethylene to Vinyl Acetate in the presence of
Oxygen, CO2 and Acetic Acid. In the above equation we can see that the overall amount of material
present passing through the reactor is defined by the rate of the flow in to the preheater multiplied by
the mole fraction of each component present in the inlet feed. This fraction is calculated by dividing
the molar flow rate of the feed in for each component divided by the total molar flow rate for the
stream in or the stream out of the reactor. We assume that no accumulation occurs inside the reactor
as the reactor is operated at steady state. As the reactions that occur only have a collective 30%
conversion rate there is still a large amount of reactants present in the outlet stream, for this reason the
mole fraction will be different in the outlet stream compared to the inlet stream, the stream inputs and
outputs are shown below:
Table 3. Material balance around Reactor
Component Input(kmol/h) Consumed(Kmol/h
)
Output(Kmol/h
)
Mole
Fraction
Acetic Acid 855.9280188 213.9820047 641.9460141 0.176133185
Ethylene 1747.621118 429.5934086 1318.027709 0.361632307
Oxygen 961.4858647 500.1318647 461.354 0.126583463
VAC 10.19419971 0 436.9052794 0.119875374
Reactor
REACT-IN REACT-OUT

7
Water 68.1870525 0 611.6673911 0.167825523
Carbon
Dioxide
1.246252032 0 174.7621118 0.047950149
Total 3644.662505 3644.662505 1
1.2.2 Energy Balance for Reactor
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡
+ 𝑄 + 𝐻𝑅 (13)
Fin
= Flow rate of feed in (kg/h)
CPx
in
= Specific heat capacity of component in feed in (kJ/kg.k)
Fout
= Flow rate of feed out (kg/h)
CPx
out
= Specific heat capacity of component in feed out (kJ/kg.k)
Q = Heat Required by process (kJ/h)
HR = Total heat from reaction (kJ/h)
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (14)
𝑜𝑢𝑡
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡 ) (15)
In the above equation we can see that the overall energy balance of the components passing through
the Reactor is defined by: the rate of the flow in to the reactor multiplied by the specific heat capacity
of the fraction of each component present in the inlet feed. This fraction is calculated by dividing the
specific heat capacity of feed in, by the specific heat capacity of one mole of the component. The feed
in is subject to the overall specific heat capacity of the components exiting the reactor too. As a
reaction is taking place, new products are formed and the amount of product formed is equal to the
amount of product entering the reactor with a 25% conversion of Ethylene, Acetic Acid and Oxygen.
We assume that no heat is lost to the surroundings. The flow rate of the outlet stream is likely to be
different from the inlet stream to take in to account the reaction, which is assumed to be a in steady
state. The Q value is dependent on the product of the mass of each component, the specific heat
capacity and the change in temperature in the process, thus defining the heat energy required by the
process. HR is defined by the total energy change from produced by the reaction, in this case the
reaction is exothermic and a negative value will be expected. The table below show the results
obtained for the stream in and out for the reactor.

8
Table 4. Energy balance around Reactor
UNIT BLOCK INPUT (KJ/HR) OUTPUT (KJ/HR)
Reactor 114,347,154 153,635,788
The block diagram below represents the cooler used to cool the reactor outlet stream to the inlet of the
separator. And the energy profile is shown in the table below.
Table 5. Energy balance around cooler
Cooler 153,635,788 138,677,465.3
Note: the enthalpy of each stream is calculated in Kj/h and then converted to cal/sec using the
correlation stated in equation (11).
1.3 Separator
1.3.1 Mass Balance for Separator
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
𝑜𝑢𝑡
+ 𝑅𝑖 (16)
Fin
Xi
in
Fout
Cooler
Reactor
REACT-OUT COOL-OUT

9
Xi
in
Ri = Reaction
𝑅𝑖 = 0
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑣𝑎𝑝 ∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+
𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑙𝑖𝑞 ∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) (17)
The phase separator exists to remove the gas created from the reaction. We assume that the separator
is 95% effective in riding the CO2, Ethylene and Oxygen in the below figures as opposed to 100%
effective in the equation. In the above equation we can see that the overall amount of material present
passing through the separator is defined by the rate of the flow in to the separator multiplied by the
fraction of each component present in the inlet feed. This fraction is calculated by dividing the mass
of feed in by the mass of one mole of the component. The feed in, is subject to the overall mass
exiting the too. We assume that no accumulation occurs inside the separator. As no reaction occurs in
the separator we can assume no new compounds and products are formed. The flow rate exiting the
reactor is assumed to be different from the inlet flow to account for the reaction process and is thus
defined by a separate variable: Fout1&2
. We assume the values for this process below.
Table 6. Mass balance around Separator
Comp Input(Kmol/h) Output Gas
CO2-IN
Output Liq
Dis-IN
Mole Fr Gas Mole Fr Liq
Acetic
Acid
641.9460141 32.0973007 609.8487134 0.016536795 0.35795547
Ethylene 1318.027709 1252.126323 65.90138545 0.64510586 0.038681333
Oxygen 461.354 438.2863 23.0677 0.225808734 0.013539767
VAC 436.9052794 21.84526397 415.0600154 0.011254861 0.243622721
Water 611.6673911 30.58336956 581.0840216 0.015756805 0.341071809
Carbon
Dioxide
174.7621118 166.0240062 8.738105588 0.085536944 0.005128899
Total 3644.662505 1940.962564 1703.699941 1 1
Separator
Liquid
Vapour

10
1.3.2 Energy Balance for Separator
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡
+ 𝑄 (18)
Fin
CPx
in
= Specific heat capacity of component in feed in (kj/kg.k)
Fout
CPx
out
= Specific heat capacity of component in feed out (kj/kg.k)
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (19)
𝑙𝑖𝑞
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
) (20)
𝑣𝑎𝑝
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
) (21)
𝑑𝑈
𝑑𝑡
= ((𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
) − (𝐹𝑜𝑢𝑡 ∗ 𝐶𝑃𝑥
𝑜𝑢𝑡1
𝑜𝑢𝑡2
)) + 𝑄 (22)
the Separator is defined by: the rate of the flow in to the Separator multiplied by the specific heat
capacity of the fraction of each component present in the inlet feed. The feed in is subject to the
overall specific heat capacity of the components exiting the separator too. As no reaction is taking
place thus no new products are formed the amount of product in is equal to the amount of product
leaving the Separator. The Q value is dependent on the product of the mass of each component. The
table below shows the results obtained for the streams in and out for the separator.
Separator
liquid
Vapour

11
Table 7. Energy balance around Separator
Separator 138677465 26439828.8
69132693.4
The block diagram below represents the heater used to heat the separator vapour outlet stream to the
inlet of the absorber. And the energy profile is shown in the table below.
Table 8. Energy balance around Heater
Heater 26464969.68 31469997
1.4 CO2 Separator
1.4.1 Mass Balance for CO2-Separator
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
𝑜𝑢𝑡
+ 𝑅𝑖 (23)
Fin
Xi
in
Fout
Xi
in
Ri = Reaction
Heater
liquid
Vapour

12
𝑅𝑖 = 0
𝑑𝑛𝑖
𝑑𝑡
=
𝐹 𝑖𝑛
∗ (𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑐 ∗ (𝑋 𝐶𝑂2
𝑜𝑢𝑡
+ 𝑋 𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐸𝑇𝐻
𝑜𝑢𝑡
) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑚 ∗
(𝑋 𝐶𝑂2
𝑜𝑢𝑡
) (24)
The CO2 separator exists to remove the CO2 gas content created from the reaction. We assume that
the separator is 95% effective in riding the CO2, Ethylene and Oxygen in the below figures as
opposed to 100% effective in the equation. In the above equation we can see that the overall amount
of material present passing through the separator is defined by the rate of the flow in to the separator
multiplied by the fraction of each component present in the inlet feed. This fraction is calculated by
dividing the mass of feed in by the mass of one mole of the component. The feed in, is subject to the
overall mass exiting the separator too. We assume that no accumulation occurs inside the separator.
As no reaction occurs in the separator we can assume no new compounds and products are formed.
The flow rate exiting the reactor is assumed to be different from the inlet flow to account for the
reaction process and is thus defined by a separate variable: Fout1&2
. We assume the values for this
process below.
Table 9. Mass balance around CO2 Separator
Compon
ents
Input
Sep
Input
Flash
Total
input
Output
CO2
Output
Rec
Mole fr
Output
CO2
Mole fr
Output
Rec
Acetic
Acid
32.09730
07
28.96781
388
61.06511
459
1.221302
292
59.84381
23
0.005841
81
0.031557
106
Ethylene 1252.126
323
59.47600
036
1311.602
324
26.23204
648
1285.370
277
0.125474
782
0.677807
189
Oxygen 438.2863 20.81859
925
459.1048
993
9.182097
985
449.9228
013
0.043920
391
0.237255
299
VAC 21.84526
397
19.71535
073
41.56061
47
0.831212
294
40.72940
241
0.003975
907
0.021477
61
Water 30.58336
956
27.60149
102
58.18486
058
1.163697
212
57.02116
337
0.005566
27
0.030068
654
Carbon
Dioxide
166.0240
062
7.886140
293
173.9101
465
170.4319
435
3.478202
929
0.815220
839
0.001797
459
CO2
Separator
Recycle
CO2

13
Total 2105.427
959
209.0622
998
1896.365
66
1 1
1.4.2 Energy Balance for CO2-Separator
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡
+ 𝑄 (25)
Fin
CPx
in
Fout
CPx
out
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (26)
𝑙𝑖𝑞
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡 ) (27)
𝑣𝑎𝑝
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡 ) (28)
𝑑𝑈
𝑑𝑡
= ((𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡1
𝑜𝑢𝑡2
)) + 𝑄 (29)
the CO2 Separator is defined by: the rate of the flow in to the Separator multiplied by the specific heat
table below shows the results obtained for the streams in and out for the CO2 separator.

14
Table 10. Energy balance around CO2 Separator
CO2 Separator 31469996.69
5029970.237
1629751
34566663.37
1.5 Distillation Column
1.5.1 Mass Balance for Distillation Column
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
𝑜𝑢𝑡
+ 𝑅𝑖 (30)
Fin
Xi
in
Fout
Xi
in
Ri = Reaction
𝑅𝑖 = 0
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑣𝑎𝑝 ∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑙𝑖𝑞 ∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) (31)
Distillation
Column
Liquid
Vapour
CO2
Separator
Recycle
CO2

15
The distillation column exists to remove the Acid from our system due to the higher boiling point of
the acid it stays in the bottom of the distillation column and enters the recycle stream. We assume that
the Distillation column is 95% effective in ridding the acid in the below figures; as opposed to 100%
effective in the equation. In the above equation we can see that the overall amount of material present
passing through the distillation column is defined by the rate of the flow in to the separator multiplied
by the fraction of each component present in the inlet feed. The feed in, is subject to the overall molar
flow rate exiting the distillation column to the collective mass from the 2 outlet streams. We assume
that no accumulation occurs inside the column. As no reaction occurs in the separator we can assume
no new compounds and products are formed. The flow rate exiting the reactor is assumed to be
different from the inlet flow to account for the reaction process and is thus defined by a separate
variable: Fout1&2
. We assume the values for this process below.
Table 11. Mass balance around Distillation column
Components Input Dis-in
(Kmol/h)
Output
bottoms
(Kmol/h)
Output
Distillate
(Kmol/h)
Mole fr
Bottoms
Mole fr
Distillate
Acetic Acid 609.8487134 579.3562777 30.49243567 0.9137407 0.028506899
Ethylene 65.90138545 3.295069272 62.60631617 0.0051968 0.058529661
Oxygen 23.0677 1.153385 21.914315 0.0018190 0.020487349
VAC 415.0600154 20.75300077 394.3070146 0.0327309 0.368631429
Water 581.0840216 29.05420108 552.0298205 0.0458232 0.516084001
Carbon
Dioxide
8.738105588 0.436905279 8.301200308 0.0006890 0.007760662
Total 1703.69994 634.0488391 1069.651102 1 1
1.5.2 Energy Balance for Distillation Column
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡
+ 𝑄 (32)
Fin
CPx
in
Fout
CPx
out

16
Q = Heat required by process (kJ/h)
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (33)
𝑙𝑖𝑞
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
) (34)
𝑣𝑎𝑝
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡 ) (35)
𝑑𝑈
𝑑𝑡
= ((𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
) − (𝐹𝑜𝑢𝑡,𝑣𝑎𝑝 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥
𝑜𝑢𝑡,𝑣𝑎𝑝
) − (𝐹𝑜𝑢𝑡,𝑙𝑖𝑞 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥
𝑜𝑢𝑡,𝑙𝑖𝑞
)) + 𝑄 (36)
the Distillation Column is defined by: the rate of the flow in to the Column multiplied by the specific
heat capacity of the fraction of each component present in the inlet feed. As no reaction is taking place
thus no new products are formed the amount of product in is equal to the amount of product leaving
the Column. The Q value is dependent on the product of the mass of each component, the specific
heat capacity and the change in temperature in the process, thus defining the heat energy required by
the process. The table below shows the results obtained for the streams in and out for the distillation
column.
Table 12. Energy balance around Distillation Column
Distillation column 69132693.4 57250227.6
3800889.67
Distillation
Column
Liquid
Vapour

17
1.6 Flash Drum
1.6.1 Mass Balance for Flash Drum
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
𝑜𝑢𝑡
+ 𝑅𝑖 (37)
Fin
Xi
in
Fout
Xi
in
Ri = Reaction
𝑅𝑖 = 0
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑐 ∗ (𝑋 𝐶𝑂2
𝑜𝑢𝑡
+ 𝑋 𝑂2
𝑜𝑢𝑡
+ 𝑋 𝐸𝑇𝐻
𝑜𝑢𝑡
) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑚 ∗ (𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
+ 𝑋 𝐻2𝑂
𝑜𝑢𝑡
) (38)
The flash drum is the third and final separator which exists to remove any excess gas that was not
separated at the first phase separator due to its error. We assume that the separator is 95% effective in
riding the CO2, Ethylene and Oxygen in the below figures as opposed to 100% effective in the above
equation. In the above equation we can see that the overall amount of material present passing through
the separator is defined by the rate of the flow in to the separator multiplied by the fraction of each
component present in the inlet feed. The feed in, is subject to the overall mass exiting the separator
too. We assume that no accumulation occurs inside the separator. As no reaction occurs in the
separator we can assume no new compounds and products are formed.
Table 13. Mass balance around Flash drum
Component
s
Input Flash
(kmol/h)
Output Liq
(kmol/h)
Output Vap
(kmol/h)
Mole fr Liq Mass fr
Vap
Flash drum
liquid
Vapor

18
Acetic Acid 30.49243567 1.524621783 28.96781388 0.00168431
9
0.17613318
5
Ethylene 62.60631617 3.130315809 59.47600036 0.00345820
3
0.36163230
7
Oxygen 21.914315 1.09571575 20.81859925 0.00121048
7
0.12658346
3
VAC 394.3070146 374.5916639 19.71535073 0.41382852
3
0.11987537
4
Water 552.0298205 524.4283295 27.60149102 0.57935993
2
0.16782552
3
Carbon
Dioxide
8.301200308 0.415060015 7.886140293 0.00045853
6
0.04795014
9
Total 1069.651102 905.1857067 164.4653955 1 1
1.6.2 Energy Balance for Flash Drum
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡
+ 𝑄 (39)
Fin
CPx
in
Fout
CPx
out
Q = Heat required by process (kJ/h)
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (40)
𝑙𝑖𝑞
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡 ) (41)
𝑣𝑎𝑝
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
) (42)

19
𝑑𝑈
𝑑𝑡
= ((𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡1
𝑜𝑢𝑡2
)) + 𝑄 (43)
the Flash Separator is defined by: the rate of the flow in to the Separator multiplied by the specific
heat capacity of the fraction of each component present in the inlet feed. The feed in is subject to the
table below shows the results obtained for the streams in and out for the Flash separator.
Table 14. Energy balance around Flash drum
Flash drum 42052861 5029970
46854259
1.7 Decanter
1.7.1 Mass Balance for Decanter
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝑋𝑖
𝑖𝑛
𝑜𝑢𝑡
+ 𝑅𝑖 (44)
Fin
Xi
in
Fout
Xi
in
= Concentration of component in feed out
Ri = Reaction
Flash drum
liquid
Vapor

20
𝑅𝑖 = 0
𝑑𝑛𝑖
𝑑𝑡
= 𝐹 𝑖𝑛
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑎𝑞𝑢𝑒𝑜𝑢𝑠 ∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+
𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) − 𝐹𝑜𝑢𝑡,𝑜𝑟𝑔𝑎𝑛𝑖𝑐 ∗ (𝑋 𝑉𝐴𝐶
𝑖𝑛
+ 𝑋 𝐶𝑂2
𝑖𝑛
+ 𝑋 𝐻2𝑂
𝑖𝑛
+ 𝑋 𝑂2
𝑖𝑛
+ 𝑋 𝐸𝑇𝐻
𝑖𝑛
+ 𝑋𝐴𝐶𝐸
𝑖𝑛
) (45)
The decanter is our final process, which separates our final products of water and Vinyl Acetate. We
assume that the Decanter is 95% effective in separating the mixture in the below figures; as opposed
to 100% effective in the equation. In the above equation we can see that the overall amount of
material present passing through the decanter is defined by the rate of the flow in to the decanter
multiplied by the fraction of each component present in the inlet feed. The feed in, is subject to the
overall mass exiting the decanter in the 2 outlet streams collectively due to the assumption that no
accumulation occurs in the decanter.. As no reaction occurs in the decanter we can assume no new
compounds and products are formed. The flow rate exiting the reactor is assumed to be different from
the inlet flow to account for the decantation process and is thus defined by a separate variable: Fout1&2
.
We assume the values for this process below.
Table 15. Mass balance around Decanter
Components Input
Decanter
Output
Water
Output Vac Mole fr
Acetic Acid 1.524621783 1.311174734 0.21344705 0.000557381
Ethylene 3.130315809 2.692071595 0.438244213 0.0011444
Oxygen 1.09571575 0.942315545 0.153400205 0.000400578
VAC 374.5916639 18.7295832 355.8620807 0.929273143
Water 524.4283295 498.206913 26.22141647 0.068472758
Carbon Dioxide 0.415060015 0.356951613 0.058108402 0.00015174
Total 905.1857067 522.2390097 382.9466971 1
Decanter
Organic
Aqueous

21
1.7.2 Energy Balance for Decanter
𝑑𝑈
𝑑𝑡
= 𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
𝑜𝑢𝑡
+ 𝑄 (46)
Fin
CPx
in
Fout
CPx
out
𝑖𝑛
= (𝐶𝑝 𝑂2
𝑖𝑛
∗ 𝑋 𝑂2
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐶𝑂2
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
𝑖𝑛
∗ 𝑋 𝐻𝑂2
𝑖𝑛
𝑖𝑛
∗
𝑋 𝑉𝐴𝐶
𝑖𝑛
) (47)
𝑙𝑖𝑞
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡
) (48)
𝑣𝑎𝑝
= (𝐶𝑝 𝑂2
𝑜𝑢𝑡
∗ 𝑋 𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐶𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
𝑜𝑢𝑡
∗ 𝑋 𝐻𝑂2
𝑜𝑢𝑡
𝑜𝑢𝑡
∗
𝑋 𝑉𝐴𝐶
𝑜𝑢𝑡 ) (49)
𝑑𝑈
𝑑𝑡
= ((𝐹 𝑖𝑛
∗ 𝐶𝑃𝑥
𝑖𝑛
) − (𝐹𝑜𝑢𝑡,𝑎𝑞𝑢𝑒𝑜𝑢𝑠 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥
𝑜𝑢𝑡,𝑎𝑞𝑢𝑒𝑝𝑢𝑠
) − (𝐹𝑜𝑢𝑡,𝑜𝑟𝑔𝑎𝑛𝑖𝑐 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥
𝑜𝑢𝑡,𝑜𝑟𝑔𝑎𝑛𝑖𝑐
)) + 𝑄 (50)
the Decanter is defined by: the rate of the flow in to the Decanter multiplied by the specific heat
leaving the Decanter. The Q value is dependent on the product of the mass of each component, the
specific heat capacity and the change in temperature in the process, thus defining the heat energy
required by the process. The table below shows the results obtained for the streams in and out for the
Decanter.

22
Table 16. Energy balance around Decanter
Decanter 46854259 4509222
36022793
Decanter
organic
aqueous

23
1.8 Overall Process
1.8.1 Overall Process Material Balance
𝑑𝑛𝑖
𝑑𝑡
= 𝐹𝑂2−𝐹𝑒𝑒𝑑
𝑖𝑛
+ 𝐹𝐸𝑇𝐻−𝐹𝑒𝑒𝑑
𝑖𝑛
+𝐹𝐴𝑐𝑒𝑡𝑖𝑐−𝐹𝑒𝑒𝑑
𝑖𝑛
− 𝐹𝐴𝑞𝑢𝑒𝑜𝑢𝑠
𝑜𝑢𝑡
− 𝐹𝑂𝑟𝑔𝑎𝑛𝑖𝑐
𝑜𝑢𝑡
− 𝐹𝐶𝑂2−𝑅𝑒𝑚𝑜𝑣𝑒𝑑
𝑜𝑢𝑡
− 𝐹𝑃𝑢𝑟𝑔𝑒
𝑜𝑢𝑡
= 0 (51)
The block diagram represents the material inputs and outputs for the Vinyl Acetate Production
process.
Vinyl Acetate ProductionETH-FEED
O2-FEED
Acetic Feed
Organic
Aqueous
CO2-REMOVED
Recycle
The table below shows the material inputs and outputs for the process
Table 17. Overall Process Mass balance
Stream Input(kmol/h) Stream Output (kmol/h)
O2-Feed 523.416 Aqueous 522.24
ETH-Feed 890 Organic 382.94
Acetic-FEED 490 Co2-Removed 209.062
- - Recycle 788.616
Total 1903.416 Total 1903.416
Note: in the material balance calculations, it was assumed that a very negligible fraction was split to
the purge stream, therefore, the recycle stream is an output for the overall process.

24
1.8.2 Overall Process Energy Balance
𝑑𝑈
𝑑𝑡
= 𝐹𝑂2−𝐹𝑒𝑒𝑑
𝑖𝑛
𝐶𝑝 𝑂2
𝑖𝑛
+ 𝐹𝐸𝑇𝐻−𝐹𝑒𝑒𝑑
𝑖𝑛
𝐶𝑝 𝐸𝑡ℎ
𝑖𝑛
+𝐹𝐴𝑐𝑒𝑡𝑖𝑐−𝐹𝑒𝑒𝑑
𝑖𝑛
𝐶𝑝 𝐴𝑐𝑒𝑡𝑖𝑐
𝑖𝑛
− 𝐹𝐴𝑞𝑢𝑒𝑜𝑢𝑠
𝑜𝑢𝑡
𝐶𝑝 𝐴𝑞𝑢𝑒𝑜𝑢𝑠
𝑜𝑢𝑡
− 𝐹𝑂𝑟𝑔𝑎𝑛𝑖𝑐
𝑜𝑢𝑡
− 𝐹𝐶𝑂2−𝑅𝑒𝑚𝑜𝑣𝑒𝑑
𝑜𝑢𝑡
𝐶𝑝 𝐶𝑂2
𝑜𝑢𝑡
− 𝐹𝑅𝑒𝑐𝑦𝑐𝑙𝑒
𝑜𝑢𝑡
𝐶𝑝 𝑅𝑒𝑐𝑦𝑐𝑙𝑒
𝑜𝑢𝑡
(52)
Vinyl Acetate ProductionETH-FEED
O2-FEED
Acetic Feed
Organic
Aqueous
CO2-REMOVED
Recycle
The table below shows the energy inputs and outputs for the process
Table 18. Overall process energy balance
Stream Input(kJ/h) Stream Output (kJ/h)
O2-Feed 1158465 Aqueous 36022793
ETH-Feed 4338842 Organic 4509222
Acetic-FEED 7890228 Co2-Removed 1629751
- - Recycle 34566663.37
Total 12877533.89 Total 76728429.37
2.0 Simulating the System
This section will attempt to provide a step-by-step explanation with regards to setting up the
simulation using the Aspen Plus V8.4 software. Although a process brief was provided, operating
conditions for the units were not presented in the document. A reference was cited instead, from
which the necessary information can be extracted and adapted to suit this particular simulation. The
first step is to specify the components in the system, followed by selecting a suitable thermodynamic.

25
As per the process brief handed out at the beginning of this report, the following components are
selected: The method chosen is WILS-LR.
Component Formula
Acetic acid C2H4O2
Carbon dioxide CO2
Ethylene C2H4
Oxygen O2
Vinyl acetate C4H6O2
Water H2O
Table 19: Components in the system
The first unit to be inserted in to the flowsheet is the mixer, on to which three feed streams of acetic
acid, ethylene and oxygen are then added. All three streams are specified as coming in at 30o
C and
150 psia (Luyben & Tyreus, 1998). The flow rates of all three feeds were entered as specified on the
process brief.
Figure 1: The mixer with the three main feed streams going in
The stream coming out of the Mixer is then directed into a feed Preheater, which is specified as being
at 148.5o
C and 128 psia (Luyben & Tyreus, 1998), with the ‘Vapour-Liquid’ option selected as the
valid phase. The stream coming out is then directed in to a stoichiometric Reactor, which is specified
as being at 158.9o
C and 90 psia. The two reactions taking place inside the reactor are also specified,
along with the conversions – as shown in Figures 3-4.
Figure 2: Reactor with pre-heated feed stream, and exit stream passing through cooler

26
Figure 3: Reaction 1 inside the reactor, with 25% conversion of ethylene
Figure 4: Reaction 2 inside the reactor, with 5% conversion of ethylene
The exit stream from the Reactor is then cooled to 134o
C (Luyben & Tyreus, 1998), and it is assumed
that there is negligible pressure drop across the unit i.e. the pressure of the exit stream from the Cooler
is 90 psia. The stream is then passed through a Separator, which is at 42.5o
C and 84 psia (Luyben &
Tyreus, 1998), to separate out the vapour phase from the liquid phase (conditions). The liquid phase
goes into a distillation column, to remove the vinyl acetate (VAC) and the water from the acetic acid
(ACE), where the purity of the VAC is required to be 0.86. The purity depends on the number of the
stages, which is known to be 20 (Luyben & Tyreus, 1998), and the reflux ratio. The latter is not
specified in the reference material and in order to get an initial estimate, the distillation column is
simulated using the DSTWU unit as it is a shortcut method used to estimate initial Reflux Ratio for a
specified number of pages. The RadFrac distillation column used in the simulation is a rigorous model
which leads us to using the DSTWU for the above stated estimation.

27
Figure 5: Liquid phase from Separator going in to the DSTWU
For the DSTWU unit, the number of stages is specified as being 20 while the recovery of VAC (light
key) in the distillate is specified at 0.9999 (Luyben & Tyreus, 1998). The concentration of ACE in the
Decanter has to be less than 600 moles per million. Since the feed to the Decanter comes from the
distillate of the DSTWU, the concentration in this stream cannot exceed 0.0006 – the recovery of the
heavy key in the distillate is therefore specified as being at this value. The pressure of both the
Condenser and the Reboiler are set at 30 psia (Luyben & Tyreus, 1998). The Condenser is specified as
being a partial condenser with all vapour distillate.
At this point, the simulation is ready to be re-initialised and run. Once it has converged, clicking on
the Results of the DSTWU shows that the minimum reflux ratio is approximately 2.
Figure 6: Results of the DSTWU
The DSTWU can be now be replaced by a RadFrac. Under the Configurations tab, the number of
stages is again set to 20, while the type of condenser is specified as being partial-vapour. The reflux
ratio is set to 2, as per the results from the DSTWU unit. Since it is desired for as much of the ACE as
possible to leave in the bottoms, the ‘Bottoms to feed ratio’ is specified as 1. By clicking on the Feed
Basis button, the key component to be removed in the bottoms is specified i.e. ACE. Next, under the
Streams tab, the feed is set to come in at Stage 6 with the Convention set at ‘On-Stage’ – the product
streams are set by default. Under the Pressure tab, the pressure of the Condenser is set at 30 psia. The
simulation is then re-initialised and run so that it converges.

28
The distillate is then passed through a Flash Drum, specified as being at 80o
C and 128 psia (Luyben &
Tyreus, 1998) to separate the vapour phase from the liquid. The liquid stream then passes through a
Decanter, where the organic and the aqueous phases are separated out. Under the Specifications tab
for the Decanter, the pressure and temperature are set at 18psia and 40o
C, respectively (Luyben &
Tyreus, 1998). The key components are identified as VAC and water. The ‘Component mole fraction’
is set to 0.5 by default. Next, under the Calculation Options tab, the valid phases are set to ‘Liquid-
FreeWater’ – all other settings are left in their default mode.
The vapour from the Flash Drum is sent to an Absorber, which is modelled by a separation unit in
order to simplify the process. Under the Specification tab, the Outlet stream is set to CO2-OUT,
which refers to the CO2 removed from the entire process. The Stream spec is left in its default mode
of ‘Split fraction’, and the 1st
Spec column pertaining to Split fraction is filled out. The Split fraction
of CO2 is set to 1, meaning that the flow of this component in the feed stream should go in its entirety
to the CO2-OUT stream – all other components are set to zero. Note that the 2nd
Spec column does not
need to be filled in, as only one column is required. Under the Feed Flash tab, the pressure is set to
128 psia. Finally, under the Outlet Flash tab, the temperature and pressure are set to 40.4o
C and 128
psia (Luyben & Tyreus, 1998) while ‘Vapour-Only’ is chosen as the valid phase. The simulation is
then re-initialised and run so that it converges. The bottoms from the RadFrac unit i.e. the ACE-REC
stream is one of the two recycle streams in the process, where the other is the RECYCLE stream
coming out of the Absorber.
Figure 7: Products streams from Flash Drum going in to Decanter and Absorber
A final intermediate stream, going from the Separator to the Absorber needs to be connected before
the recycle streams can be connected back to the Mixer. The vapour phase from the Separator is
passed through a Heater, to bring the temperature of the stream up to 80o
C – the pressure of the unit is
set to 128 psia (Luyben & Tyreus, 1998). The simulation is then re-initialised and run once again.
The RECYCLE stream is connected to a Divider, where it is split into two further streams. Under the
Specifications tab for the Divider unit, the two relevant streams (PURGE and ETH-REC) are selected.
The Specification column is set to ‘Split fraction’, which is the fraction of the feed in each respective
outlet stream, by default. Since a larger fraction of the feed stream is desired in the ETH-REC stream,
an initial guess of 0.8 is entered in to the Value column for this stream. The simulation is then re-
initialised before running.

29
An initial run did not converge, and returned an error message as shown in Figure 7. An investigation
into the problem showed that there was zero flow of O2 in the outlet stream of the Reactor.
Additionally, the RadFrac was not in Mass Balance, and the simulation did not converge in 30
iterations – which is the default setting. The Split Fraction of ETH-REC in the Divider was therefore
decreased to 0.7. This resolved the issue with the Reactor, but not the Radfrac – which had a relative
error of 0.122E-03 in the mass balance between the inlet and outlet streams.
Clearly, the Split Fraction is in the region of 0.7-0.8. This parameter was adjusted in increments of
0.01, with the relative error in the mass balance of the RadFrac decreasing slowly. Through a trial-and
error run, it was determined that the threshold value for the Split Fraction of ETH-REC at which the
simulation converges is 0.789.
Figure 8: Aspen Plus V8.4 simulation of a vinyl acetate monomer process

30
2.1 Simulation Input file
The input file of the simulation from Aspen Plus is attached below;
;
;Input Summary created by Aspen Plus Rel. 30.0 at 07:45:11 Tue Dec 16, 2014
;Directory C:ProgramDataAspenTechAspen Plus V8.4Group 6 trial(rec) Filename
C:UsersBILLYB~1AppDataLocalTemp~ape5e3.txt
;
DYNAMICS
DYNAMICS RESULTS=ON
IN-UNITS MET PRESSURE=bar TEMPERATURE=C DELTA-T=C PDROP=bar &
INVERSE-PRES='1/bar'
DEF-STREAMS CONVEN ALL
MODEL-OPTION
DATABANKS 'APV84 PURE28' / 'APV84 AQUEOUS' / 'APV84 SOLIDS' / &
'APV84 INORGANIC' / NOASPENPCD
PROP-SOURCES 'APV84 PURE28' / 'APV84 AQUEOUS' / 'APV84 SOLIDS' &
/ 'APV84 INORGANIC'
COMPONENTS
CO2 CO2 /
C2H4 C2H4 /
O2 O2 /

31
ACE C2H4O2-1 /
VINYL-01 C4H6O2-1 /
WATER H2O
SOLVE
RUN-MODE MODE=SIM
FLOWSHEET
BLOCK MIXER IN=ETHYLENE ACETIC O2-FEED C2H4-REC ACID-REC &
OUT=FEED
BLOCK PREHEAT IN=FEED OUT=REAC-IN
BLOCK REACTOR IN=REAC-IN OUT=REAC-OUT
BLOCK SEP IN=COOL-OUT OUT=SEP-OUT DIST-IN
BLOCK DIST IN=DIST-IN OUT=DIST-VAP ACID-REC
BLOCK FLASH IN=DIST-VAP OUT=CO2 DEC-IN
BLOCK DECANTER IN=DEC-IN OUT=ORGANIC AQUEOUS
BLOCK ABSORBER IN=CO2 ABS-IN OUT=CO2-OUT RECYCLE
BLOCK SPLIT IN=RECYCLE OUT=PURGE C2H4-REC
BLOCK COOLER IN=REAC-OUT OUT=COOL-OUT
BLOCK HEATER2 IN=SEP-OUT OUT=ABS-IN
PROPERTIES WILS-LR
PROPERTIES WILS-LR
ESTIMATE ALL
STREAM ACETIC
SUBSTREAM MIXED TEMP=30.00000000 PRES=10.34213590 &

32
MOLE-FLOW=490.0000000
MOLE-FLOW ACE 490.0000000
STREAM ETHYLENE
MOLE-FLOW=890.0000000
MOLE-FLOW C2H4 890.0000000
STREAM O2-FEED
MOLE-FLOW=523.4166666
MOLE-FLOW O2 523.4166666
BLOCK MIXER MIXER
PARAM
BLOCK SPLIT FSPLIT
PARAM PRES=150. <psia>
FRAC C2H4-REC 0.789
STREAM-ORDER PURGE 2 / C2H4-REC 1
BLOCK ABSORBER SEP2
PARAM PRES=8.825289340
FRAC STREAM=CO2-OUT SUBSTREAM=MIXED COMPS=CO2 C2H4 O2 &
ACE VINYL-01 WATER FRACS=1. 0. 0. 0. 0. 0.
FLASH-SPECS CO2-OUT TEMP=40.40000000 PRES=8.825289340 &
NPHASE=1 PHASE=V FREE-WATER=NO
FLASH-SPECS RECYCLE TEMP=47.7 PRES=128. <psia>

33
BLOCK COOLER HEATER
PARAM TEMP=134.0000000 PRES=6.205281560
BLOCK HEATER2 HEATER
PARAM TEMP=80.00000000 PRES=8.825289340
BLOCK PREHEAT HEATER
PARAM TEMP=148.5000000 PRES=8.825289340
BLOCK FLASH FLASH2
PARAM TEMP=80.00000000 PRES=128. <psia>
BLOCK SEP FLASH2
PARAM TEMP=42.50000000 PRES=5.791596130
BLOCK DECANTER DECANTER
PARAM TEMP=40.00000000 PRES=18. <psia> LL-METH=EQ-SOLVE &
L2-COMPS=WATER L2-CUTOFF=0.5
BLOCK-OPTION FREE-WATER=YES
BLOCK DIST RADFRAC
PARAM NSTAGE=20 ALGORITHM=STANDARD INIT-OPTION=STANDARD &
MAXOL=30 NPHASE=2 DAMPING=NONE
COL-CONFIG CONDENSER=PARTIAL-V
RATESEP-ENAB CALC-MODE=EQUILIBRIUM
FEEDS DIST-IN 6
PRODUCTS DIST-VAP 1 V / ACID-REC 20 L

34
P-SPEC 1 140. <psia> / 2 30. <psia>
COL-SPECS B:F=0.975 MOLE-RR=2.5
DB:F-PARAMS COMPS=ACE
SC-REFLUX OPTION=0
BLOCK-OPTION FREE-WATER=NO
BLOCK REACTOR RSTOIC
PARAM TEMP=158.9000000 PRES=6.205281560
STOIC 1 MIXED C2H4 -1. / ACE -1. / O2 -0.5 / VINYL-01 &
1. / WATER 1.
STOIC 2 MIXED C2H4 -1. / O2 -3. / CO2 2. / WATER 2.
CONV 1 MIXED C2H4 0.25
CONV 2 MIXED C2H4 0.05
DESIGN-SPEC VAC
DEFINE VAC MOLE-FRAC STREAM=ORGANIC SUBSTREAM=MIXED &
COMPONENT=VINYL-01
SPEC "VAC" TO "0.86"
TOL-SPEC "0.005"
VARY BLOCK-VAR BLOCK=DECANTER VARIABLE=TEMP SENTENCE=PARAM &
UOM="C"
LIMITS "5" "100"
EO-CONV-OPTI
SENSITIVITY S-1
DEFINE VAC MOLE-FLOW STREAM=REAC-OUT SUBSTREAM=MIXED &
COMPONENT=VINYL-01 UOM="kmol/hr"

35
TABULATE 1 "VAC"
VARY MOLE-FLOW STREAM=O2-FEED SUBSTREAM=MIXED COMPONENT=O2 &
UOM="kmol/day"
RANGE LOWER="12562" UPPER="25124" NPOINT="10"
CONV-OPTIONS
DIRECT MAXIT=100
STREAM-REPOR MOLEFLOW
PROPERTY-REP PCES
;
;
;
;
;
2.2 Simulation Report file
The simulation report file from Aspen Plus is attached below;
BLOCK: ABSORBER MODEL: SEP2
----------------------------
INLET STREAMS: CO2 ABS-IN
OUTLET STREAMS: CO2-OUT RECYCLE
PROPERTY OPTION SET: WILS-LR WILSON / IDEAL GAS
*** MASS AND ENERGY BALANCE ***
IN OUT RELATIVE DIFF.
TOTAL BALANCE
MOLE(KMOL/HR ) 1507.51 1507.51 0.301654E-15

36
MASS(KG/HR ) 46767.6 46767.6 0.155577E-15
ENTHALPY(CAL/SEC ) 0.147164E+07 750305. 0.490157
*** CO2 EQUIVALENT SUMMARY ***
FEED STREAMS CO2E 8069.11 KG/HR
PRODUCT STREAMS CO2E 8069.11 KG/HR
NET STREAMS CO2E PRODUCTION 0.00000 KG/HR
UTILITIES CO2E PRODUCTION 0.00000 KG/HR
TOTAL CO2E PRODUCTION 0.00000 KG/HR
*** INPUT DATA ***
INLET PRESSURE BAR 8.82529
FLASH SPECS FOR STREAM CO2-OUT
ONE PHASE TP FLASH SPECIFIED PHASE IS VAPOR
SPECIFIED TEMPERATURE C 40.4000
SPECIFIED PRESSURE BAR 8.82529
MAXIMUM NO. ITERATIONS 30
CONVERGENCE TOLERANCE 0.000100000
FLASH SPECS FOR STREAM RECYCLE
TWO PHASE TP FLASH

37
SPLIT FRACTION
SUBSTREAM= MIXED
STREAM= CO2-OUT CPT= CO2 FRACTION= 1.00000
C2H4 0.0
O2 0.0
ACE 0.0
VINYL-01 0.0
WATER 0.0
*** RESULTS ***
HEAT DUTY CAL/SEC -0.72133E+06
STREAM= CO2-OUT SUBSTREAM= MIXED
COMPONENT = CO2 SPLIT FRACTION = 1.00000
STREAM= RECYCLE SUBSTREAM= MIXED
COMPONENT = C2H4 SPLIT FRACTION = 1.00000
COMPONENT = O2 SPLIT FRACTION = 1.00000
COMPONENT = ACE SPLIT FRACTION = 1.00000
COMPONENT = VINYL-01 SPLIT FRACTION = 1.00000
COMPONENT = WATER SPLIT FRACTION = 1.00000
BLOCK: COOLER MODEL: HEATER
------------------------------
INLET STREAM: REAC-OUT

38
OUTLET STREAM: COOL-OUT
TOTAL BALANCE
MOLE(KMOL/HR ) 3097.06 3097.06 0.00000
MASS(KG/HR ) 124729. 124729. 0.00000
ENTHALPY(CAL/SEC ) 0.900052E+07 0.742285E+07 0.175287
*** INPUT DATA ***
TWO PHASE TP FLASH
*** RESULTS ***
OUTLET TEMPERATURE C 134.00

39
OUTLET PRESSURE BAR 6.2053
OUTLET VAPOR FRACTION 0.91675
V-L PHASE EQUILIBRIUM :
COMP F(I) X(I) Y(I) K(I)
CO2 0.61407E-01 0.10215E-02 0.66891E-01 65.484
C2H4 0.42985 0.88158E-02 0.46808 53.096
O2 0.63243E-03 0.24540E-05 0.68964E-03 281.02
ACE 0.12910 0.40015 0.10449 0.26112
VINYL-01 0.16028 0.17866 0.15861 0.88777
WATER 0.21873 0.41135 0.20124 0.48922
BLOCK: DECANTER MODEL: DECANTER
--------------------------------
INLET STREAM: DEC-IN
FIRST LIQUID OUTLET: ORGANIC
SECOND LIQUID OUTLET: AQUEOUS
FREE WATER OPTION SET: SYSOP12 ASME STEAM TABLE
SOLUBLE WATER OPTION: THE MAIN PROPERTY OPTION SET (WILS-LR ).
TOTAL BALANCE

40
MOLE(KMOL/HR ) 1202.86 1202.86 -0.189027E-15
MASS(KG/HR ) 54902.9 54902.9 0.00000
ENTHALPY(CAL/SEC ) 768569. -0.121784E+08 1.06311
*** INPUT DATA ***
LIQUID-LIQUID SPLIT, TP SPECIFICATION
FREE WATER CONSIDERED
CONVERGENCE TOLERANCE ON EQUILIBRIUM 0.10000E-03
MAXIMUM NO ITERATIONS ON EQUILIBRIUM 30
EQUILIBRIUM METHOD EQUATION-SOLVING
KLL COEFFICIENTS FROM OPTION SET OR EOS
KLL BASIS MOLE
KEY COMPONENT(S): WATER
*** RESULTS ***

41
CALCULATED HEAT DUTY CAL/SEC -0.12947E+08
MOLAR RATIO 1ST LIQUID / TOTAL LIQUID 0.45493
L1-L2 PHASE EQUILIBRIUM :
COMP F X1 X2 K
CO2 0.0056814 0.012489 0.0 0.0
C2H4 0.040626 0.089303 0.0 0.0
O2 0.180849-05 0.397532-05 0.0 0.0
ACE 0.011454 0.025178 0.0 0.0
VINYL-01 0.39062 0.85863 0.0 0.0
WATER 0.55162 0.014395 1.00000 69.4690
BLOCK: DIST MODEL: RADFRAC
-------------------------------
INLETS - DIST-IN STAGE 6
OUTLETS - DIST-VAP STAGE 1
ACID-REC STAGE 20
TOTAL BALANCE
MOLE(KMOL/HR ) 1633.86 1633.86 0.00000
MASS(KG/HR ) 79423.8 79423.8 -0.181386E-13
ENTHALPY(CAL/SEC ) 546774. 0.476780E+07 -0.885319

42
**********************
**** INPUT DATA ****
**********************
**** INPUT PARAMETERS ****
NUMBER OF STAGES 20
ALGORITHM OPTION STANDARD
ABSORBER OPTION NO
INITIALIZATION OPTION STANDARD
HYDRAULIC PARAMETER CALCULATIONS NO
INSIDE LOOP CONVERGENCE METHOD BROYDEN
DESIGN SPECIFICATION METHOD NESTED
MAXIMUM NO. OF OUTSIDE LOOP ITERATIONS 30
MAXIMUM NO. OF INSIDE LOOP ITERATIONS 10
MAXIMUM NUMBER OF FLASH ITERATIONS 30
FLASH TOLERANCE 0.000100000
OUTSIDE LOOP CONVERGENCE TOLERANCE 0.000100000
**** COL-SPECS ****

43
MOLAR VAPOR DIST / TOTAL DIST 1.00000
MOLAR REFLUX RATIO 2.50000
BOTTOMS TO FEED RATIO 0.97500
**** PROFILES ****
P-SPEC STAGE 1 PRES, BAR 9.65266
2 2.06843
*******************
**** RESULTS ****
*******************
*** COMPONENT SPLIT FRACTIONS ***
OUTLET STREAMS
--------------
DIST-VAP ACID-REC
COMPONENT:
CO2 1.0000 0.0000
C2H4 1.0000 0.0000
O2 1.0000 0.0000
ACE .34780E-01 .96522
VINYL-01 1.0000 .30024E-06
WATER .99420 .58001E-02

44
*** SUMMARY OF KEY RESULTS ***
TOP STAGE TEMPERATURE C 169.728
BOTTOM STAGE TEMPERATURE C 142.577
TOP STAGE LIQUID FLOW KMOL/HR 3,117.93
BOTTOM STAGE LIQUID FLOW KMOL/HR 386.685
TOP STAGE VAPOR FLOW KMOL/HR 1,247.17
BOILUP VAPOR FLOW KMOL/HR 6,323.32
MOLAR REFLUX RATIO 2.50000
MOLAR BOILUP RATIO 16.3526
CONDENSER DUTY (W/O SUBCOOL) CAL/SEC -5,639,190.
REBOILER DUTY CAL/SEC 9,860,210.
**** MAXIMUM FINAL RELATIVE ERRORS ****
DEW POINT 0.82625E-06 STAGE= 6
BUBBLE POINT 0.27978E-05 STAGE= 5
COMPONENT MASS BALANCE 0.27679E-05 STAGE= 6 COMP=CO2
ENERGY BALANCE 0.30034E-05 STAGE= 8
**** PROFILES ****
**NOTE** REPORTED VALUES FOR STAGE LIQUID AND VAPOR RATES ARE THE
FLOWS
FROM THE STAGE INCLUDING ANY SIDE PRODUCT.
ENTHALPY
STAGE TEMPERATURE PRESSURE CAL/MOL HEAT DUTY

45
C BAR LIQUID VAPOR CAL/SEC
1 169.73 9.6527 4709.9 12320. -.56392+07
2 115.25 2.0684 2748.9 11535.
4 118.22 2.0684 2693.6 11449.
5 119.17 2.0684 2749.0 11406.
6 120.71 2.0684 2845.8 11459.
7 122.89 2.0684 2776.1 11456.
18 141.84 2.0684 4577.2 10276.
19 142.31 2.0684 4624.7 10251.
20 142.58 2.0684 4652.3 10237. .98602+07
STAGE FLOW RATE FEED RATE PRODUCT RATE
KMOL/HR KMOL/HR KMOL/HR
LIQUID VAPOR LIQUID VAPOR MIXED LIQUID VAPOR
1 3118. 1247. 1247.1722
2 2443. 4365.
4 2461. 3687.
5 2529. 3708. 69.0147
6 4428. 3707. 1564.8427
7 4418. 4042.
18 6638. 6133.
19 6710. 6251.
20 386.7 6323. 386.6851
**** MASS FLOW PROFILES ****
STAGE FLOW RATE FEED RATE PRODUCT RATE

46
KG/HR KG/HR KG/HR
LIQUID VAPOR LIQUID VAPOR MIXED LIQUID VAPOR
1 0.1269E+06 0.5637E+05 .56365+05
2 0.8183E+05 0.1833E+06
4 0.7625E+05 0.1323E+06
5 0.8079E+05 0.1326E+06 2251.9402
6 0.1484E+06 0.1349E+06 .77172+05
7 0.1370E+06 0.1254E+06
18 0.3895E+06 0.3528E+06
19 0.3978E+06 0.3664E+06
20 0.2306E+05 0.3747E+06 .23058+05
**** MOLE-X-PROFILE ****
STAGE CO2 C2H4 O2 ACE VINYL-01
1 0.15477E-03 0.13996E-02 0.50425E-07 0.26411E-01 0.31670
2 0.18859E-04 0.15594E-03 0.41320E-08 0.48962E-01 0.19719
4 0.20723E-04 0.17053E-03 0.47574E-08 0.10955 0.12282
5 0.20344E-04 0.16774E-03 0.47041E-08 0.15059 0.11163
6 0.67512E-05 0.55174E-04 0.22955E-09 0.19989 0.10426
7 0.43366E-07 0.42756E-06 0.31703E-12 0.21889 0.55711E-01
18 0.68655E-32 0.66664E-30 0.51423E-44 0.96734 0.34469E-05
19 0.33554E-34 0.40819E-32 0.62016E-47 0.98170 0.11352E-05
20 0.16331E-36 0.24901E-34 0.74633E-50 0.98997 0.36677E-06
**** MOLE-X-PROFILE ****
STAGE WATER
1 0.65533
2 0.75367

47
4 0.76743
5 0.73759
6 0.69578
7 0.72540
18 0.32656E-01
19 0.18301E-01
20 0.10030E-01
**** MOLE-Y-PROFILE ****
1 0.97818E-02 0.67314E-01 0.10854E-04 0.11060E-01 0.37876
2 0.29054E-02 0.20232E-01 0.31371E-05 0.22025E-01 0.33443
4 0.33231E-02 0.22886E-01 0.36749E-05 0.54023E-01 0.22451
5 0.33037E-02 0.22753E-01 0.36536E-05 0.76428E-01 0.20890
6 0.11191E-02 0.76150E-02 0.17988E-06 0.10633 0.20272
7 0.73971E-05 0.60452E-04 0.25151E-09 0.12430 0.11424
18 0.14834E-29 0.11507E-27 0.45169E-41 0.94000 0.10934E-04
19 0.72902E-32 0.70787E-30 0.54604E-44 0.96594 0.36375E-05
20 0.35596E-34 0.43300E-32 0.65804E-47 0.98119 0.11822E-05
**** MOLE-Y-PROFILE ****
STAGE WATER
1 0.53307
2 0.62040
4 0.69526
5 0.68861
6 0.68221
7 0.76139

48
18 0.59984E-01
19 0.34056E-01
20 0.18807E-01
**** K-VALUES ****
1 63.202 48.096 215.24 0.41877 1.1959
2 154.05 129.74 759.19 0.44983 1.6960
4 160.35 134.20 772.42 0.49312 1.8279
5 162.38 135.64 776.62 0.50753 1.8713
6 165.75 138.01 783.53 0.53196 1.9443
7 170.56 141.38 793.25 0.56787 2.0505
18 216.07 172.61 878.38 0.97174 3.1720
19 217.27 173.42 880.48 0.98395 3.2043
20 217.97 173.89 881.70 0.99113 3.2232
**** K-VALUES ****
STAGE WATER
1 0.81344
2 0.82317
4 0.90596
5 0.93360
6 0.98049
7 1.0496
18 1.8368
19 1.8609
20 1.8750

49
**** MASS-X-PROFILE ****
1 0.16734E-03 0.96463E-03 0.39642E-07 0.38966E-01 0.66985
2 0.24776E-04 0.13059E-03 0.39469E-08 0.87772E-01 0.50676
4 0.29435E-04 0.15440E-03 0.49132E-08 0.21234 0.34127
5 0.28025E-04 0.14730E-03 0.47117E-08 0.28306 0.30083
6 0.88648E-05 0.46181E-04 0.21915E-09 0.35815 0.26781
7 0.61548E-07 0.38681E-06 0.32715E-12 0.42390 0.15467
18 0.51491E-32 0.31871E-30 0.28041E-44 0.98997 0.50571E-05
19 0.24909E-34 0.19316E-32 0.33474E-47 0.99444 0.16485E-05
20 0.12053E-36 0.11715E-34 0.40049E-50 0.99697 0.52952E-06
**** MASS-X-PROFILE ****
STAGE WATER
1 0.29005
2 0.40531
4 0.44621
5 0.41593
6 0.37398
7 0.42143
18 0.10026E-01
19 0.55615E-02
20 0.30303E-02
**** MASS-Y-PROFILE ****
1 0.95254E-02 0.41784E-01 0.76847E-05 0.14696E-01 0.72149
2 0.30454E-02 0.13519E-01 0.23908E-05 0.31502E-01 0.68573

50
4 0.40753E-02 0.17891E-01 0.32769E-05 0.90404E-01 0.53860
5 0.40654E-02 0.17848E-01 0.32690E-05 0.12833 0.50287
6 0.13534E-02 0.58704E-02 0.15817E-06 0.17547 0.47958
7 0.10495E-04 0.54675E-04 0.25946E-09 0.24065 0.31707
18 0.11348E-29 0.56110E-28 0.25123E-41 0.98120 0.16361E-04
19 0.54731E-32 0.33876E-30 0.29806E-44 0.98953 0.53419E-05
20 0.26435E-34 0.20498E-32 0.35531E-47 0.99428 0.17174E-05
**** MASS-Y-PROFILE ****
STAGE WATER
1 0.21249
2 0.26620
4 0.34903
5 0.34688
6 0.33772
7 0.44221
18 0.18783E-01
19 0.10466E-01
20 0.57172E-02
BLOCK: FLASH MODEL: FLASH2
------------------------------
INLET STREAM: DIST-VAP
OUTLET VAPOR STREAM: CO2
OUTLET LIQUID STREAM: DEC-IN

51
TOTAL BALANCE
MOLE(KMOL/HR ) 1247.17 1247.17 0.00000
MASS(KG/HR ) 56365.4 56365.4 -0.387257E-15
ENTHALPY(CAL/SEC ) 0.426808E+07 818109. 0.808319
*** INPUT DATA ***
TWO PHASE TP FLASH
*** RESULTS ***
VAPOR FRACTION 0.35530E-01

52
CO2 0.97818E-02 0.56814E-02 0.12109 21.313
C2H4 0.67314E-01 0.40626E-01 0.79177 19.489
O2 0.10854E-04 0.18085E-05 0.25639E-03 141.77
ACE 0.11060E-01 0.11454E-01 0.35489E-03 0.30982E-01
VINYL-01 0.37876 0.39062 0.56920E-01 0.14572
WATER 0.53307 0.55162 0.29607E-01 0.53673E-01
BLOCK: HEATER2 MODEL: HEATER
------------------------------
INLET STREAM: SEP-OUT
OUTLET STREAM: ABS-IN
TOTAL BALANCE
MOLE(KMOL/HR ) 1463.20 1463.20 0.00000
MASS(KG/HR ) 45305.1 45305.1 -0.160599E-15
ENTHALPY(CAL/SEC ) 675913. 0.142210E+07 -0.524707

53
*** INPUT DATA ***
TWO PHASE TP FLASH
*** RESULTS ***
HEAT DUTY CAL/SEC 0.74619E+06
CO2 0.12164 0.16576E-01 0.12164 21.313
C2H4 0.85246 0.12704 0.85246 19.489
O2 0.13294E-02 0.27235E-04 0.13294E-02 141.77
ACE 0.22080E-02 0.20699 0.22080E-02 0.30982E-01
VINYL-01 0.16408E-01 0.32704 0.16408E-01 0.14572
WATER 0.59567E-02 0.32233 0.59567E-02 0.53673E-01

54
BLOCK: MIXER MODEL: MIXER
-----------------------------
INLET STREAMS: ETHYLENE ACETIC O2-FEED C2H4-REC ACID-REC
OUTLET STREAM: FEED
TOTAL BALANCE
MOLE(KMOL/HR ) 3334.87 3334.79 0.244274E-04
MASS(KG/HR ) 124734. 124729. 0.395686E-04
ENTHALPY(CAL/SEC ) -0.216557E+07 -0.216553E+07 -0.187479E-04
*** INPUT DATA ***
TWO PHASE FLASH
OUTLET PRESSURE: MINIMUM OF INLET STREAM PRESSURES
BLOCK: PREHEAT MODEL: HEATER

55
------------------------------
INLET STREAM: FEED
OUTLET STREAM: REAC-IN
TOTAL BALANCE
MOLE(KMOL/HR ) 3334.79 3334.79 0.00000
MASS(KG/HR ) 124729. 124729. -0.233337E-15
ENTHALPY(CAL/SEC ) -0.216553E+07 0.530555E+07 -1.40816
*** INPUT DATA ***
TWO PHASE TP FLASH

56
*** RESULTS ***
C2H4 0.57030 0.13271E-01 0.57030 43.210
O2 0.15742 0.74342E-03 0.15742 212.91
ACE 0.26247 0.97175 0.26247 0.27159
VINYL-01 0.62768E-02 0.73613E-02 0.62768E-02 0.85737
WATER 0.35358E-02 0.68763E-02 0.35358E-02 0.51703
BLOCK: REACTOR MODEL: RSTOIC
------------------------------
INLET STREAM: REAC-IN
OUTLET STREAM: REAC-OUT
IN OUT GENERATION RELATIVE DIFF.
TOTAL BALANCE
MOLE(KMOL/HR ) 3334.79 3097.06 -237.727 0.00000
MASS(KG/HR ) 124729. 124729. 0.233337E-15

57
ENTHALPY(CAL/SEC ) 0.530555E+07 0.900052E+07 -0.410529
*** INPUT DATA ***
STOICHIOMETRY MATRIX:
REACTION # 1:
SUBSTREAM MIXED :
C2H4 -1.00 O2 -0.500 ACE -1.00 VINYL-01 1.00
WATER 1.00
REACTION # 2:
SUBSTREAM MIXED :
CO2 2.00 C2H4 -1.00 O2 -3.00 WATER 2.00
REACTION CONVERSION SPECS: NUMBER= 2
REACTION # 1:
SUBSTREAM:MIXED KEY COMP:C2H4 CONV FRAC: 0.2500
REACTION # 2:
SUBSTREAM:MIXED KEY COMP:C2H4 CONV FRAC: 0.5000E-01

58
TWO PHASE TP FLASH
SIMULTANEOUS REACTIONS
GENERATE COMBUSTION REACTIONS FOR FEED SPECIES NO
*** RESULTS ***
VAPOR FRACTION 1.0000
REACTION EXTENTS:
REACTION REACTION
NUMBER EXTENT
KMOL/HR
1 475.45
2 95.091

59
CO2 0.61407E-01 0.11759E-02 0.61407E-01 87.542
C2H4 0.42985 0.10622E-01 0.42985 67.836
O2 0.63243E-03 0.33288E-05 0.63243E-03 318.49
ACE 0.12910 0.43071 0.12910 0.50247
VINYL-01 0.16028 0.17812 0.16028 1.5085
WATER 0.21873 0.37938 0.21873 0.96652
BLOCK: SEP MODEL: FLASH2
------------------------------
INLET STREAM: COOL-OUT
OUTLET VAPOR STREAM: SEP-OUT
OUTLET LIQUID STREAM: DIST-IN
TOTAL BALANCE
MOLE(KMOL/HR ) 3097.06 3097.06 0.00000
MASS(KG/HR ) 124729. 124729. 0.00000
ENTHALPY(CAL/SEC ) 0.742285E+07 0.122269E+07 0.835281

60
*** INPUT DATA ***
TWO PHASE TP FLASH
*** RESULTS ***
VAPOR FRACTION 0.47245
CO2 0.61407E-01 0.74667E-02 0.12164 16.291
C2H4 0.42985 0.51383E-01 0.85246 16.590
O2 0.63243E-03 0.82849E-05 0.13294E-02 160.46
ACE 0.12910 0.24274 0.22080E-02 0.90961E-02
VINYL-01 0.16028 0.28912 0.16408E-01 0.56750E-01
WATER 0.21873 0.40928 0.59567E-02 0.14554E-01
BLOCK: SPLIT MODEL: FSPLIT

61
------------------------------
INLET STREAM: RECYCLE
OUTLET STREAMS: PURGE C2H4-REC
TOTAL BALANCE
MOLE(KMOL/HR ) 1324.17 1324.17 0.00000
MASS(KG/HR ) 38698.5 38698.5 0.00000
ENTHALPY(CAL/SEC ) 659639. 659639. 0.00000
*** INPUT DATA ***
FRACTION OF FLOW STRM=C2H4-REC FRAC= 0.78900
STREAM CALCULATION ORDER:
STREAM ORDER
PURGE 2

62
C2H4-REC 1
*** RESULTS ***
STREAM= PURGE SPLIT= 0.21100 KEY= 0 STREAM-ORDER= 2
C2H4-REC 0.78900 0 1

63
2.3 Simulation Heat and Mass balance table
The table below shows the heat and mass balance results from the process simulation on Aspen Plus.
Table 20. Heat and Material balance Table
Heat and Material Balance Table
Stream ID A BS-IN A CETIC A CID-REC A QUEOUS C 2H4-REC C O2 C O2-OUT C OO L-OUT DEC -IN DIST-IN DIST-VA P ETHYLENE FEED O 2-FEED O RGA NIC PURGE REA C -IN REA C -OUT REC YCLE SEP-O UT
From HEA TER2 DIST DEC A NTER SPLIT FLASH A BSO RBER C OO LER FLASH SEP DIST MIXER DEC A NTER SPLIT PREHEA T REA C TO R A BSO RBER SEP
To A BSO RBER MIXER MIXER MIXER A BSO RBER SEP DEC A NTER DIST FLASH MIXER PREHEA T MIXER REA C TO R C OO LER SPLIT HEA TER2
Phase V APO R LIQUID LIQUID LIQUID MIXED V APO R V APO R MIXED LIQUID LIQUID V APO R V APO R MIXED V APO R LIQUID MIXED V APO R V APO R MIXED V APO R
Substream: MIXED
Mole Flow kmol/hr
CO 2 177.9530 0.0 6.4220E-35 0.0 0.0 5.393792 183.3468 190.1808 6.833972 12.22776 12.22776 0.0 6.4309E-35 0.0 6.833972 0.0 6.4309E-35 190.1808 0.0 177.9530
C2H4 1247.119 0.0 9.7976E-33 0.0 1011.807 35.27273 0.0 1331.266 48.87394 84.14667 84.14667 890.0000 1901.808 0.0 48.87394 270.5846 1901.808 1331.266 1282.392 1247.119
O2 1.972217 0.0 2.9823E-48 0.0 1.565198 .0115573 0.0 1.985976 2.20144E-3 .0137588 .0137588 0.0 524.9832 523.4167 2.20144E-3 .4185764 524.9832 1.985976 1.983774 1.972217
AC E 3.250926 490.0000 386.1555 0.0 2.577561 .0159449 0.0 403.3199 13.89753 400.0690 13.91348 0.0 878.7719 0.0 13.89753 .6893099 878.7719 403.3199 3.266871 3.250926
VINYL-01 23.94928 0.0 1.43423E-4 0.0 20.89643 2.535428 0.0 496.3468 469.8620 472.3976 472.3974 0.0 20.89484 0.0 469.8620 5.588273 20.89484 496.3468 26.48471 23.94928
WA TER 8.694784 0.0 3.911612 655.6420 7.900726 1.318810 0.0 677.4452 663.5200 668.7504 664.8388 0.0 11.81242 0.0 7.878028 2.112868 11.81242 677.4452 10.01359 8.694784
Total Flow kmol/hr 1462.939 490.0000 390.0672 655.6420 1044.747 44.54826 183.3468 3100.544 1202.990 1637.605 1247.538 890.0000 3338.270 523.4167 547.3477 279.3937 3338.270 3100.544 1324.141 1462.939
Total Flow kg/hr 45294.83 29425.75 23260.11 11811.57 30531.18 1470.274 8069.057 1.24936E+5 54910.64 79641.02 56380.91 24967.85 1.24936E+5 16748.71 43099.06 8164.866 1.24936E+5 1.24936E+5 38696.05 45294.83
Total Flow l/min 81120.86 460.0146 427.4864 198.1222 44532.30 2470.228 9026.664 2.58185E+5 1076.122 1441.345 79316.73 36150.42 5.25046E+5 21260.37 816.9256 11909.15 2.21014E+5 2.99148E+5 66215.54 1.10487E+5
Temperature C 80.00000 30.00000 142.5775 36.45885 48.20477 80.00000 40.40000 134.0000 80.00000 42.50000 169.7259 30.00000 42.37128 30.00000 36.45885 48.20477 148.5000 158.9000 47.70000 42.50000
Pressure bar 8.825289 10.34214 2.068427 1.241056 10.34214 8.825289 8.825289 6.205282 8.825289 5.791596 9.652660 10.34214 2.068427 10.34214 1.241056 10.34214 8.825289 6.205282 8.825289 5.791596
V apor Frac 1.000000 0.0 0.0 0.0 .9897130 1.000000 1.000000 .9149709 0.0 0.0 1.000000 1.000000 .7428943 1.000000 0.0 .9897130 1.000000 1.000000 .9924615 1.000000
Liquid Frac 0.0 1.000000 1.000000 1.000000 .0102870 0.0 0.0 .0850290 1.000000 1.000000 0.0 0.0 .2571057 0.0 1.000000 .0102870 0.0 0.0 7.53848E-3 0.0
Solid Frac 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0
Enthalpy cal/mol 3498.149 873.2567 4652.306 -68062.72 1792.624 4024.646 1780.208 8616.401 2300.246 1204.876 12319.37 957.1858 -2330.836 -24356.22 1427.707 1792.624 5732.260 10462.36 1792.624 1662.172
Enthalpy cal/gm 112.9837 14.54154 78.01822 -3778.055 61.34182 121.9439 40.45026 213.8340 50.39410 24.77506 272.5901 34.11970 -62.27966 -761.1605 18.13154 61.34182 153.1653 259.6454 61.34182 53.68508
Enthalpy cal/sec 1.42155E+6 1.18860E+5 5.04087E+5 -1.2396E+7 5.20233E+5 49803.05 90665.40 7.42098E+6 7.68659E+5 5.48086E+5 4.26913E+6 2.36638E+5 -2.1614E+6 -3.5412E+6 2.17070E+5 1.39124E+5 5.31551E+6 9.01084E+6 6.59357E+5 6.75460E+5
Entropy cal/mol-K 17.65259 3.032029 13.67130 -38.30636 10.54996 18.92140 10.63330 29.68676 9.163621 6.593331 33.26491 7.292506 -3.922623 -93.62346 6.004709 10.54996 17.42947 34.47307 10.84579 11.84919
Entropy cal/gm-K .5701461 .0504895 .2292648 -2.126326 .3610092 .5733047 .2416122 .7367391 .2007578 .1355743 .7360511 .2599475 -.1048120 -2.925843 .0762583 .3610092 .4657133 .8555214 .3711322 .3827067
Density mol/cc 3.00568E-4 .0177530 .0152077 .0551546 3.91007E-4 3.00568E-4 3.38528E-4 2.00150E-4 .0186315 .0189360 2.62143E-4 4.10323E-4 1.05968E-4 4.10323E-4 .0111668 3.91007E-4 2.51739E-4 1.72743E-4 3.33291E-4 2.20681E-4
Density gm/cc 9.30604E-3 1.066117 .9068555 .9936269 .0114266 9.91997E-3 .0148985 8.06501E-3 .8504398 .9209109 .0118472 .0115110 3.96587E-3 .0131298 .8792939 .0114266 9.42139E-3 6.96064E-3 9.73992E-3 6.83263E-3
A verage MW 30.96153 60.05256 59.63102 18.01528 29.22352 33.00408 44.00980 40.29481 45.64514 48.63261 45.19375 28.05376 37.42532 31.99880 78.74166 29.22352 37.42532 40.29481 29.22352 30.96153
Liq V ol 60F l/min 1964.203 470.6450 372.0793 197.2390 1467.151 58.96492 163.6609 3411.194 1015.946 1446.991 1074.911 1256.555 4033.685 467.2174 818.7075 392.3560 4033.685 3411.194 1859.507 1964.203

64
3.0 Sensitivity Analysis
A sensitivity analysis was carried out on the simulation to see the impact of the feed streams flowrate
on the VAC produced using Aspen Plus.
This was achieved by setting up a sensitivity analysis using the feed streams molar flow rate as
manipulated variables (i.e. molar flow rates of Fresh Acetic acid feed, Fresh Ethylene feed and
Oxygen feed) and the flow rate of the VAC produced (i.e. VAC flow from the reactor outlet stream).
The limits of the manipulated variables was set as;
 Lower bound – Current given flowrate
 Upper bound – Twice the current given flowrate
And the number of points chosen (iterations) was selected as 10 points. And this is run for each
individual feed stream to analyze its impact on VAC yield.
The following screenshots show the sensitivity analysis set up and the respective results for each of
the feed streams.
Figure 9. Manipukated variable setup for fresh Acetic Acid feed

65
Figure 10. Measured Variable Setup for VAC
Figure 2 is a generic setup for measured variable VAC yield flow rate for all the sensitivity analysis
carried out on all required feed streams. This will only be shown once here on this page but applies to
all streams.

66
Figure 11. Sensitivity Analysis Result for Acetic Acid Feed stream

67
Figure 12. Error Message for Acetic feed stream
The four screenshots above (Figures 1 – 4) show the sensitivity analysis setup and results for the
acetic acid feed stream impact on VAC production. We can see the analysis was completed with
errors which may arise from the fact that the Acetic acid is being recycled back into the system.

68
Figure 13.Manipulated Variable Setup for Fresh Ethylene feed

69
Figure 14. Sensitivity Analysis Result for Fresh Ethylene Feed

70
Figure 15. Warning message for Ethylene feed stream
Figures 5 – 7 shows the sensitivity analysis setup and results for the ethylene feed stream impact on
VAC production. The analysis completed with warnings and it is safe to assume that this is as a result
of the recycling of the ethylene back into the system.

71
Figure 16. Manipulated Variable Setup for Oxygen feed

72
Figure 17. Sensitivity Analysis Result for Oxygen Feed

73
Figure 18. Status message with no Errors or Warnings for Oxygen feed.
Finally, figures 8 – 10 show the sensitivity analysis setup and results for the oxygen feed stream
impact on VAC production. The analysis completed with no errors or warnings owing to the premise
that oxygen has no dedicated recycle, this is an assumption and is no way a statement of fact.
It is seen from the results in figure 9 above that an increase in the oxygen feed flow rate corresponds
to an increase in the amount of VAC produced.
NOTE: That the streams selected for this sensitivity analysis are;
1. Oxygen, Fresh Ethylene and Fresh Acetic Feeds at the initial input of the system (i.e. before
mixing)
2. Reactor-Out feed to check for VAC produced.
The impact of an increase in the flowrate of oxygen on VAC produced is shown on the Sensitivity
analysis graph below

75
4.0 Verification of Results
The results from the energy and material balances showed some differences when compared for each,
the main difference was seen in the reactor block , the values for the other results were not very far
from the simulation results, however for the energy balances there were some big differences in some
of the streams, the impact of the material balance differences has an effect on the energy balances,
also the simulator takes other factors into account when calculating the energy streams such as heat
loss through block, the specific heat capacities used by Aspen plus are probably more accurate than
the one used from correlations or the other sources used for the energy balances calculation, this could
explain the differences shown when comparing the results with simulation.
The purity obtained from the material balance for the final product VAC was 0.92 compared to the
Apsen plus simulation (0.86), this could explained by material losses encountered by Aspen in block
units such distillation or the separator. It is convincing to obtain a better purity in theory compared to
simulation, as the simulator has a closer interpretation to a real life process.

Simulation of Vinyl Acetate Production

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