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 – 16

th December 2014

Date of Submission

16

th 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 148oC 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= 𝐹𝑖𝑛 ∗ 𝑋𝑖

𝑖𝑛 − 𝐹𝑜𝑢𝑡 ∗ 𝑋𝑖𝑜𝑢𝑡 + Ri = 0 (4)

Fin

= Flow rate of feed in (kmol/h)

Xiin

= molar concentration of component in feed in

Fout

= Flow rate of feed out (kmol/h)

Xiin

= molar concentration of component in feed out

Ri = Product made

𝑑𝑛𝑖

𝑑𝑡= 𝐹𝑖𝑛 ∗ (𝑋𝑉𝐴𝐶

𝑖𝑛 + 𝑋𝐶𝑂2𝑖𝑛 + 𝑋𝐻2𝑂

𝑖𝑛 + 𝑋𝑂2𝑖𝑛 + 𝑋𝐸𝑇𝐻

𝑖𝑛 + 𝑋𝐴𝐶𝐸𝑖𝑛 ) − 𝐹𝑜𝑢𝑡 ∗ (𝑋𝑉𝐴𝐶

𝑜𝑢𝑡 + 𝑋𝐶𝑂2𝑜𝑢𝑡 + 𝑋𝐻2𝑂

𝑜𝑢𝑡 + 𝑋𝑂2𝑜𝑢𝑡 + 𝑋𝐸𝑇𝐻

𝑜𝑢𝑡

+ 𝑋𝐴𝐶𝐸𝑜𝑢𝑡 + 𝑅𝑖 (5)

𝐹𝑖𝑛 = 𝐹𝑜𝑢𝑡 = 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







𝑜𝑢𝑡

+ 𝑋𝐴𝐶𝐸𝑜𝑢𝑡) (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= 𝐹𝑖𝑛 ∗ 𝐶𝑃𝑥

𝑖𝑛 − 𝐹𝑜𝑢𝑡 ∗ 𝐶𝑃𝑥𝑜𝑢𝑡 + 𝑄 (7)

Fin

= Flow rate of feed in (Kmol/h)

CPxin

= Specific heat capacity of component in feed in (Kg/Kj.K)

Fout

= Flow rate of feed out (Kmol/h)

CPxout

= Specific heat capacity of component in feed out (Kg/Kj.K)

Q = Heat Required by process

𝐶𝑝𝑥𝑚𝑖𝑥𝑖𝑛 = (𝐶𝑝𝑂2

𝑖𝑛 ∗ 𝑋𝑂2𝑖𝑛 + 𝐶𝑝𝐶𝑂2

𝑖𝑛 ∗ 𝑋𝐶𝑂2𝑖𝑛 + 𝐶𝑝𝐸𝑇𝐻

𝑖𝑛 ∗ 𝑋𝐸𝑇𝐻𝑖𝑛 + 𝐶𝑝𝐴𝐶𝐸

𝑖𝑛 ∗ 𝑋𝐴𝐶𝐸𝑖𝑛 + 𝐶𝑝𝐻𝑂2

𝑖𝑛 ∗ 𝑋𝐻𝑂2𝑖𝑛 + 𝐶𝑝𝑉𝐴𝐶

𝑖𝑛 ∗

𝑋𝑉𝐴𝐶𝑖𝑛 ) (8)







𝑖𝑛 ∗


𝑑𝑈

𝑑𝑡= 𝐹 ∗ (𝐶𝑃𝑥,𝑚𝑖𝑥

𝑖𝑛 − 𝐶𝑃𝑥,𝑚𝑖𝑥𝑜𝑢𝑡 ) + 𝑄 (10)

5

𝐹𝑖𝑛 = 𝐹𝑜𝑢𝑡 = 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 × 𝑄𝑘𝐽/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


𝑖𝑛 − 𝐹𝑜𝑢𝑡 ∗ 𝑋𝑖𝑜𝑢𝑡 + 𝑅𝑖 = 0 (11)

Fin

= Flow rate of feed in (Kmol/h)

Xiin


Fout

= Flow rate of feed out (Kmol/h)

Xiin


Ri = Reaction of ethylene

𝑑𝑛𝑖







𝑜𝑢𝑡 +

𝑋𝐴𝐶𝐸𝑜𝑢𝑡) (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


𝑖𝑛 − 𝐹𝑜𝑢𝑡 ∗ 𝐶𝑃𝑥𝑜𝑢𝑡 + 𝑄 + 𝐻𝑅 (13)

Fin

= Flow rate of feed in (kg/h)

CPxin

= Specific heat capacity of component in feed in (kJ/kg.k)

Fout

= Flow rate of feed out (kg/h)

CPxout

= 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 ∗


𝐶𝑝𝑥𝑚𝑖𝑥𝑜𝑢𝑡 = (𝐶𝑝𝑂2

𝑜𝑢𝑡 ∗ 𝑋𝑂2𝑜𝑢𝑡 + 𝐶𝑝𝐶𝑂2

𝑜𝑢𝑡 ∗ 𝑋𝐶𝑂2𝑜𝑢𝑡 + 𝐶𝑝𝐸𝑇𝐻

𝑜𝑢𝑡 ∗ 𝑋𝐸𝑇𝐻𝑜𝑢𝑡 + 𝐶𝑝𝐴𝐶𝐸

𝑜𝑢𝑡 ∗ 𝑋𝐴𝐶𝐸𝑜𝑢𝑡 + 𝐶𝑝𝐻𝑂2

𝑜𝑢𝑡 ∗ 𝑋𝐻𝑂2𝑜𝑢𝑡 + 𝐶𝑝𝑉𝐴𝐶

𝑜𝑢𝑡 ∗

𝑋𝑉𝐴𝐶𝑜𝑢𝑡 ) (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


𝑖𝑛 − 𝐹𝑜𝑢𝑡 ∗ 𝑋𝑖𝑜𝑢𝑡 + 𝑅𝑖 (16)

Fin


Xiin


Fout


Cooler

Reactor

REACT-OUT COOL-OUT

9

Xiin


Ri = Reaction

𝑅𝑖 = 0

𝑑𝑛𝑖




𝑖𝑛 + 𝑋𝐴𝐶𝐸𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑣𝑎𝑝 ∗ (𝑋𝑉𝐴𝐶



𝑖𝑛 +

𝑋𝐴𝐶𝐸𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑙𝑖𝑞 ∗ (𝑋𝑉𝐴𝐶



𝑖𝑛 + 𝑋𝐴𝐶𝐸𝑖𝑛 ) (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



Fin


CPxin

= Specific heat capacity of component in feed in (kj/kg.k)

Fout


CPxout

= Specific heat capacity of component in feed out (kj/kg.k)








𝑖𝑛 ∗


𝐶𝑝𝑥𝑚𝑖𝑥𝑙𝑖𝑞

= (𝐶𝑝𝑂2𝑜𝑢𝑡 ∗ 𝑋𝑂2

𝑜𝑢𝑡 + 𝐶𝑝𝐶𝑂2𝑜𝑢𝑡 ∗ 𝑋𝐶𝑂2

𝑜𝑢𝑡 + 𝐶𝑝𝐸𝑇𝐻𝑜𝑢𝑡 ∗ 𝑋𝐸𝑇𝐻

𝑜𝑢𝑡 + 𝐶𝑝𝐴𝐶𝐸𝑜𝑢𝑡 ∗ 𝑋𝐴𝐶𝐸

𝑜𝑢𝑡 + 𝐶𝑝𝐻𝑂2𝑜𝑢𝑡 ∗ 𝑋𝐻𝑂2

𝑜𝑢𝑡 + 𝐶𝑝𝑉𝐴𝐶𝑜𝑢𝑡 ∗


𝐶𝑝𝑥𝑚𝑖𝑥𝑣𝑎𝑝








𝑑𝑈

𝑑𝑡= ((𝐹𝑖𝑛 ∗ 𝐶𝑃𝑥

𝑖𝑛) − (𝐹𝑜𝑢𝑡 ∗ 𝐶𝑃𝑥𝑜𝑢𝑡1) − (𝐹𝑜𝑢𝑡 ∗ 𝐶𝑃𝑥

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



Fin


Xiin


Fout


Xiin


Ri = Reaction

Heater

liquid

Vapour

12

𝑅𝑖 = 0

𝑑𝑛𝑖

𝑑𝑡=

𝐹𝑖𝑛 ∗ (𝑋𝐶𝑂2𝑖𝑛 + 𝑋𝑂2

𝑖𝑛 + 𝑋𝐸𝑇𝐻𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑐 ∗ (𝑋𝐶𝑂2


𝑜𝑢𝑡 ) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑚 ∗

(𝑋𝐶𝑂2𝑜𝑢𝑡) (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



Fin


CPxin


Fout


CPxout









𝑖𝑛 ∗


















𝑑𝑈



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



Fin


Xiin


Fout


Xiin


Ri = Reaction

𝑅𝑖 = 0

𝑑𝑛𝑖




𝑖𝑛 + 𝑋𝐴𝐶𝐸𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑣𝑎𝑝 ∗ (𝑋𝑉𝐴𝐶



𝑖𝑛

+ 𝑋𝐴𝐶𝐸𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑙𝑖𝑞 ∗ (𝑋𝑉𝐴𝐶



𝑖𝑛

+ 𝑋𝐴𝐶𝐸𝑖𝑛 ) (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



Fin


CPxin


Fout


CPxout


16

Q = Heat required by process (kJ/h)







𝑖𝑛 ∗


















𝑑𝑈


𝑖𝑛) − (𝐹𝑜𝑢𝑡,𝑣𝑎𝑝 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥𝑜𝑢𝑡,𝑣𝑎𝑝

) − (𝐹𝑜𝑢𝑡,𝑙𝑖𝑞 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥𝑜𝑢𝑡,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



Fin


Xiin


Fout


Xiin


Ri = Reaction

𝑅𝑖 = 0

𝑑𝑛𝑖




𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑐 ∗ (𝑋𝐶𝑂2𝑜𝑢𝑡 + 𝑋𝑂2

𝑜𝑢𝑡 + 𝑋𝐸𝑇𝐻𝑜𝑢𝑡 ) − 𝐹𝑜𝑢𝑡,𝑟𝑒𝑚 ∗ (𝑋𝑉𝐴𝐶

𝑜𝑢𝑡

+ 𝑋𝐻2𝑂𝑜𝑢𝑡 ) (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



Fin


CPxin


Fout


CPxout


Q = Heat required by process (kJ/h)







𝑖𝑛 ∗


















19

𝑑𝑈



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



Fin


Xiin


Fout


Xiin

= Concentration of component in feed out

Ri = Reaction

Flash drum

liquid

Vapor

20

𝑅𝑖 = 0

𝑑𝑛𝑖




𝑖𝑛 + 𝑋𝐴𝐶𝐸𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑎𝑞𝑢𝑒𝑜𝑢𝑠 ∗ (𝑋𝑉𝐴𝐶


𝑖𝑛 + 𝑋𝑂2𝑖𝑛 +

𝑋𝐸𝑇𝐻𝑖𝑛 + 𝑋𝐴𝐶𝐸

𝑖𝑛 ) − 𝐹𝑜𝑢𝑡,𝑜𝑟𝑔𝑎𝑛𝑖𝑐 ∗ (𝑋𝑉𝐴𝐶𝑖𝑛 + 𝑋𝐶𝑂2

𝑖𝑛 + 𝑋𝐻2𝑂𝑖𝑛 + 𝑋𝑂2

𝑖𝑛 + 𝑋𝐸𝑇𝐻𝑖𝑛 + 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



Fin


CPxin


Fout


CPxout









𝑖𝑛 ∗


















𝑑𝑈


𝑖𝑛) − (𝐹𝑜𝑢𝑡,𝑎𝑞𝑢𝑒𝑜𝑢𝑠 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥𝑜𝑢𝑡,𝑎𝑞𝑢𝑒𝑝𝑢𝑠

) − (𝐹𝑜𝑢𝑡,𝑜𝑟𝑔𝑎𝑛𝑖𝑐 ∗ 𝐶𝑃𝑥,𝑚𝑖𝑥𝑜𝑢𝑡,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−𝐹𝑒𝑒𝑑𝑖𝑛 +𝐹𝐴𝑐𝑒𝑡𝑖𝑐−𝐹𝑒𝑒𝑑

𝑖𝑛 − 𝐹𝐴𝑞𝑢𝑒𝑜𝑢𝑠𝑜𝑢𝑡 − 𝐹𝑂𝑟𝑔𝑎𝑛𝑖𝑐

𝑜𝑢𝑡 − 𝐹𝐶𝑂2−𝑅𝑒𝑚𝑜𝑣𝑒𝑑𝑜𝑢𝑡 − 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 𝐶𝑝𝑂2𝑖𝑛 + 𝐹𝐸𝑇𝐻−𝐹𝑒𝑒𝑑

𝑖𝑛 𝐶𝑝𝐸𝑡ℎ𝑖𝑛 +𝐹𝐴𝑐𝑒𝑡𝑖𝑐−𝐹𝑒𝑒𝑑

𝑖𝑛 𝐶𝑝𝐴𝑐𝑒𝑡𝑖𝑐𝑖𝑛 − 𝐹𝐴𝑞𝑢𝑒𝑜𝑢𝑠

𝑜𝑢𝑡 𝐶𝑝𝐴𝑞𝑢𝑒𝑜𝑢𝑠𝑜𝑢𝑡 − 𝐹𝑂𝑟𝑔𝑎𝑛𝑖𝑐

𝑜𝑢𝑡

− 𝐹𝐶𝑂2−𝑅𝑒𝑚𝑜𝑣𝑒𝑑𝑜𝑢𝑡 𝐶𝑝𝐶𝑂2

𝑜𝑢𝑡

− 𝐹𝑅𝑒𝑐𝑦𝑐𝑙𝑒𝑜𝑢𝑡 𝐶𝑝𝑅𝑒𝑐𝑦𝑐𝑙𝑒

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 30oC 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 148.5oC 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 158.9oC 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 134oC (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 42.5oC 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 80oC 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 40oC, 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 40.4oC 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 80oC – 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:\ProgramData\AspenTech\Aspen Plus V8.4\Group 6 trial(rec) Filename

C:\Users\BILLYB~1\AppData\Local\Temp\~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

F rom 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 F low 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 F low 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 F low 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 F low 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 F rac 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.

76

5.0 Peer Review