1. BEEF, Inc.
Memorandum
Date: October 16, 2015
To: Drs. James E. Maneval, Elif E. Miskioglu, and Ryan C. Snyder
Project Supervisors, Process Engineering Department
From: Andrew J. Fox, Alexander Kempf, and Joel Toro
CHEG 400 Senior Engineers, Process Engineering Department
Subject: PR3.A9: Project Report 3 Full Draft
Abstract
Our team has continued in the process design and has determined the complete structure of our
design. We have developed a series of goals for each separation section and determined the
necessary separation units to accomplish such a separation. Each unit has been thoroughly
designed and investigated to ensure the achievement of our overall goals. Using a series of
heuristics, the size necessary of each unit was determined and the capital investment and
operating costs of using each unit was calculated. By combining each unit, a complete process
flow diagram was created and evaluated to confirm the technical feasibility of our design. The
financial feasibility of our process was then analyzed determined using net present value and
found to be $18,227,945. Process improvements, mainly energy integration, were evaluated and
found to improve the profitability potential by raising the net present value to $18,972,634.
Based on our analysis, we recommended that our design be implemented.

2.
Background
The Marcellus formation, known colloquially as Marcellus shale, is a sedimentary rock
formation, running from New York to West Virginia. Over millions of years, organic material
trapped in the rock decomposed and, due to the unique conditions within the shale, created an
estimated 1 trillion cubic feet of natural gas. The natural gas found within the formation was
thought to be unobtainable, until the development of hydraulic fracturing, or “fracking”, made
extraction of the gas possible. This natural gas, composed primarily of methane with other trace
components, represents a new source of chemical resources that could potentially be used as a
raw material. Our team has sought to use this potential chemical resource to create a product that
is economically viable and competitive.
Our team evaluated the technical and economic feasibility of the production of chemicals from
Marcellus­shale natural gas. Through our previous investigations in PR1 and PR2, we have
developed an appropriate base for further investigation. We have determined that we would
produce chloromethane by chlorinating methane using chlorine gas in a continuous method
operation. After identifying the destinations of each chemical within our process, an
input­output structure of our block flow diagram was created with a material balance. Using this,
the annual profit potential for our process was determined and used to select various operating
conditions, leading to the development of the recycle structure of our process, including reactors
and necessary separations sections.
In this memo, our team will complete stages 5 and 6. In Key Finding 1, we develop the specifics
of the separation section, developing the size and design variables necessary to achieve our
separation goals. In Key Finding 2, we consider alternative design possibilities and examine an
energy integration analysis. Finally, we summarize the results and make recommendations for
completing the project.

3.
Key Finding I: Separation
We developed a base case design for our system that is based on the results of our PR2 block
flow diagram. The base case process flow diagram is shown in Figure 1 on the following page.
Throughout this flow diagram there are many separation units that are necessary. Key Finding 1
aims to address these separation units throughout our process. These units aim to separate out the
necessary components that we need either for a relatively pure feed into our reactor, to separate
out any components that can be recycled, or to separate out our products. We have broken the
units down into three different sections that each focus on one of these goals. Each section
contains the individual separator units, their specific goals, and the key data necessary to run
each section.
Separator Section 1:
Figure 2. Process flow diagram of separation section 1.
S1 Cooler
The flash drum works by separating a gaseous stream from a liquid stream based on the boiling
points of the components. The key data for performing this separation is using the boiling points
of the components of the natural gas stream. As shown below in Table 1, there is a sharp divide
between the boiling points of carbon dioxide and methane (­60.86 and ­141.2°C respectively).
Using this information, the separation can be conducted using a flash drum instead of other
separation methods, where a temperature gradient is not required. The separation will retain a
small amount of nitrogen in the stream, which is acceptable as the 99% purity of methane can
still be achieved and the nitrogen does not affect the reaction.
Component BP (°C)
Ethane ­57.78
Carbon Dioxide ­60.86
Methane ­141.2
Nitrogen ­181.4
Table 1:​ The boiling points of each component in the original natural gas stream at 60 psia.

4.
Figure 1: A process flow diagram showing the initial base case design of the process.

5.
In order to obtain the required temperature for this separation, a working fluid must be used that
can reach these temperatures. We decided to use liquid nitrogen which is able to reach a
temperature of 195.8°C at 1 atm, allowing it to fully cool our natural gas stream. We would
implement a refrigeration unit on­site to keep recycling the cryogenic nitrogen for cooling
throughout not just this cooling, but others throughout our process as well. The refrigeration flow
diagram is shown below in Figure 3. This is blown­up version of the S1 Cooler unit in our PFD.
Figure 3. Process flow diagram for the refrigeration system of S1 Cooler.
Streams 4.1­4.4 contain the liquid nitrogen while streams S1.1 and S1.2 contain our natural gas
feed. In the evaporator the natural gas stream enters at 25°C and exits at ­139°C while liquid
nitrogen enters at ­195.8°C to cool the natural gas stream. As the nitrogen goes through its
refrigeration process, it evaporates into a vapor, compresses to a higher pressure, is cooled with a
another stream of working fluid which turns it into a liquid, expands to atmospheric pressure to
its saturation temperature of ­195.8°C and repeats the cycle. In order to determine the cost of
implementing this we looked at the Carnot efficiency of this mechanical refrigeration system
which can be expressed by the reversible coefficient of performance. After determining the
temperatures that we would need for the condenser and evaporator we used HYSYS to model
this system as best as it could. We were able to get an evaporator heat load from HYSYS and
calculate the work required for this cycle. Knowing this we were able to take our value of
$0.07/hr of electricity and calculate an annual cost of $14,216 for this refrigeration unit. The
detailed calculations for this are located in Appendix D.
S1.2 ­ Flash drum
We have chosen to use a flash drum as our first separation unit in the process. The natural gas
stream contains methane, ethane, nitrogen, and carbon dioxide. The drum aims to separate our
natural gas feed stream into two streams, one with a high purity of methane while the other
contains the primary unwanted component of ethane that would react in our reactor, creating
undesirable chloroethanes. A list of the main goals for this separation is below.

6.
Goals:
● Achieve a 99% purity of methane in the methane stream
● Limit amount of ethane entering the methane stream
● Minimize the amount of methane in the waste stream
These goals were met by this separation process as shown in the stream table below. All of the
methane was able to be separated into stream S1.3 from the inlet stream of S1.2 with a limited
amount of ethane entering that stream.
Stream Table for S1.2 – Flash Drum: ​All values in kg/hr.
Stream ID: S1.2 S1.3 S1.4
Total 225.2000 210.2000 14.9800
CH4 204.2261 204.2261 0.0000
C2H6 16.1175 1.9789 14.1386
N2 1.8769 1.8769 0.0000
CO2 2.9486 2.1079 0.8408
We found the flash drum to have a volume of 0.4m​3​
with a base module cost of $30,125.
Calculations for determining these values is found in Appendix D. The total cost for this
separation section includes the costs of the flash drum, cooler and operating cost of the
refrigeration cycle. The capital cost is $38,383 with an operating cost of $14,216 per year.
After this separation unit, the pure methane stream is mixed with chlorine gas and sent to the
reactor. After the reactor, the product stream exits at a temperature of 440°C and enters
separation section 2.

7.
Separator Section 2:
Figure 4. Process flow diagram of separation section 2.
S2 Cooler
The product stream exits the reactor at 440°C and needs to be cooled before it enters the first
distillation column. This is because we need to get the stream to a temperature where we can get
a separation of our chloromethane products in stream S2.s7 and recycle components in stream
S2.s3. This cooler uses the same refrigeration cycle system as the S1 Cooler with liquid nitrogen
as the working fluid. The process flow diagram is depicted below in Figure 5.
Figure 5. Process flow diagram for the refrigeration system of S2 Cooler.

8.
Streams 6.1­6.4 contain the liquid nitrogen while streams S2.s1 and S2.s2 contain our post
reactor products stream. In the evaporator the products stream enters at 440°C and exits at ­50°C
while liquid nitrogen enters at ­195.8°C to cool the stream. In order to determine the cost of
implementing this, the Carnot efficiency of this mechanical refrigeration system was again
calculated from the same equations used in the S1 Cooler section. After determining the
temperatures that we would need for the condenser and evaporator we used HYSYS to model
this system as best as it could. We were able to get an evaporator heat load from HYSYS and
calculate the work required for this cycle. Knowing this we were able to take our value of
$0.07/hr of electricity and calculate an annual cost of $204,735 for this refrigeration unit.
S2.2 ­ Distillation column
We chose to use a distillation column for the separation that occurs after our reactor since we
require a temperature gradient throughout to get pure products in the distillate and bottoms
streams.
Goals:
● Separate the chlorinated methane products from the remaining components
As seen in the stream table below the chlorinated methane products are separated from stream
S2.s2 into S2.s7 while there is some chloromethane that goes into stream S2.s3. This stream is
recycled back into the reactor which does not go to waste and ends up being part of our final
products streams.
Stream Table for S2.2 – Distillation Column:​ All values in kg/hr.
Stream ID: S2.s2 S2.s3 Flare S2.s7
Total 1759.6755 1025.1873 1.0465 733.4417
CH4 413.1240 412.6481 0.4759 0.0000
C2H6 9.9877 9.9876 0.0000 0.0001
N2 11.8986 11.3281 0.5705 0.0000
CO2 2.1079 2.1078 0.0000 0.0001
Cl2 66.9121 15.8493 0.0000 51.0628
CH3Cl 349.4269 33.9080 0.0000 315.5189
CH2Cl2 259.5216 0.0001 0.0000 259.5215
CHCl3 85.1140 0.0000 0.0000 85.1140
CCl4 15.6690 0.0000 0.0000 15.9960
HCl 545.9137 539.3583 0.0001 6.5553
From our HYSYS simulation we were able to size this distillation column as well as the other
distillation columns that are present throughout separation section 3. We found this distillation
column to have a volume of 7.2 m​3​
with 13 trays. The bare module costing calculations can be
found in Appendix D, where we calculated a cost of $170,365.

9.
Additionally, we looked at sizing the condenser and reboiler for the distillation column along
with calculating the costs for each. We plan on using liquid nitrogen for the condenser and
saturated steam for the reboiler. We found that the reboiler surface area is 1.2m​2​
, and the
condenser is 3.55m​2​
. From these values we calculated the cost of each of these and found them to
be a bare module cost of $7,008 for the condenser, and $4,145 for the reboiler. These
calculations can be found in Appendix D.
S2.2 ­ Absorption column
An absorption column is necessary here to watch out as much of the hydrochloric acid in the
recycle stream as possible. This is simply because we do not want HCl to get into the reactor at
high concentrations. The goals for this section are listed below.
Goals:
● Remove as much HCl as reasonably possible from waste stream in order to recycle the
valuable reactant.
● Achieve a purity of 37% hydrochloric acid in the outlet stream
As seen in the stream table below for the absorption column we were actually able to
successfully separate out all of the HCl from stream S2.s4 into stream S2.s5 with stream S2.p1 as
the recycle stream.
Stream Table for S2.2 – Absorption Column:​ All values in kg/hr.
Stream ID: S2.s4 S2.fH2O S2.s5 S2.p1
Total 1025.1874 749.0805 478.7566 1295.5112
CH4 412.6481 0.0000 412.4831 0.1651
C2H6 9.9876 0.0000 8.9010 1.0867
N2 11.3281 0.0000 11.1718 0.1563
CO2 2.1078 0.0000 0.0000 2.1078
Cl2 15.8493 0.0000 12.2927 3.5566
CH3Cl 33.9080 0.0000 33.9080 0.0000
HCl 539.3583 0.0000 0.0000 539.3583
H2O 0.0000 749.0805 0.0000 749.0805
As shown in Table 3, the solubility of hydrochloric acid is much higher than those of the
remaining components, followed by the chlorinated components and carbon dioxide. Because of
this we were able to wash the stream with water where the HCl had a higher affinity for and
would separate out into the water stream instead of the recycle stream.

10.
Component
Solubility
(g/L)
Solubility
(mol/L)
HCl 720 19.7
CH2Cl2 17.5 0.21
CHCl3 8.09 0.07
Chlorine Gas 7.2 0.1
CHCl3 5.325 0.11
Carbon
Dioxide 1.45 0.033
CCl4 0.81 0.0053
Ethane 0.0568 0.0019
Methane 0.0227 0.0014
Nitrogen 0.019 0.00068
Table 3. The solubility of each component of the product stream from the reactor in water
The resulting stream of mainly hydrochloric acid and water is actually a higher concentration of
37% HCl, so we mixed more water with the stream after separation to achieve this concentration
so it could be sold.
In sizing this unit we looked at equilibrium data for hydrochloric acid and water. We were able
to determine a height of the column to be 4.177m. We also needed to know the diameter in order
to get the volume of the unit, which was calculated based on the flow of our gas through the
system and heuristics from Turtin. We calculated a diameter of 0.36m. From this we calculated a
volume for the absorption column to be 0.4m​3​
. The bare module cost was then calculated to be
$26,623. Calculations for determining the height, diameter, and cost of the column are found in
Appendix D.
From HYSYS we were able to determine a flow of water that we needed for this separation as
well, which adds to the operating costs of this unit. At a flow rate of 50.58 kgmol/hr of H​2​O with
a price of $0.5735/100 gallons of H​2​O, we determined an annual cost of $23.84.

11.
Separation Section 3
The third and final separation section of our process includes three distillation columns. We
chose to use distillation columns because we need to have each stream come out at a high
enough purity of our products to be able to be sold. A PFD for this section is shown below along
with the goals for each of the separate distillation columns.
S3.1 ­ Distillation column
The purpose of this distillation column is to further separate out our products form any
undesirable components. These include Cl​2​ and HCl. This distillation column separates these
components into stream S3.s1, which is a waste stream. The costing for this waste stream can be
found in Key Finding 2.
Goals:
● Further separate out the chlorinated methane products from remaining components, specifically
chlorine
S3.2 ­ Distillation column
The purpose of this distillation column is to obtain our primary product, chloromethane, in a pure
stream while the rest of our products are sent to the S3.3 – Distillation Column.
Goals:
● Separate out chloromethane from the product stream
● Achieve a chloromethane purity of 99%
S3.3 ­ Distillation column
The purpose of this distillation column is to obtain one of our valuable byproducts of
dichloromethane as a pure stream, while the other stream contains any waste that is leftover from
our process. The costing for the disposal of this waste stream is found in Key Finding 2.
Goals:
● Separate out dichloromethane from the product stream
● Achieve a dichloromethane purity of 99.8%

12.
Based on the stream tables for each of the distillation columns shown below, we were able to
meet our goals of separating out high product purities. In S3.1 – Distillation Column we were
able to separate out all of the hydrochloric acid (HCl) and most of the chlorine (Cl2) from stream
S2.s7 into stream S3.s1, where stream S3.s2 contains our valuable products. For S3.2 –
Distillation Column we were able to separate out most of the chloromethane (CH3Cl), where
stream S3.s3 has a chloromethane purity of 99.2%, which meets our goal. For S3.3 – Distillation
Column we were able to separate out most of the dichloromethane (CH2Cl2), where stream
S3.s5 has a dichloromethane purity of 99.8%, which also meets our goal.
Stream Table for S3.1 – Distillation Column:​ All values in kg/hr.
Stream ID: S2.s7 S3.s1 S3.s2
Total 733.4417 57.2714 676.1703
Cl2 51.0628 49.8812 1.1815
CH3Cl 315.5189 0.8347 314.6842
CH2Cl2 259.5215 0.0000 259.5215
CHCl3 85.1140 0.0000 85.1140
CCl4 15.6690 0.0000 15.6690
HCl 6.5553 6.5553 0.0000
Stream Table for S3.2 – Distillation Column:​ All values in kg/hr.
Stream ID: S3.s2 S3.s3 S3.s4
Total 676.1703 316.8983 359.2720
Cl2 1.1815 1.1810 0.0005
CH3Cl 314.6842 314.4740 0.2102
CH2Cl2 259.5215 1.2432 258.2783
CHCl3 85.1140 0.0000 85.1140
CCl4 15.6690 0.0000 15.6690
Stream Table for S3.3 – Distillation Column:​ All values in kg/hr.
Stream ID: S3.s4 S3.s5 S3.s6
Total 359.2720 254.7190 104.5531
Cl2 0.0005 0.0005 0.0000
CH3Cl 0.2102 0.2102 0.0000
CH2Cl2 258.2783 254.2844 3.9940
CHCl3 85.1140 0.1996 84.9144
CCl4 15.6690 0.0243 15.6447
The key data to performing this separation in the distillation columns are the boiling points of the
components of the reactor products, and in the absorption column are the solubility of these
components in water. As shown below in Table 2, the boiling points of all of the chlorinated
methane products are higher than those of the remaining components. In addition, the boiling
points of the chloromethanes increase as the degree of chlorination increases, with
chloromethane having the lowest boiling point temperature and carbon tetrachloride having the
highest boiling point temperature.

13.
Component BP (°C)
CCl4 77.08
CHCl3 62.05
CH2Cl2 39.95
CH3Cl ­24.01
Chlorine Gas ­33.15
HCl ­84.78
Carbon Dioxide ­86.58
Ethane ­88.61
Methane ­161.6
Nitrogen ­195.7
Table 2. The boiling points of each component in the product stream from the reactor
Additionally, we looked at sizing the condensers and reboilers for the distillation columns along
with calculating the costs for each. We plan on using liquid nitrogen for the condensers and
saturated steam for the reboilers. The following table shows the surface area and bare module
cost for the condensers and reboilers for each of the distillation columns in this section. The
calculations that achieved these values can be found in Appendix D.
Surface area (m^2) Bare Module Cost
S2.3­ Condenser 3.63 $7,166
S2.3­ Reboiler 2.4058 $7,915
S2.4­ Condenser 14.89 $5211.5
S2.4­ Reboiler 14.19 $4966.5
S2.5­ Condenser 70.16 $17540
S2.5­ Reboiler 102.2 $30660

14.
Key Finding II: Energy Integration and Process Improvements
After completing the base case design (Figure 1), Team 7 carefully examined how we could decrease the
operating costs of the process. Potential areas for cost reduction including energy integration, waste
stream management, reduction of capital costs and reduction of operating (utility) costs.
We looked at the temperatures of all streams, especially ones right before and after heaters and coolers.
The goal of this examination is to decrease our annual costs and to increase the net present value of the
project.
In the creation of the base case (Figure 1), unit and operating cost were highly considered at all times. The
smallest and cheapest units were used when possible (for example, flash drums were first installed and if
they didn’t work, then a distillation column was used). Since the creation of our base case design was so
iterative, there is not much cost reduction that would be further possible in terms of capital cost.
In order to meet our goals of decreasing the annual costs and to increase the net present value of the
project, we completed the following:
1. Identify and assess improvements with respect to energy use
2. Identify and assess improvements with respect to utility costs
Energy Use
Table 1.1: Shows the inlet and outlet temperatures from the heaters and coolers in the base case.
Stream Inlet temp (C)
Outlet temp
(C)
Heat/Cool load
kJ/hr
S1
Cooler Natural gas feed 25 ­139 ­84000
S1
Heater
HCl and recycle from heat
exchanger 3 45.6 440 930000
S2
Cooler HCl and recycle 440 ­50 ­1400000
S2
Heater
Reactor feed from heat
exchanger 1 ­170 25 800000
Net energy
requirement 246000

15.
Table 1.2: Shows the streams that are crossed in heat exchangers and their inlet and outlet temperatures in
the energy integration case.
Cold stream
Cold
stream
inlet temp
(C)
Cold
stream
outlet
temp (C) Hot stream
Hot
stream
inlet temp
(C)
Hot stream
outlet temp
(C)
Heat
exchanger 1 Reactor feed 44.4 430 Reactor products 430 68
Heat
exchanger 2
HCl and
recycle from
heat
exchanger 3 ­159 ­81.7
Reactor products
from heat
exchanger 1 68 ­50
Heat
exchanger 3
HCl and
recycle ­171 ­159 Natural gas feed 25 ­139
Fired heater 1
Reactor feed
from heat
exchanger 1 430 440
Ethane stream
from natural gas
purification
This
stream is
burned
Exhaust is
vented to
atmosphere
Table 1.2 shows the inlet and outlet stream temperatures and how the heat exchangers reached their
required temperatures. We did not need to use composite curves in this energy integration since there was
enough energy available to eliminate all of the heaters and coolers and replace them with the heat
exchangers. We chose which streams to cross through an iterative process in HYSYS. There are a lot of
very high and very low temperature streams which made crossing them to exchange heat a relatively
simple process. There was only one stream that was not able to be fully heated by crossing it with other
streams, so this was dealt with by heating it with a fired heater which burned an ethane waste stream.
Quantitatively, this can be shown by the net energy input required for all of the heaters and coolers in the
base case. The 2.5x10​5
kJ/hr of energy that is needed (assuming perfect energy integration) is more than
covered by the direct fired heater. In the energy integration case, the direct fired heater is able to deliver
1.1x10​6​
kJ/hr of heat.
All of the heat exchangers and direct fired heaters were added in order to replace heaters and coolers that
were initially in the base case. Making changes in this manner removes the capital cost of the heaters and
coolers, as well as the utility costs of them. However it does add the capital cost of the heat exchanger and
direct fired heater. As shown by the present value of the project in Table 3, this is a profitable
replacement.

16.
Figure 1: A process flow diagram showing the initial base case design of the process.

17.
Figure 2: A process flow diagram showing how streams have been rerouted to cross other streams in an
effort to reduce overall costs.

18.
S1.1­ Heat Exchanger 1:
Figure 3: A portion of Figure 2 showing the streams involved with heat exchanger 1.
Goals: This heat exchanger serves to cool the natural gas feed stream down while heating up the distillate
(HCl and recycle) stream of the first distillation column after the reactor.
This heat exchanger is replacing a heater that would have heated the HCl and recycle stream, as well as a
cooler that would have cooled the natural gas feed stream.
S1.5­ Heat Exchanger 2:
Figure 4: A portion of Figure 2 showing the streams involved with heat exchanger 2.

19.
Goals: The goals of this heat exchanger are essentially to heat the reactor feed stream with the product
stream.
Additional required heating of the reactor feed stream is added by a direct fired heater discussed below.
This replaces a heater that would have heated the reactor feed stream and a cooler that would have cooled
the products stream.
S2.1­ Heat Exchanger 3:
Figure 5: A portion of figure 2(from above) showing the streams involved with heat exchanger 3.
Goals: The goals of this heat exchanger are to further cool the reactor products, or first distillation column
feed, with the distillate stream from the first distillation column.
This heat exchanger replaces a cooler that would cool the feed to the first distillation column, as well as a
heater that would have heated the distillate stream.

20.
S1.6­ Direct Fired Heater 1:
Figure 6: A portion of Figure 2 showing the streams involved with direct fired heater 1.
Goals: This fired heater serves to burn a waste stream in order to decrease waste disposal costs. It also
heats the reactor feed stream the additional 10 degrees that heat exchanger S1.5 was unable to
accomplish.
This unit partially replaces the heater that was needed to preheat the feed stream.

21.
Utilities
In an effort to reduce utility costs, the breakdown of the utilities was calculated below.
Table 2.1: Utility breakdown and percentage of total utility cost for base case.
Unit Utility Cost of utility
Percentage of
total
S1 Cooler
Nitrogen $2 0.00011%
electricity $14,216 1.03287%
S1 Heater Saturated Steam $2,089 0.15176%
Reactor Cooling water $1 0.00005%
S2 Cooler
Nitrogen $15 0.00111%
Electricity $204,735 14.87548%
S2 Heater Steam $101 0.00735%
Compressor Electricity $19,254 1.39898%
S2.1 nitrogen $2 0.00014%
Saturated Steam $488 0.03548%
S2.2 Product water $24 0.00173%
S2.3 nitrogen $2 0.00014%
Saturated Steam $81 0.00586%
S2.4 nitrogen $2 0.00014%
Saturated Steam $421 0.03057%
S2.5 nitrogen $2 0.00014%
Saturated Steam $2,089 0.15175%
Water pipe service charge Water­ general $38,664 2.80922%
Waste Waste disposal $1,094,138 79.49711%
Net Utilities $1,376,325

22.
Table 2.2: Utility breakdown and percentage of total utility cost for energy integration. Utilities that are
less than 0.5% of the total are not shown here. This includes nitrogen and steam costs.
Unit Utility Annual cost of utility
Percentage of
total
Waste
Waste
disposal $1,094,138 94.96509%
Water pipe service charge Water­ general $38,664 3.35582%
Compressor Electricity $19,254.48 1.67118%
Net Utilities $1,152,148
As shown by Tables 2.1 and 2.2, the energy integration did decrease the net utility cost by approximately
$200,000. However, as shown in Table 2.2, the primary utility cost in this design is the waste disposal. As
discussed above in the Waste Stream Management section, it would be very difficult to separate the waste
stream into valuable byproducts, adding multiple large distillation columns as well as adding the utility
costs for those units.
Net Impact of Energy Integration:
The costing information described below only represents the changes that occurred in capital costs and
utility costs due to the energy integration, as well as the overall project costs (see appendix C). To see a
detailed breakdown of how the calculations were carried out, please refer to Project Report 2.
Table 3: Cost comparison between the base case and the energy integrated case. This table only shows the
utilities that had a significant change between the base case and energy integration case.
Base case Energy integration
FCI $7,299,822 $8,417,570
Steam $5,167 $59
Electricity $238,205 $19,254
Water $38,688 $38,688
Net Utilities $1,376,224 $1,152,148
NPV $18,227,945 $18,972,634
In an effort to increase the overall efficiency of our plant, decrease our utility costs and therefore make the
overall process more profitable, Team 7 has closely examined ways to save energy, decrease the amount
of utilities used, decrease waste disposal costs, and decrease the capital cost of the equipment. We have
accomplished this by using different temperature streams to heat or cool each other, rather than paying for
a heater and a cooler. We also utilized direct fired heaters to burn our waste streams and provide further
heating.
As shown by Table 3, the net present value of the project increases the capital cost of the units, but it
decreases the utility costs. Additionally, it decreases electricity consumption which is more
environmentally friendly. Overall there is a relatively small change in the net present value of the project,

23.
and increase of roughly 4%. This means that the project is more profitable after the energy integration, as
well as more environmentally friendly, meeting our goal of becoming more profitable.
Waste Stream Management
Team 7 also examined our waste streams, and how they can best be dealt with. From the base case design
(Figure 1), there is stream (S1.s4) that will be sent to a flare, and two streams (S3.s1 and S3.s6) that need
to be disposed of as hazardous waste. They could not be burned since they are primarily chloromethane.
The flared streams provide energy to the fired heaters currently in our process. As for the waste streams,
the only alternative option would be to separate them further and sell the products. However, in order to
purify the stream to a high enough purity for sale, two additional distillation columns that were larger than
any other columns in the system would have to be added. Since these products are not our primary
product and since they would have had an enormous capital cost, we chose to simply dispose of the waste.

24.
Conclusions
Our team has developed the structure of the separations necessary for the chlorination of methane using
chlorine gas to produce chloromethane. Our key findings are as follows:
1. Our base case has a net present value of $18,227,945 and total capital investment of $8,077,467
2. Our energy integrated design has a net present value of $18,972,634 and total capital investment
of $9,306,990
Recommendations
We recommend the following actions to be taken based on these findings:
● Recommendation 1: The series of separations detailed in Key Finding 1 be used to successfully
separate and purify our primary product and valuable byproducts
● Recommendation 2: The energy integration detailed in Key Finding 2 be used to reduce the
overall cost and increase the net present value.
● Recommendation 3: Our process for the production of chloromethane be implemented

25.
References
1. Fair, J. Henry. "Natural Gas Drilling in the Marcellus Shale Uncovered: Community,
Environment, and Law." ​Environmental Studies Capstone ::​. Swarthmore College, n.d. Web. 16
Sept. 2015. <http://www.swarthmore.edu/environmental­studies­capstone>.
2. "History and Background ­ What Is Marcellus Shale?" ​Natural Gas Development​. Cornell
University, n.d. Web.
<http://www2.cce.cornell.edu/naturalgasdev/documents/pdfs/history%20and%20background1.pd
f>.
3. "The Marcellus Shale, Explained." ​StateImpact​. NPR, n.d. Web. 16 Sept. 2015.
<https://stateimpact.npr.org/pennsylvania/tag/marcellus­shale/>.
4. United States of America. Department of Conservation and Natural Resoruces. ​Marcellus Shale​.
PA DCNR ­ Geology, n.d.
<http://www.dcnr.state.pa.us/topogeo/econresource/oilandgas/marcellus/marcellus_faq/marcellus
_shale/index.htm>
5. United States of America. Ohio Department of Natural Resources. Ohio Geological Study. ​The
Marcellus Shale Play: Geology, History, and Oil & Gas Potential in Ohio​. By Chris Perry and
Larry Wickstrom. Ohio Geological Survey, n.d. Web.
<http://geosurvey.ohiodnr.gov/portals/geosurvey/Energy/Marcellus/The_Marcellus_Shale_Play_
Wickstrom_and_Perry.pdf>.

26.
Appendix A: Unit Sizing:
Unit Size Cp B1 B2
F
m Cbm = Cp*(B1+B2Fm)
Reactor 10 liters $128,000 1.49 1.52 9.4 $2,019,584
S2 Compressor 42.1 hp $12,800 9.8 $125,440
Column
Diamete
r (m)
Numbe
r of
trays
Colum
n
height
(m)
Colum
n
volume
(m^3)
Base
price
Pressur
e factor
Materi
al
Factor B1 B2
Bare
module
cost
S1­ Flash
Drum ­
Vertical 0.4572 N/A 2.515 0.4 $2,192 0.672 9.4
2.2
5
1.8
2 $30,125
S2.1­
Distillatio
n column 1.19 13 6.5 7.2 $9,235 0.947 9.4
2.2
5
1.8
2
$170,36
5
S2.2­
Absorptio
n column 0.360 N/A 4.177 0.4 $2,226 0.568 9.4
2.2
5
1.8
2 $26,623
S2.3­
Distillatio
n column 1.193 61 30.5 34.1
$27,59
8 0.948 9.4
2.2
5
1.8
2
$509,67
0
S2.4­
Distillatio
n column 1.193 33 16.5 18.4
$17,41
5 0.835 9.4
2.2
5
1.8
2
$287,93
1
S2.5­
Distillatio
n column 1.193 73 36.5 40.8
$31,78
4 0.806 9.4
2.2
5
1.8
2
$509,67
6

27.
Surface area
(m^2)
Cp (from graph in
Turtin) B1 B2
Bare Module
Cost
S1 Natural Gas
Cooler 5.02 2510 1.74 1.55 $8,258
S1 Reactor
Preheater 31.81 11133.5 1.74 1.55 $36,629
S2 Products Cooler 49.61 17363.5 1.74 1.55 $57,126
S2 Heater 48.4 16940 1.74 1.55 $55,733
S2.1­ Condenser 3.55 2130 1.74 1.55 $7,008
S2.1­ Reboiler 1.2 1260 1.74 1.55 $4,145
S2.3­ Condenser 3.63 2178 1.74 1.55 $7,166
S2.3­ Reboiler 2.4058 2405.8 1.74 1.55 $7,915
S2.4­ Condenser 14.89 5211.5 1.74 1.55 $17,146
S2.4­ Reboiler 14.19 4966.5 1.74 1.55 $16,340
S2.5­ Condenser 70.16 17540 1.74 1.55 $57,707
S2.5­ Reboiler 102.2 30660 1.74 1.55 $100,871
Unit Cost Summary
Reactor $2,019,584
S1­ Flash Drum ­ Vertical $30,125
S1 Natural Gas Cooler $170,365
S1 Reactor Preheater $26,623
S1 Products Cooler $509,670
S2 Heater $287,931
S2 Compressor $125,440
S2.1­ Distillation column $181,518
S2.2­ Absorption column $26,623
S2.3­ Distillation column $524,751
S2.4­ Distillation column $321,416
S2.5­ Distillation column $668,254
Total unit costs $4,055,457

28.
Fixed capital investment (FCI) $7,299,822 =total unit cost * 1.2 * 1.5
Cl2
kgmol/hr 15.76 from HYSYS
$/kgmol $15 from Alibaba.com
7 day supply $39,715 =kgmol/hr * $/kgmol * 24hr * 7 days
Natural Gas
kgmol/hr 13.40 from HYSYS
$/kgmol $0.0017 from Alibaba.com
7 day supply $4 =kgmol/hr * $/kgmol * 24hr * 7 days
Feed inventories (FI) $39,719
=7 day supply Cl2 + 7 day supply natural
gas
Working capital (WC) $777,645 =1.2 * FI + 0.1 * FCI

29.
Appendix B: Utilities
Heat
exchanger
utilities
Tf C
kgmol
/hr
Heat
flow
kJ/hr
Needed
flow of
steam(for
reboilers)
nitrogen
(for
condensers
) g/s
mol/mont
h H2O
gal/mont
h H2O
$/100 gal
(condenser
)
$/1000 kg
(reboiler)
Annual
cost
S1 Natural
Gas Cooler ­139 13.4 ­83540 77.7 55944 266 0.5735 $1.53
S1 Reactor
Preheater 440 53.26 927900 2033.33 1463998 6962 30
$2,088.6
6
S2 Cooler ­50 53.26
­14180
00 778 560160 2664 0.5735 $15.28
S2 Heater 25 42.2 800600 98.53 70942 337 30 $101.21
S2.1­
Condenser ­170.2 42.2
­74000
0 100 72000 342 0.5735 $1.96
S2.1­
Reboiler 11.6 11.2 107000 475.4 342288 1628 30 $488
S2.3­
Condenser ­43.5 0.9
­63000
0 100 72000 342 0.5735 $2
S2.3­
Reboiler 19.3 10.1 640000 78.55 56556 269 30 $81
S2.4­
Condenser ­24.1 6.3
­33000
00 100 72000 342 0.5735 $2
S2.4­
Reboiler 65.8 3.9
333000
0 409.61 294919 1403 30 $421
S2.5­
Condenser 39.8 2.1
­20000
0 100 72000 342 0.5735 $2
S2.5­
Reboiler 67.6 1.78 200000 2033.23 1463926 6962 30 $2,089

30.
S2.2 and HCl utilities
kgmol/hr H2O 50.58
mol/month H2O 36418
gal/month H2O 173
$/100 gal 0.5735
Annual cost $23.84
Waste streams
Flow rate­ kgmol/hr
Waste1­ kgmol/hr 0.9
Waste2­ kgmol/hr 0.86
Waste1­ kg/hr 57.27
Waste2­ kg/hr 104.6
Annual production­ ton/year 1563
Disposal cost­ $/ton 700
Annual disposal cost $1,094,138

31.
Unit Utility Costs
Unit Utility
Cost of
utility
S1 Cooler
Nitrogen $1.53
electricity $14,216
S1 Heater
Saturated
Steam $2,088.66
Reactor Cooling water $0.65
S2 Cooler
Nitrogen $15.28
Electricity $204,735
Compressor Electricity $19,254
S2.1 nitrogen $1.96
Saturated
Steam $488
S2.2 Product water $23.84
S2.3 nitrogen $2
Saturated
Steam $81
S2.4 nitrogen $2
Saturated
Steam $421
S2.5 nitrogen $2
Saturated
Steam $2,089
Water pipe service charge Water­ general $38,664

32.
Appendix C
Table 2­ This table show the total unit costs for the base case.
Reactor $2,019,584
S1­ Flash Drum ­ Vertical $30,125
S1 Natural Gas Cooler $170,365
S1 Reactor Preheater $26,623
S1 Products Cooler $509,670
S2 Heater $287,931
S2 Compressor $125,440
S2.1­ Distillation column $181,518
S2.2­ Absorption column $26,623
S2.3­ Distillation column $524,751
S2.4­ Distillation column $321,416
S2.5­ Distillation column $668,254
Total unit costs $4,055,457
Table 3­ This table show the total unit costs for the energy integrated case.
Reactor $2,019,584
Fired_heater1 $740
Heat_recovery1 $58,957
Heat_recovery2 $12,015
Heat_recovery3 $626,913
Heater $80,092
S2 Compressor $125,440
S2.1­ Distillation column $804,285
S2.2­ Absorption column $26,623
S2.3­ Distillation column $520,981
S2.4­ Distillation column $312,991
S2.5­ Distillation column $583,723
Total unit costs $4,676,428

33.
Table 4­ This table shows the utility costs of each unit in the base case.
Unit Utility Cost of utility
S1 Cooler
Nitrogen $2
electricity $14,216
S1 Heater Saturated Steam $2,089
Reactor Cooling water $1
S2 Cooler
Nitrogen $15
Electricity $204,735
Compressor Electricity $19,254
S2.1 nitrogen $2
Saturated Steam $488
S2.2 Product water $24
S2.3 nitrogen $2
Saturated Steam $81
S2.4 nitrogen $2
Saturated Steam $421
S2.5 nitrogen $2
Saturated Steam $2,089
Water pipe service charge Water­ general $38,664
Waste Waste disposal $1,094,138
Net Utilities $1,376,224

34.
Table 5­ This table shows the utility costs of each unit in the energy integrated case.
Unit Utility Cost of utility
Reactor Cooling water $0.65
Compressor Electricity $19,254.48
S2.1 nitrogen $1.96
Saturated Steam $9.34
S2.2 Product water $23.84
S2.3 nitrogen $1.96
Saturated Steam $1.54
S2.4 nitrogen $1.96
Saturated Steam $8.04
S2.5 nitrogen $1.96
Saturated Steam $39.93
Water pipe service charge Water­ general $38,664
Waste Waste disposal $1,094,138
Net Utilities $1,152,148

35.
Appendix D: Unit Sheets
Project: Chloromethane synthesis from methane
Unit Name/ID: S1.1 Flash Drum
Unit Type: Flash Drum (vertical)
Sizing Basis: Volume, m​3
Unit Diagram
Stream Table: ​All values in kg/hr.
Stream ID: S1.2 S1.3 S1.4
Total 225.2000 210.2000 14.9800
CH4 204.2261 204.2261 0.0000
C2H6 16.1175 1.9789 14.1386
N2 1.8769 1.8769 0.0000
CO2 2.9486 2.1079 0.8408
Flash Drum Size Summary
Using the Souders­Brown equation shown below we were able to size our drum.
k) V = (
√ ρV
ρ −ρL V
where:
V = maximum allowable vapor velocity (m/s)
= liquid density, kg/m​3
ρL
= vapor density, kg/m​3
ρV
k = 0.107 m/s
From the vapor velocity and knowing the flow rate of the natural gas stream from our HYSYS
model we could get the cross sectional area (A) and diameter (D) of the drum with the following
equations:
A = V
vapor flow rate, in m /s3
D = [ ]π
4(A) 0.5

36.
The height of the vessel (h) was then calculated with the equation below assuming a liquid
residence time of 5 minutes and that the liquid level is about the midpoint of the vessel height:
h = A
2 (5 min) (liquid flow rate, in m /min)* *
3
With the height and diameter of the vessel the volume (V) can be calculated as follows:
( ) hV = π 2
D 2
We calculated the flash drum to have a volume of 0.4m​3​
.
Flash Drum Cost Summary
We found the flash drum to have a volume of 0.4m​3​
. Looking in Table A.1 of Turton we
determined the purchased equipment cost at ambient operating pressure and using carbon steel
construction with the following equation:
C log (A) [log (A) ]log10
0
p = K1 + K2 10 + K3 10
2
where:
A = 0.4m​3
K​1​ = 3.4974
K​2​ = 0.4485
K​3​ = 0.1074
The bare module cost (C​BM​) for this vessel has to account for its material factors (F​M​) and
pressure factor (F​P​).
(B F F )CBM = C0
p 1 + B2 M P
FP,vessel = 0.0063
+0.00315(P+1)D
2[850−0.6(P+1)]
where:
D = 0.46m
B​1​ = 2.25
B​2​ = 1.82
F​M​ = 9.4
P = 3.01 barg
The base module cost (C​BM​) was found to be $30,125.
S1 Cooler Size Summary

37.
Using the following set of equations as well as information from HYSYS, we were able to size
the heat exchanger between liquid nitrogen and our natural gas stream.
A
Q
= dT
+∑
1
hi k
x
pipe
hi = D
k Nui
u .023Re PrN = 0 0.8 0.3
eR = μ
ρvD
r p( )P = C k
μ
Td =
ln ( ) T −Tsat in
T −Tsat out
(T −T )−(T −T )sat out sat in
A = ( )A
Q
Qneeded
where:
i = the liquid nitrogen stream and natural gas stream
k​pipe​ = 400 W/mC
D = 0.1407m
x = 0.0021m
Q​needed​, viscosity ( ), thermal conductivity (k), density ( ). T​in​, and T​out​ data was found from ourμ ρ
HYSYS simulation for each stream.
We calculated a surface area of 5.02m​2​
for this heat exchanger.
S1 Cooler Cost Summary – Capital Cost
Knowing the surface area (A) of the heat exchanger, we used Figure A.5 in Appendix A of
Turtin to find a C​p​
0​
/A($/m​2​
) where:
C0
p =
C0
p
A( )$
m2
* A
To calculate the bare module cost (C​BM​) of this heat exchanger we used the following equation:
(B F F )CBM = C0
p 1 + B2 M P
where:
B​1​ = 1.74

38.
B​2​ = 1.55
F​M​ = 1
F​P​ = 1
We calculated a bare module cost of $8,258 for this heat exchanger.
S1 Cooler Cost Summary – Operating Cost
The condenser in the refrigeration unit that creates the liquid nitrogen used in the heat exchanger
requires power. Therefore the cost of the electricity to run it was calculated. We looked at the
Carnot efficiency of this mechanical refrigeration system which can be expressed by the
reversible coefficient of performance, COP​REV​:
COPREV =
evaporator temperature (T )1
temperature difference between condenser and evaporator (T −T )2 1
ork requiredw = COP
evaporator heat load
where:
T​1​ = ­195.8°C
T​2​ = 0°C
evaporator heat load = 83460 kJ/hr
We calculated a value of 23.2 kW of work required, and were able to take our value of $0.07/hr
of electricity to calculate an annual operating cost of $14,216 for this refrigeration unit.
Total Capital Cost: $38,383
Total Operating Cost: $14,216/year
Project: Chloromethane synthesis from methane
Unit Name/ID: S2.1 Distillation Column

39.
Unit Type: Distillation Column
Sizing Basis: Volume, m​3
Unit Diagram
Stream Table:​ All values in kg/hr.
Stream ID: S2.s2 S2.s3 Flare S2.s7
Total 1759.6755 1025.1873 1.0465 733.4417
CH4 413.1240 412.6481 0.4759 0.0000
C2H6 9.9877 9.9876 0.0000 0.0001
N2 11.8986 11.3281 0.5705 0.0000
CO2 2.1079 2.1078 0.0000 0.0001
Cl2 66.9121 15.8493 0.0000 51.0628
CH3Cl 349.4269 33.9080 0.0000 315.5189
CH2Cl2 259.5216 0.0001 0.0000 259.5215
CHCl3 85.1140 0.0000 0.0000 85.1140
CCl4 15.6690 0.0000 0.0000 15.9960
HCl 545.9137 539.3583 0.0001 6.5553
S2.2 – Distillation Column Size Summary
Column Diameter = 1.19m
Number of Trays = 13
Column Height = 6.5m (assuming the distance between each tray is 0.5m)
olumn V olume π olumn HeightC = * ( )2
Column Diameter 2
* C
Column Volume = 7.2m​3
S2.2 – Distillation Column Cost Summary

40.
We found the distillation column to have a volume of 7.2m​3​
. Looking in Table A.1 of Turton we
determined the purchased equipment cost at ambient operating pressure and using carbon steel
construction with the following equation:
C log (A) [log (A) ]log10
0
p = K1 + K2 10 + K3 10
2
where:
A = 7.2m​3
K​1​ = 3.4974
K​2​ = 0.4485
K​3​ = 0.1074
The bare module cost (C​BM​) for this vessel has to account for its material factors (F​M​) and
pressure factor (F​P​).
(B F F )CBM = C0
p 1 + B2 M P
FP,vessel = 0.0063
+0.00315(P+1)D
2[850−0.6(P+1)]
where
D = 1.19m
B​1​ = 2.25
B​2​ = 1.82
F​M​ = 9.4
P = 3.01 barg
The base module cost (C​BM​) was found to be $170,365.
S2.2 Condenser and Reboiler Size Summary
Using the following set of equations as well as information from HYSYS, we were able to size
the condensers as well as reboilers in the distillation column. For the condensers we used liquid
nitrogen and for the reboilers we used saturated steam.
A
Q
= dT
+∑
1
hi k
x
pipe
hi = D
k Nui
u .023Re PrN = 0 0.8 0.3
eR = μ
ρvD
r p( )P = C k
μ

41.
Td =
ln ( ) T −Tsat in
T −Tsat out
(T −T )−(T −T )sat out sat in
A = ( )A
Q
Qneeded
where:
i = each stream in the condenser or reboiler
k​pipe​ = 400 W/mC
D = 0.1407m
x = 0.0021m
Q​needed​, viscosity ( ), thermal conductivity (k), density ( ). T​in​, and T​out​ data was found from ourμ ρ
HYSYS simulation for each exit stream of either the condenser or reboiler.
From this set of equation we were able to calculate the surface area of the heat exchanger for the
reboiler and condenser in the distillation column. The reboiler surface area is 1.2m​2​
, and the
condenser is 3.55m​2​
.
S2.2 Condenser and Reboiler Cost Summary
Knowing the surface area (A) of the condenser and reboiler, we used Figure A.5 in Appendix A
of Turtin to find a C​p​
0​
/A($/m​2​
) where:
C0
p =
C0
p
A( )$
m2
* A
To calculate the bare module cost (C​BM​) of this heat exchanger we used the following equation:
(B F F )CBM = C0
p 1 + B2 M P
where:
B​1​ = 1.74
B​2​ = 1.55
F​M​ = 1
F​P​ = 1
We calculated a bare module cost of $7,008 for the condenser, and $4,145 for the reboiler.
Total Capital Coat = $181,518
Project: Chloromethane synthesis from methane
Unit Name/ID: S2.2 Absorption Column
Unit Type: Absorption Column

42.
Sizing Basis: Volume, m​3
Unit Diagram
Stream Table:​ All values in kg/hr.
Stream ID: S2.s4 S2.fH2O S2.s5 S2.p1
Total 1025.1874 749.0805 478.7566 1295.5112
CH4 412.6481 0.0000 412.4831 0.1651
C2H6 9.9876 0.0000 8.9010 1.0867
N2 11.3281 0.0000 11.1718 0.1563
CO2 2.1078 0.0000 0.0000 2.1078
Cl2 15.8493 0.0000 12.2927 3.5566
CH3Cl 33.9080 0.0000 33.9080 0.0000
HCl 539.3583 0.0000 0.0000 539.3583
H2O 0.0000 749.0805 0.0000 749.0805
S2.2 ­ Absorption Column Size Summary
The following set of equation was used in order to determine the height of the absorption
column.
x )LM( * − x = ln ( ) (x2 −x2)*
(x1 −x1)*
(x1 −x1)−(x2 −x2)* *
LM(1 )− x *
= ln ( ) (1−x1)
(1−x1 )*
(1−x1)−(1−x1 )*
.6634Kya = 0 * L0.82
TUN = (x1−x2)
(x −x)LM*
TUH = L
K (1−x) LMya*
*

43.
TU TUz = H * N
where:
x1 = 0.262
x2 = 0
x1​*​
= 0.9693
x2​*​
= 0.9693
L = 41.6 kgmol/hr
We found that the height of the column (z) is 4.177m.
We then used the following equation to determine the diameter of the column.
D = √ π
( )3600
gas flow, in kg/hr
average gas density, in 1kg
m3* s
m
* 2
The diameter of the column was calculated to be 0.36m.
We then used the following equation to determine the volume of the absorption column.
πV = * ( )2
D 2
* z
The column has a volume of 0.4 m​3​
.
S2.2 ­ Absorption Column Capital Cost Summary
We found the absorption column to have a volume of 0.4m​3​
. Looking in Table A.1 of Turton we
determined the purchased equipment cost at ambient operating pressure and using carbon steel
construction with the following equation:
C log (A) [log (A) ]log10
0
p = K1 + K2 10 + K3 10
2
where:
A = 0.4m​3
K​1​ = 3.4974
K​2​ = 0.4485
K​3​ = 0.1074
The bare module cost (C​BM​) for this vessel has to account for its material factors (F​M​) and
pressure factor (F​P​).
(B F F )CBM = C0
p 1 + B2 M P
FP,vessel = 0.0063
+0.00315(P+1)D
2[850−0.6(P+1)]

44.
where
D = 0.36m
B​1​ = 2.25
B​2​ = 1.82
F​M​ = 9.4
P = 3.01 barg
The base module cost (C​BM​) was found to be $26,623.
S2.2 Absorption Column Operating Cost Summary
We require a feed stream of pure water at a rate of 50.58kgmol/hr to obtain the separation that
we need through the absorption column. This equates to 173gallons of H​2​O per month. With a
current price of $0.5735/100 gal of water, we found an annual cost of $23.84 for the water
required to run the absorption column.
Total Capital Cost = $26,623
Total Operating Cost = $23.84/year
Project: Chloromethane synthesis from methane
Unit Name/ID: S3.1 Distillation Column
Unit Type: Distillation Column
Sizing Basis: Volume, m​3

45.
Unit Diagram
Stream Table:​ All values in kg/hr.
Stream ID: S2.s7 S3.s1 S3.s2
Total 733.4417 57.2714 676.1703
Cl2 51.0628 49.8812 1.1815
CH3Cl 315.5189 0.8347 314.6842
CH2Cl2 259.5215 0.0000 259.5215
CHCl3 85.1140 0.0000 85.1140
CCl4 15.6690 0.0000 15.6690
HCl 6.5553 6.5553 0.0000
S3.1 – Distillation Column Size Summary
Column Diameter = 1.193m
Number of Trays = 61
Column Height = 30.5m (assuming the distance between each tray is 0.5m)
olumn V olume π olumn HeightC = * ( )2
Column Diameter 2
* C
Column Volume = 34.1m​3
S3.1 – Distillation Column Cost Summary
We found the distillation column to have a volume of 34.1m​3​
. Looking in Table A.1 of Turton
we determined the purchased equipment cost at ambient operating pressure and using carbon
steel construction with the following equation:

46.
C log (A) [log (A) ]log10
0
p = K1 + K2 10 + K3 10
2
where:
A = 34.1m​3
K​1​ = 3.4974
K​2​ = 0.4485
K​3​ = 0.1074
The bare module cost (C​BM​) for this vessel has to account for its material factors (F​M​) and
pressure factor (F​P​).
(B F F )CBM = C0
p 1 + B2 M P
FP,vessel = 0.0063
+0.00315(P+1)D
2[850−0.6(P+1)]
where
D = 1.193m
B​1​ = 2.25
B​2​ = 1.82
F​M​ = 9.4
P = 3.01 barg
The base module cost (C​BM​) was found to be $509,670
S3.1 Condenser and Reboiler Size Summary
Using the following set of equations as well as information from HYSYS, we were able to size
the condensers as well as reboilers in the distillation column. For the condensers we used liquid
nitrogen and for the reboilers we used saturated steam.
A
Q
= dT
+∑
1
hi k
x
pipe
hi = D
k Nui
u .023Re PrN = 0 0.8 0.3
eR = μ
ρvD
r p( )P = C k
μ
Td =
ln ( ) T −Tsat in
T −Tsat out
(T −T )−(T −T )sat out sat in

47.
A = ( )A
Q
Qneeded
where:
i = each stream in the condenser or reboiler
k​pipe​ = 400 W/mC
D = 0.1407m
x = 0.0021m
Q​needed​, viscosity ( ), thermal conductivity (k), density ( ). T​in​, and T​out​ data was found from ourμ ρ
HYSYS simulation for each exit stream of either the condenser or reboiler.
From this set of equation we were able to calculate the surface area of the heat exchanger for the
reboiler and condenser in the distillation column. The reboiler surface area is 2.4m​2​
, and the
condenser is 3.63m​2​
.
S3.1 Condenser and Reboiler Cost Summary
Knowing the surface area (A) of the condenser and reboiler, we used Figure A.5 in Appendix A
of Turtin to find a C​p​
0​
/A($/m​2​
) where:
C0
p =
C0
p
A( )$
m2
* A
To calculate the bare module cost (C​BM​) of this heat exchanger we used the following equation:
(B F F )CBM = C0
p 1 + B2 M P
where:
B​1​ = 1.74
B​2​ = 1.55
F​M​ = 1
F​P​ = 1
We calculated a bare module cost of $7,166 for the condenser, and $7,915 for the reboiler.
Total Capital Coat = $524,751
Project: Chloromethane synthesis from methane
Unit Name/ID: S3.2 Distillation Column
Unit Type: Distillation Column
Sizing Basis: Volume, m​3
Unit Diagram

48.
Stream Table:​ All values in kg/hr.
Stream ID: S3.s2 S3.s3 S3.s4
Total 676.1703 316.8983 359.2720
Cl2 1.1815 1.1810 0.0005
CH3Cl 314.6842 314.4740 0.2102
CH2Cl2 259.5215 1.2432 258.2783
CHCl3 85.1140 0.0000 85.1140
CCl4 15.6690 0.0000 15.6690
S3.2 – Distillation Column Size Summary
Column Diameter = 1.193m
Number of Trays = 33
Column Height = 16.5m (assuming the distance between each tray is 0.5m)
olumn V olume π olumn HeightC = * ( )2
Column Diameter 2
* C
Column Volume = 18.4m​3
S3.2 – Distillation Column Cost Summary
We found the distillation column to have a volume of 18.4m​3​
. Looking in Table A.1 of Turton
we determined the purchased equipment cost at ambient operating pressure and using carbon
steel construction with the following equation:
C log (A) [log (A) ]log10
0
p = K1 + K2 10 + K3 10
2
where:
A = 18.4m​3
K​1​ = 3.4974
K​2​ = 0.4485
K​3​ = 0.1074

49.
The bare module cost (C​BM​) for this vessel has to account for its material factors (F​M​) and
pressure factor (F​P​).
(B F F )CBM = C0
p 1 + B2 M P
FP,vessel = 0.0063
+0.00315(P+1)D
2[850−0.6(P+1)]
where
D = 1.193m
B​1​ = 2.25
B​2​ = 1.82
F​M​ = 9.4
P = 3.01 barg
The base module cost (C​BM​) was found to be $287,931
S3.2 Condenser and Reboiler Size Summary
Using the following set of equations as well as information from HYSYS, we were able to size
the condensers as well as reboilers in the distillation column. For the condensers we used liquid
nitrogen and for the reboilers we used saturated steam.
A
Q
= dT
+∑
1
hi k
x
pipe
hi = D
k Nui
u .023Re PrN = 0 0.8 0.3
eR = μ
ρvD
r p( )P = C k
μ
Td =
ln ( ) T −Tsat in
T −Tsat out
(T −T )−(T −T )sat out sat in
A = ( )A
Q
Qneeded
where:
i = each stream in the condenser or reboiler
k​pipe​ = 400 W/mC
D = 0.1407m
x = 0.0021m

50.
Q​needed​, viscosity ( ), thermal conductivity (k), density ( ). T​in​, and T​out​ data was found from ourμ ρ
HYSYS simulation for each exit stream of either the condenser or reboiler.
From this set of equation we were able to calculate the surface area of the heat exchanger for the
reboiler and condenser in the distillation column. The reboiler surface area is 14.19m​2​
, and the
condenser is 14.89m​2​
.
S3.2 Condenser and Reboiler Cost Summary
Knowing the surface area (A) of the condenser and reboiler, we used Figure A.5 in Appendix A
of Turtin to find a C​p​
0​
/A($/m​2​
) where:
C0
p =
C0
p
A( )$
m2
* A
To calculate the bare module cost (C​BM​) of this heat exchanger we used the following equation:
(B F F )CBM = C0
p 1 + B2 M P
where:
B​1​ = 1.74
B​2​ = 1.55
F​M​ = 1
F​P​ = 1
We calculated a bare module cost of $17,146 for the condenser, and $16,340 for the reboiler.
Total Capital Coat = $321,417
Project: Chloromethane synthesis from methane
Unit Name/ID: S3.3 Distillation Column
Unit Type: Distillation Column
Sizing Basis: Volume, m​3

51.
Unit Diagram
Stream Table:​ All values in kg/hr.
Stream ID: S3.s4 S3.s5 S3.s6
Total 359.2720 254.7190 104.5531
Cl2 0.0005 0.0005 0.0000
CH3Cl 0.2102 0.2102 0.0000
CH2Cl2 258.2783 254.2844 3.9940
CHCl3 85.1140 0.1996 84.9144
CCl4 15.6690 0.0243 15.6447
S3.3 – Distillation Column Size Summary
Column Diameter = 1.193m
Number of Trays = 73
Column Height = 36.5m (assuming the distance between each tray is 0.5m)
olumn V olume π olumn HeightC = * ( )2
Column Diameter 2
* C
Column Volume = 40.8m​3
S3.3 – Distillation Column Cost Summary
We found the distillation column to have a volume of 40.8m​3​
. Looking in Table A.1 of Turton
we determined the purchased equipment cost at ambient operating pressure and using carbon
steel construction with the following equation:
C log (A) [log (A) ]log10
0
p = K1 + K2 10 + K3 10
2
where:
A = 40.8m​3
K​1​ = 3.4974
K​2​ = 0.4485
K​3​ = 0.1074

52.
The bare module cost (C​BM​) for this vessel has to account for its material factors (F​M​) and
pressure factor (F​P​).
(B F F )CBM = C0
p 1 + B2 M P
FP,vessel = 0.0063
+0.00315(P+1)D
2[850−0.6(P+1)]
where
D = 1.193m
B​1​ = 2.25
B​2​ = 1.82
F​M​ = 9.4
P = 3.01 barg
The base module cost (C​BM​) was found to be $509,676
S3.3 Condenser and Reboiler Size Summary
Using the following set of equations as well as information from HYSYS, we were able to size
the condensers as well as reboilers in the distillation column. For the condensers we used liquid
nitrogen and for the reboilers we used saturated steam.
A
Q
= dT
+∑
1
hi k
x
pipe
hi = D
k Nui
u .023Re PrN = 0 0.8 0.3
eR = μ
ρvD
r p( )P = C k
μ
Td =
ln ( ) T −Tsat in
T −Tsat out
(T −T )−(T −T )sat out sat in
A = ( )A
Q
Qneeded
where:
i = each stream in the condenser or reboiler
k​pipe​ = 400 W/mC
D = 0.1407m
x = 0.0021m

53.
Q​needed​, viscosity ( ), thermal conductivity (k), density ( ). T​in​, and T​out​ data was found from ourμ ρ
HYSYS simulation for each exit stream of either the condenser or reboiler.
From this set of equation we were able to calculate the surface area of the heat exchanger for the
reboiler and condenser in the distillation column. The reboiler surface area is 102.2m​2​
, and the
condenser is 70.16m​2​
.
S3.3 Condenser and Reboiler Cost Summary
Knowing the surface area (A) of the condenser and reboiler, we used Figure A.5 in Appendix A
of Turtin to find a C​p​
0​
/A($/m​2​
) where:
C0
p =
C0
p
A( )$
m2
* A
To calculate the bare module cost (C​BM​) of this heat exchanger we used the following equation:
(B F F )CBM = C0
p 1 + B2 M P
where:
B​1​ = 1.74
B​2​ = 1.55
F​M​ = 1
F​P​ = 1
We calculated a bare module cost of $57,707 for the condenser, and $100,871 for the reboiler.
Total Capital Coat = $668,254