Acetone Retrofit_Senior Design

Executive Summary
The process design is for a retrofit to an existing plant in the Gulf Coast, USA. This process
would upgrade waste streams from an existing plant to produce product-grade acetone for
sale. Currently, these waste streams are burned as fuel and produce high pressure steam.
Environmental regulations have changed, and the company’s boilers no longer meet
environmental regulations, so one of these three options to treat the waste streams needs to
be implemented:
Burn: upgrade the existing boilers so that the waste can continue to be burned
Sell: sell the waste streams to WasteCo.
Build: build an acetone recovery unit and sell the acetone
Our recommendation is to sell the waste streams to WasteCo.
The designed plant produces 189 MMlb of acetone per year and features 7 distillation
columns, 1 isopropanol (IPOH) reactor, and 2 carbon beds. Two waste streams, one
concentrated in isopropanol and one concentrated in acetone, feed into the plant from an
existing operation. The IPOH reactor utilizes a copper on alumina catalyst to convert IPOH
to acetone. The catalyst achieves a conversion of 90% and has a selectivity to acetone of
90%. The plant is estimated to cost 194 MM$ in capital and 40 MM$ in operating costs per
year in order to sell product-grade acetone at the current market price of 40 ¢/lb.
The economic analysis of the base case gave an After Tax Rate of Return (ATROR) of
11.51%, a Net Present Worth (NPW) of 28.3 $MM, and a 5.7 year payback period. The
company currently has multiple projects that it would like to execute over the next three
years which will occupy capital funds and manpower. These projects are all economically
attractive, with ATRORs of 20% or greater and NPWs of $30 MM. The company is also
reluctant to create a new product line, which would be the case if the plant is constructed.
Therefore, the acetone retrofit plant, which costs 194 MM$ to build, requires 60 laborers,
has an ATROR less than 20%, and requires the company to start a new product line is not
economically or strategically attractive. For all of these reasons, building the acetone retrofit
plant is not recommended.
The waste streams are currently burned for use as fuel to produce high pressure steam. This
gives the streams a combined worth of 10 MM$/yr. However, the capital cost of upgrading
the boiler in order to continue burning these streams would be a large capital investment for
the company. It is estimated that this would reduce the worth of these streams to 7.7
MM$/yr when they are used as a fuel source. WasteCo. values the two waste streams at a
combined worth of 17 ¢/lb, which gives a revenue of 32 MM$/year. Selling to WasteCo.
represents the best option economically, and is in line with the company’s objectives as this
option requires no capital investment and does not require the creation of a new product
line.

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Table of Contents
Section One: Background................................................................................................... 6
Background..................................................................................................................... 7
Product Background .................................................................................................... 7
Feed Background ........................................................................................................ 7
Market Survey ............................................................................................................. 8
Section Two: Process Description....................................................................................... 9
Process Description .......................................................................................................10
Overview ....................................................................................................................10
Feed Streams.............................................................................................................10
IPOH Reactor .............................................................................................................10
Block and Process Flow Diagrams .............................................................................12
Separations ................................................................................................................14
Process Specifications ...................................................................................................16
Achieving Hard and Soft Specifications ......................................................................16
Separation Specifications ...........................................................................................19
Reactor Specifications ................................................................................................19
Mass Balance .............................................................................................................20
Energy Balance ..........................................................................................................23
Section Three: Process and Equipment Design.................................................................26
Process/Equipment Design ............................................................................................27
Distillation Column Key Variables ...............................................................................27
General Optimization Technique.................................................................................27
Shell and Tube Reactor Key Variables .......................................................................44
General Optimization Technique.................................................................................45
Reactor and Catalyst Maintenance .............................................................................47
Detailed Equipment Lists................................................................................................47
Inside Battery Limit (IBL).............................................................................................47
Outside Battery Limit (OBL) ........................................................................................48
Section Four: Alternative Cases.........................................................................................49
Alternative Studies .........................................................................................................50
Acetone-Methanol Separation.....................................................................................50
Section Five: Outside Battery Limit ....................................................................................54

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Outside Battery Limit......................................................................................................55
Section Six: Environmental, Safety and Special Design Considerations ............................56
Environmental/Safety Information...................................................................................57
Chemical Information..................................................................................................57
Waste Considerations.................................................................................................59
Safety Precautions......................................................................................................59
Process Hazard Analysis (PHA) .................................................................................60
Discussion of the Process Hazard Analysis: ...............................................................63
Standard Operating Procedure (Startup & Shutdown Procedure):..................................64
Startup Procedure.......................................................................................................64
Shutdown Procedure ..................................................................................................65
Process Control Strategies .........................................................................................65
Special Design Considerations ...................................................................................67
Section Seven: Capital Estimate........................................................................................69
Capital Estimate.............................................................................................................70
Basis...........................................................................................................................70
Summary of Capital Cost Calculations........................................................................71
ICARUS List of Assumptions ......................................................................................71
Section Eight: Operating Costs..........................................................................................74
Overview........................................................................................................................75
Raw Materials ................................................................................................................75
Fixed Costs ....................................................................................................................75
Utilities ...........................................................................................................................76
Section Nine: Economic Evaluation ...................................................................................79
Basis ..............................................................................................................................80
Plant Economics.........................................................................................................80
Fixed Costs.................................................................................................................80
Future Prospects for the Acetone Market....................................................................80
Chemical Commodity Historical and Future Pricing ....................................................81
Basis for Utility Costs..................................................................................................84
Base Case Economic Analysis.......................................................................................85
Sensitivity Analysis.........................................................................................................87
Case 1: Not all of the product can be sold - 30 MMlb/yr surplus..................................87

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Case 2: The Price of Acetone Changes......................................................................88
Case 3: Capital Costs Increase...................................................................................88
Case 4: The Price of Natural Gas Changes ................................................................89
Section Ten: PDRI.............................................................................................................92
PDRI Discussion ............................................................................................................93
Section Eleven: Outstanding Issues...................................................................................96
Technical........................................................................................................................97
Economical ....................................................................................................................97
Environmental/Safety .....................................................................................................97
Section Twelve: Conclusion and Recommendations..........................................................99
Conclusions .................................................................................................................100
Recommendations .......................................................................................................102
Based on the sensitivity analysis ..............................................................................102
Based on the alternative case studies.......................................................................102
Based on the economic analysis:..............................................................................102
Supporting Information for Recommendations:.............................................................103
Sensitivity Analysis ...................................................................................................103
Alternative Cases......................................................................................................103
Section Thirteen: References...........................................................................................105
References...................................................................................................................106
GATE 1 References..................................................................................................106
GATE 2 References..................................................................................................106
GATE 3 References..................................................................................................106
GATE 4 References..................................................................................................107
GATE 5 References..................................................................................................107
Section Fourteen: Appendix.............................................................................................109
Equipment Sizing Calculation Methodologies...............................................................110
Distillation Columns ..................................................................................................110
Reflux Drums............................................................................................................112
Heat Exchangers ......................................................................................................113
Shell and Tube Reactor ............................................................................................115
Hot Oil System..........................................................................................................116
Pumps ......................................................................................................................117

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Compressors ............................................................................................................120
Holdup Tanks ...........................................................................................................121
Carbon Beds.............................................................................................................122
Deciding Where to Place Holding Tanks ......................................................................122
Before the reactor.....................................................................................................122
Before Separator 300 ...............................................................................................123
Before Separator 500 ...............................................................................................123
Before and After Separator 700 ................................................................................123
Material and Type of Holding Tank Consideration.....................................................123
Equipment Specification Sheets...................................................................................124
Distillation Column....................................................................................................124
Heat Exchanger........................................................................................................126
Pumps ......................................................................................................................128
Compressor..............................................................................................................131
Equipment Sizing Calculations by Unit Operation.........................................................133
Tower 100.................................................................................................................133
Tower 200.................................................................................................................137
Tower 300.................................................................................................................141
Tower 400.................................................................................................................146
Tower 500.................................................................................................................151
Tower 600.................................................................................................................156
Tower 700.................................................................................................................160
Reactor.....................................................................................................................165
Hot Oil System .............................................................................................................167
Economic Calculation Methodologies (ICARUS Inputs):...............................................168
Assumptions.............................................................................................................168
Sizing Inputs.............................................................................................................171
Alternate Cases: .......................................................................................................176
ICARUS Individual Equipment Prices .......................................................................177
Price Correlation Curves ..............................................................................................179
Alternative Case Capital and Cash Flow Sheets ..........................................................181
Extractive Distillation.................................................................................................181
Hydrogen-Propylene Separator ................................................................................183

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Sensitivity Analysis:......................................................................................................185
Case 2: Acetone Price Changes...............................................................................187
.....................................................................................................................................187
Case 3: Capital Cost changes...................................................................................188
Case 4: Natural Gas Price Changes .........................................................................190
HYSYS Model ..............................................................................................................191

Background
Our team has undergone the task of determining the best method of treatment for the two
waste streams associated with our current production process. Until now, we have been
burning these waste streams to produce high pressure steam. WasteCo has recently shown
interest in purchasing these streams for their company in order to recover key components.
After taking their intentions into consideration, we feel that it may be possible for our
company to upgrade these waste streams ourselves. More specifically, a copper on alumina
catalyst could be used to convert isopropanol to acetone, which could be combined with the
acetone already present and processed further in order to produce a highly purified acetone
product.
Product Background
Acetone is a commodity chemical with many practical laboratory and household uses. It is a
polar organic compound that is miscible in water and is capable of dissolving many organic
compounds. As a result, it is commonly used as a cleaning agent for glassware in chemical
laboratories. It is also a relatively safe chemical and is therefore much more desirable than
other polar compounds such as methanol or ethanol, which have higher flashpoints and are
therefore more likely to catch fire. Acetone is also commonly used as the main component
in nail polish remover, as it is capable of dissolving the nitrocellulose layer on the surface of
the nail without causing much damage to the nail itself. Acetone’s simple chemical structure
makes it fairly easy to produce in large quantities.
Feed Background
In order to determine whether the acetone
production plant could be profitable, our
company compared the potential profits
of this plant with the amount earned from
burning or selling our waste streams.
Burning acetone yields a high pressure
steam product. Fluctuations in the price
of HPS is assumed to follow the trends of
natural gas. The data in table __ was used
to determine the price of HPS based on
the price of natural gas of
$2.50/MMBTU. This yields an assumed
HPS price of roughly 9 ¢/lb acetone
product. By combusting both the waste
acetone and waste isopropanol streams
from our current production process
using its lower heating value and
assuming a 60% energy yield, our company
Figure 1: Utility prices as a function of natural gas price

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estimates that the current burning of these waste streams earns a profit of roughly $7.7
MM/yr.
WasteCo is currently offering our company 15 ¢/lb for the waste acetone stream and 12 ¢/lb
for the waste isopropanol streams. By considering the mass of each stream that we currently
produce in our process, our company estimates that selling these waste streams to WasteCo
would earn a profit of roughly $34.2 MM/yr. This is a greater profit than our company
currently makes by burning these streams and thus it should be considered as an alternative
practice.
Market Survey
Market prices for truck acetone had shown
decreases throughout 2015. However, recent
increases to almost 40￠/lb have occurred due
to higher raw material costs such as refinery-
grade propylene (RGP). This serves as one of
the two raw materials used in the production
of cumene, the feedstock for phenol/acetone
production.
US spot export acetone prices have also seen a
recent increase in price per lb. The
strengthening of RGP values and increases in
US domestic acetone pricing have been
reflected in export pricing. In addition, US
acetone supply has been tightened due to
upcoming plant turnarounds and lack of recent
imports.
Using the current truck acetone price of 40￠
/lb, current estimates for the design plant
indicate that roughly $75.6 MM/yr of acetone
can be produced (assuming an acetone
capacity of roughly 189 MMlb/yr). Although
acetone prices have decreased significantly
over the past year, the recent stagnation and
slight increases in price change indicate that
the profitability of an acetone production plant
may increase in the near future.
Figure 2: Acetone delivered contract price in 2015-2016 [26]
Figure 3: Acetone Free On Board spot price 2015-2016 [26]

Section Two: Process
Description

Process Description
Overview
The acetone retrofit plant is located in the US Gulf Coast. Two waste streams from an
adjacent production plant production will be fed to the acetone plant in order to produce 189
MMlb of acetone product per year with >99.9% purity. This plant uses a copper on alumina
catalyst to convert isopropanol to acetone. In order to maximize the capacity and purity of
the acetone product, one isopropanol (IPOH) conversion reactor and seven distillation
columns were optimized in this process.
Feed Streams
Two waste streams from an adjacent plant serve as the feed streams for this process. The
compositions of the feed streams are as follows:
Table 1: The compositions of the acetone and IPOH waste streams
The waste acetone is fed to the process at 16,670 lb/hr and the waste isopropanol is fed at
11,706 lb/hr. Both feed streams enter as subcooled liquids at 80 °F and atmospheric pressure
(14.7 psia).
IPOH Reactor
The IPOH reactor converts isopropanol to acetone and hydrogen gas.
Isopropanol is also consumed by several side reactions.

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A shell and tube reactor packed with copper on alumina catalyst was designed for this
process with the reactant stream fed to the tubes. The feed stream to the reactor is pumped to
50 psia and heated to 627 °F using two process streams and high pressure stream utility
(three heat exchangers in series). These conditions allow for an isopropanol conversion of
93.5% and a 90% selectivity with respect to acetone production. Because these reactions are
highly endothermic, a utility stream of hot oil was fed through the shells of the reactor in
order to keep the vessel isothermal. This prevents conversion from falling as more
isopropanol is consumed. A pressure drop of 20 psi occurs throughout the reactor. The
product stream exits as superheated vapor and is immediately compressed to 30 psi and
condensed to liquid using two process streams and refrigerant (three heat exchangers in
series).

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Block and Process Flow Diagrams
Figure 4: The block flow diagram for the acetone retrofit plant

Separations
Acetone Waste Tower (T-100)
The goal of this column is to completely remove the acetic acid from the acetone waste
stream, which limits the number of distillation columns constructed with stainless steel to
this single column. The acetone waste stream is initially pumped to 65.7 psia and heated to
152 °F using one process stream and low-low pressure steam (two heat exchangers in
series). The distillation column contains 11 actual trays. Following a flow meter and control
valve, the feed stream (37.2 psia, 152 °F) enters the column at tray 7. The reflux ratio is set
to 1.001, which results in a condenser duty of -8.645 MMBTU/hr and a reboiler duty of
8.504 MMBTU/hr. The condenser will run at a pressure of 14.7 psia and temperature of
132.7 °F with a pressure drop of 5 psia. The reboiler will run at a pressure of 21.2 psia and
temperature of 228.5 °F. The bottoms product (enriched in acetic acid) is sent to fuel, while
the distillate product is fed into column T-300 after mixing with the distillate product of
column T-200.
Isopropanol Waste Tower (T-200)
The goal of this column is to separate the acetone and methanol from isopropanol and water
present in the isopropanol waste stream. This is done to prevent acetone and methanol from
being fed to the reactor and to collect the acetone present in this feed stream. The
isopropanol waste stream is initially pumped to 65.9 psia and heated to 180 °F using one
process stream. The distillation column contains 35 actual trays. Following a flow meter and
control valve, the feed stream (60.9 psia, 180 °F) enters the column at tray 18. The reflux
ratio is set to 20.59, which results in a condenser duty of -22.66 MMBTU/hr and a reboiler
duty of 22.92 MMBTU/hr. The condenser will run at a pressure of 14.7 psia and a
temperature of 147.6 °F with a pressure drop of 4 psia. The reboiler will run at a pressure of
20.6 psia and a temperature of 193.5 °F. The distillate product is mixed with the distillate
product of tower T-100 and subsequently fed to T-300. The bottoms product is fed to the
IPOH reactor.
Acetone/Methanol Vacuum Tower (T-300)
The goal of this column is to separate the acetone from methanol present in the mixture of
product streams from T-100 and T-200. This increases the purity of the final acetone
product by removing methanol. This separation is very difficult to achieve at atmospheric
pressure due to an azeotrope formed by acetone and methanol, and thus a vacuum
distillation column was used. The feed stream is initially pumped to 49.6 psia and cooled to
43 °F using a process stream and refrigerant (2 heat exchangers in series). The distillation
column contains 55 actual trays. Following a flow meter and control valve, the feed stream
(1.8 psia, 43 °F) enters the column at tray 40. The reflux ratio is set to 7.382, which results
in a condenser duty of -30.71 MMBTU/hr and a reboiler duty of 30.05 MMBTU/hr. The
condenser will run at a pressure of 0.8 psia and a temperature 15 °F. The reboiler will run at
a pressure of 2.2 psia and a temperature of 73.48 °F. The distillate product (enriched in
acetone) is sent to the final column, T-700, after mixing with the distillate of T-500 and the
bottoms product (enriched in methanol) is sent to fuel.

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Gas Products Tower (T-400)
The goal of this column is to separate hydrogen and propylene from the other components in
present in the outlet of the IPOH reactor. Hydrogen and propylene are gasses at STP and
thus can be easily separated from a mixture of liquid components. The feed stream
(superheated vapor) is initially compressed to 40 psia and 725 °F. It is then condensed and
cooled to 25.16 °F using two process streams and refrigerant (three heat exchangers in
series). The distillation column contains 36 actual trays. Following a flow meter and a
control valve, the feed stream (20 psia, 25.16 °F) enters the column at tray 18. The reflux
ratio is set to 10, which results in a condenser duty of -12.49 MMBTU/hr and a reboiler duty
of 12.92 MMBTU/hr. The condenser will run at a pressure of 14.7 psia and a temperature of
-132.5 °F. The reboiler will run at a pressure of 17.6 psia and a temperature of 149.1 °F. The
distillate product (enriched in H2 and propylene) is compressed to 70 psia and sent to fuel
and the bottoms product is sent to column T-500.
Acetone/Isopropanol Vacuum Tower (T-500)
The goal of this column is to separate acetone (the desired product) from all other
components leaving T-400. Although these components do not form an azeotrope, a vacuum
distillation column is necessary to achieve the sufficient separation so that final product
capacity of 189 MMlb/yr is achieved. The feed stream is initially pumped to 83.8 psia. The
distillation column contains 53 actual trays. Following a flow meter and a control valve, the
feed stream (8.1 psia, 111.5 °F) enters the column at tray 27. The reflux ratio is set to 28.68,
which results in a condenser duty of -50.19 MMBTU/hr and a reboiler duty of 49.82
MMBTU/hr. The condenser will run at a pressure of 2 psia and a temperature of 89.7 °F.
The reboiler will run at a pressure of 8.5 psia and a temperature of 150.8 °F. The distillate
product (enriched in acetone) is sent to the final column, T-700, after mixing with the
distillate of column T-300 and the bottoms product is sent to column T-600.
Water Remover (T-600)
The goal of this column is to remove water from all other components leaving T-500. This
is done to limit the amount of water recycling back to the reactor. The feed stream is
initially pumped to 64.2 psia and heated to 176.9 °F using low-low pressure steam. The
which results in a condenser duty of -1.66 MMBTU/hr and a reboiler duty of 1.63
MMBTU/hr. The condenser will run at a pressure of 14.7 psia and a temperature of 198 °F.
The reboiler will run at a pressure of 20.7 psia and a temperature of 230.4 °F. The bottoms
product (enriched in water) is removed from the process as waste and half of the distillate
product (enriched in IPOH) is recycled back to the reactor. The remaining distillate product
is sent to fuel.
Final Acetone Tower (T-700)

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The goal of this column is to remove trace amounts of methanol from the mixture of the
distillate streams from T-300 and T-500 in order to reach final product purity (≥ 99.9 % by
mass). In order to perform the difficult separation of acetone and methanol, this column
operated at a high enough pressure that the separation occurs on the right side of the
azeotrope. As a result, acetone is collected in the bottoms while methanol is collected in the
distillate. The feed stream is initially pumped to 129.12 psia and cooled to 83.4 °F using one
process stream, low-low pressure steam, and cold water (3 heat exchangers in series). The
which results in a condenser duty of -6 MMBTU/hr and a reboiler duty of 5.937
MMBTU/hr. The condenser will run at a pressure of 64 psia and a temperature of 219 °F.
The reboiler will run at a pressure of 70 psia and a temperature of 230.7 °F. The distillate
product (enriched in methanol) is sent to fuel while the bottoms product is our final acetone
product. This stream is then sent to carbon beds for final purification.
Process Specifications
Achieving Hard and Soft Specifications
Table 2: Summary of the hard and soft specifications provided to the design team
Hard Specifications Peter’s Posse’s Design
Product Capacity: 189 MMlb/yr 189.3
Product Acetone Purity: 99.90-99.93 wt% min 99.90
Product Isopropanol: 500 wt ppm max 0
Product Methanol: 1000 wt ppm max 500
Product Acetic Acid: 10 wt ppm max 0
Product Water: 1000 wt ppm max 500
Sellable Hydrogen: 95 mol% min N/A
Fuel Acetone + Methanol: 3 wt% max N/A
Fuel Water: 2 wt% max N/A
Byproduct for Sale: 99.9 wt% N/A
Soft Specifications Peter’s Posse’s Design
Acetone Recovery: 95% 91.2

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Acetic Acid to Avoid Stainless Steel: 50 wt% max 0 in all except one distillation column
4,000 MMBTU/lb Acetone Product Reboiler Duty 5,954
Reactor Feed Specifications Peter’s Posse’s Design
Isopropanol: 85 wt% min 89.7
Acetone: 5 wt% max 1.8
Methanol: 1 wt% max 0.4
Water: 10 wt% max 5.0
The soft specification of recovery was not met because the hard specs were met without it.
Meeting 95% acetone recovery would mean producing 197 MMlb/yr, which means our
design would have to use extra utilities and have slightly larger columns. This would
produce 4.3MM$/yr more in sales. Assuming the utilities and capital costs increase by the
same percentage that the capacity increased, the After Tax Rate of Return (ATROR) of the
process will change to 12.33% from 11.51%. This change is slightly more profitable, but
may only be due to the assumptions. As a higher purity is desired, the energy input
increases non-linearly. This means that the capital cost and energy inputs are probably
much larger than the 4% increase assumed based on the 4% increase in capacity. The
design would probably be less profitable than meeting the 189 MMlb/yr capacity.
The soft specification of 4,000 MMBTU/lb of acetone product reboiler duty was also not
met. Our reboiler duty is 48% higher than the soft spec because of the high reboiler duties
in the two vacuum towers. Combined, the two vacuum towers contribute 3702 BTU/lb in
reboiler duty. This is because the separation of acetone and methanol is an azeotrope that
cannot be separated under atmospheric conditions. The high reboiler duty is compensated
for by crossing process streams later in the process to save energy.
Table 3: Summary of the conditons in each of the plant’s seven distillation columns
Column Reflux
Ratio
Trays Feed
Tray
Feed Condenser Reboiler
Temp Press Temp Press Duty
(MMBT
U/hr)
Temp Press Duty
(MMBT
U/hr)
100 1.001 11 7 152 37.2 132.7 14.7 -8.645 228.5 21.2 8.5
200 20.59 35 18 180 60.9 147.6 14.7 -8.645 193.5 20.6 22.92
300 7.382 55 40 43 24.7 15 0.8 -30.71 73.48 2.2 30.05

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400 10 36 18 25.16 20 -132.5 14.7 -12.49 149.1 17.6 12.92
500 28.68 53 27 111.5 8.1 89.7 2 -50.1 150.8 8.5 49.82
600 1.002 6 4 176.9 14.9 198 14.7 -1.66 230.4 20.7 1.63
700 107.2 43 30 83.4 25.2 219 64 -6 230.7 70 5.937

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Separation Specifications
Table 4: Summary of the key light and heavy components that were separated in each of the seven distillation
columns
Column Key Light Key Heavy
1 Water Acetic Acid
2 Methanol Isopropanol
3 Acetone Methanol
4 Propylene Acetone
5 Acetone Isopropanol
6 Isopropanol Water
7 Acetone Methanol
Reactor Specifications
Table 5: The IPOH reactor conditions
Inlet Pressure (psia) 50
Maximum Pressure Drop (psi) 20
Temperature (°F) 627.4
Feed Flow Rate (lb/hr) 9781
Weight Catalyst (lb) 39,200
Conversion of IPOH (%) 90
Acetone Selectivity (%) 90
Mesityl Oxide Selectivity (%) 8
Propylene Selectivity (%) 2

Section Three: Process and
Equipment Design

Process/Equipment Design
Distillation Column Key Variables
There are various factors that affect the design of a distillation column. Pressure is the most
important parameter. At high pressures, the relative volatility of most two-component
systems decreases and the separation becomes more difficult. Higher pressures require
either more trays (higher capital cost) or a higher reflux (greater utility cost) in order to
achieve the separation. The capital cost of a column is also intrinsically higher at higher
pressures, as a thicker material of construction is needed to be able to withstand the pressure
exerted by the vapor on the walls of the column. Separation becomes much easier at
pressures below atmospheric, but these systems require expensive vacuum equipment and
dramatically increase utility costs. Therefore, most of the columns in the acetone retrofit
system were designed to operate at atmospheric pressure in the condenser. Two of the seven
columns, however, are operating under vacuum. Column 500 operates under vacuum
because an extremely pure top stream of acetone was required to be sent to the final
separator in order to meet specifications. Column 300 also operates under vacuum because
operating at such a low pressure allowed the design group to get around the acetone-
methanol azeotrope, which is key for this process. For each column, a pressure drop of 0.1
psi was assumed for each tray, a pressure drop of 4 psi was assigned to each condenser, and
a negligible pressure drop was assumed for the reboilers.
General Optimization Technique
To begin, the pressure in the condenser of each column was set to atmospheric pressure, as
this is the lowest pressure that the column could operate at without vacuum. For each
column, an arbitrary number of trays was put into HYSYS to allow for the desired
separation to occur. Then, the number of trays was reduced until the reflux ratio began to
greatly increase. From this analysis, the number of trays was tentatively set for each column.
Based on the tentative number of trays for each column and a 0.1 psi pressure drop per tray,
the pressure in the reboiler of each column was also tentatively set.
Then, based on the pressure and temperature profile of each column, the feeds to each
column were modified using heat exchangers, pumps, and valves so that the pressure and
temperature of the feed matched the pressure and temperature at the middle of each column.
This was a major design decision because having a feed composition that matches closely
with the composition at the feed tray in the column allows for the best separation. If the feed
pressure and temperature vary greatly from the pressure and temperature of the liquid and
vapor at the feed tray in the column, then mixing will occur in a portion of the column
which will reduce the column efficiency. The feed conditions were determined before a
knee of the curve analysis was performed because the design group assumed that the
number of trays would not vary greatly from the tentative values.
Then, with the feeds at the proper pressure and temperature, a knee of the curve analysis
was performed for each column. The number of trays versus reflux ratio was plotted for
each column and the number of trays found at the knee of the curve was selected. This knee

28
of the curve analysis leads to a minimization of both capital and utility cost. A greater
number of trays in a column gives more stages for contact between the rising vapor and
downward-flowing liquid, which allows for better separation. However, increasing the
number of stages requires a taller column and a greater capital cost. Increasing the reflux
ratio results in greater flows in the column, which gives a higher mass transfer coefficient
and better separation at each stage. However, increasing the reflux ratio increases both the
condenser and reboiler duties, as more vapor must be condensed and more liquid must be
vaporized. Increasing the reflux ratio also increases capital cost, as a wider column is
required to handle the increased vapor flow rates.
With the number of trays selected, a second optimization was performed. A plot of reflux
ratio versus feed tray location was made for each column and the feed tray that gave the
minimum reflux ratio was determined. Since the condenser and reboiler duties are
proportional to the reflux ratio, the feed tray that minimizes the reflux ratio also minimizes
these duties and leads to lower utility costs.
The optimal number of trays for each column found using the knee of the curve analysis are
based on HYSYS data and are therefore the theoretical number of trays. When the columns
were sized to determine their height and diameter, the theoretical number of trays for each
column was an input used to find the actual number of trays.
The optimal feed tray location found in the following optimizations is also a theoretical
value, and was later scaled up when the column heights and diameters were determined.
Distillation Column 100 (Acetone Waste Tower)
Purpose
The purpose of this column is to remove the acetic acid that is present in the acetone waste
stream so that the remainder of the columns in the process can be made of carbon steel
instead of stainless steel, which greatly reduces the capital cost for the plant.

29
XY Analysis
The XY diagram for isopropanol and acetic acid at atmospheric pressure is shown in Figure
5. The separation is relatively easy as can be seen from the separation of the equilibrium line
from the y=x line.
Feed Condition
The pressure at the condenser was set to atmospheric pressure. The feed temperature and
pressure were specified to match the pressure and temperature at the middle of the column.
The pressure of the feed is 16.70 psia and the temperature is 140.7°F, which gives a vapor
fraction of 0.0217.
Column Sizing and Feed Tray Determination
With the feed conditions specified, a knee of the curve analysis was performed by making a
plot of number of trays versus reflux ratio, which is shown as Figure 6.
Figure 5: XY diagram for IPOH and acetic acid at atmospheric pressure

30
Using the knee of the curve method, the optimum number of theoretical trays was found to
be 5. With the number of theoretical trays determined, the next thing to be determined was
the feed tray location. Figure 7 shows a plot of reflux ratio versus feed tray location, which
was used to determine the optimum feed tray location.
From Figure 7, the optimal feed tray was found to be tray 3, as this minimizes the reflux
ratio.
Figure 6: Knee of the curve analysis to find optimal number of trays for Column 100
Figure 7: Determining the feed tray location for Column 100

31
Distillation Column 200 (Isopropanol Waste Tower)
Purpose
The purpose of the Isopropanol Waste Tower is to send almost all of the isopropanol in the
waste stream down to the reactor system so that it can react to form acetone. Almost all of
the methanol and acetone fed to the tower leaves in the top and is sent to Column 300 where
the methanol is separated from the acetone.
XY Analysis
The XY diagram for isopropanol and methanol at atmospheric pressure is shown in Figure
8. The separation is fairly difficult as can be seen from the separation of the equilibrium line
from the y=x line. This explains why this tower has a relatively high theoretical number of
stages.
Feed Condition
The feed is a subcooled liquid with a pressure of 19.70 psia and a temperature of 150.2°F.
Figure 8: XY diagram for IPOH and methanol at atmospheric pressure

32
Figure 9: Knee of the curve analysis to find the optimal number of trays for Column 200

33
ratio.
Distillation Column 300 (Acetone/Methanol Vacuum Tower)
Purpose
Column 300 is operated at vacuum in order to separate acetone from methanol. A 99.50
wt.% acetone stream leaves from the top of the column and is sent to mixing point C to mix
with acetone produced from the reactor. A 87.22 wt.% methanol stream leaves as the
bottoms and is used for fuel.
XY Analysis
The XY diagram for acetone and methanol at atmospheric pressure is shown in Figure 11.
The separation is extremely difficult as can be seen from the separation of the equilibrium
line from the y=x line. There is also an azeotrope that occurs at approximately 84 wt.%
acetone which makes a column that produces a >99 wt.% acetone stream at atmospheric
pressure impossible.
Figure 11: XY diagram for acetone and methanol at atmospheric pressure

34
The XY diagram for acetone and methanol at 1.5 psia is shown in Figure 12. At this
extremely low pressure, the separation becomes much easier and the azeotrope is no longer
present. However, the tower still has a very high number of theoretical stages because
getting the desired tops acetone purity of 99.5 wt.% adds on a greater number of stages.
Feed Condition
The pressure at the condenser was set to 0.20 psia. The feed temperature and pressure were
specified to match the pressure and temperature at the middle of the column. The feed is at a
pressure of 2 psia and a temperature of 48.86 °F, with a vapor fraction of 0.09.
Figure 12: XY diagram for acetone and methanol at 1.5 psia

35
ratio.
Figure 13: Knee of the curve analysis to determine the optimal number of trays for Column 300

36
Distillation Column 400 (Gas Products Tower)
Purpose
The purpose of the Gas Products Tower is to remove the lightest components formed during
the reaction, hydrogen and propylene. These gases are removed from the tops of this tower,
whose condenser runs at total reflux. All other species coming from the reactor come out the
bottom of this tower before being separated in the following columns.
XY Analysis
The XY diagram for propylene and acetone at atmospheric pressure is shown in Figure 15.
The separation is extremely easy as can be seen from the large distance between the
equilibrium line and the y=x line.
Feed Condition
The feed is at a pressure of 20.00 psia and a temperature of 25.14 °F, with a vapor fraction
of 0.4345.
Figure 15: XY Diagram for propylene and acetone at atmospheric pressure

37
The column would only converge in HYSYS with 10 theoretical trays, the feed at tray 5,
and a reflux ratio of 10.00.
Distillation Column 500 (Acetone/Isopropanol Vacuum Tower)
Purpose
Column 500 is operated at vacuum in order to separate acetone from isopropanol. A 99.82
wt.% acetone stream leaves from the top of the column and is sent to mixing point C to mix
with acetone coming from Column 300. Almost all of the unreacted isopropanol was sent
out of the bottoms of this column. It was desired to send the isopropanol to the bottoms
stream so that as much unreacted isopropanol as possible could be recycled back to the
reactor. This column is operated at vacuum because of the high purity specification of the
tops stream.
XY Analysis
The XY diagram for acetone and isopropanol at atmospheric pressure is shown in Figure
16. The separation is fairly easy as can be seen from the distance between the equilibrium
line and the y=x line. However, since a nearly pure acetone distillate stream was required, a
high number of theoretical trays were needed for this column.
Figure 16: XY Diagram for acetone and isopropanol at atmospheric pressure

38
Feed Condition
pressure of 8.10 psia and a temperature of 111.5 °F, with a vapor fraction of 0.0852.
plot of number of trays versus the reflux ratio of Column 700. This was done because
simply optimizing Column 500 on its own led to extremely high and unrealistic reflux ratios
in Column 700, which could not be reduced. Therefore, Column 500 was optimized with
respect to Column 700 since the distillate of Column 500 is fed to Column 700 and plays a
major role in that column’s design. The knee of the curve analysis is shown as Figure 17.
the feed tray location. Figure 18 shows a plot of reflux ratio of Column 500 versus feed tray
location, which was used to determine the optimum feed tray location.
Figure 17: Knee of the curve analysis to determine the optimal number of trays for Column 500

39
From Figure 18, the optimal feed tray was tray 16. Above this tray, the feed would have
been below stage pressure. Therefore, a plot of reflux ratio versus feed tray location was
only performed up to tray 16.
Distillation Column 600 (Water Remover)
Purpose
The purpose of Column 600 is to remove a large amount of the water in the system so that
the recycle back to the reactor meets the specification for water fed to the reactor. The
distillate contains a large amount of unreacted isopropanol that is fed back to the reactor.

40
XY Analysis
The XY diagram for isopropanol and water at atmospheric pressure is shown in Figure 19.
The separation is fairly easy as can be seen from the large distance between the equilibrium
line and the y=x line. Since an azeotrope exists between isopropanol and water at
atmospheric pressure, it was not possible to remove all of the water in the feed. This was
acceptable, however, because the reactor feed specifications were able to be met without all
of the water being removed.
Feed Condition
Figure 19: XY diagram for isopropanol and water at atmospheric pressure

41
From this plot, it can be seen that the reflux ratio does not change with number of trays.
Therefore, it was decided to use the smallest possible theoretical number of trays, 3.
The feed was decided to enter at the middle of the column at tray 2.
Distillation Column 700 (Final Acetone Tower)
Purpose
The purpose of this column is to remove trace amounts of methanol in order to meet the
acetone purity specification of 99.90%. This tower operates at high pressure to move to the
right of the acetone-methanol azeotrope, which causes acetone to be the bottoms product
and methanol to be the distillate.
XY Analysis
The XY diagram for methanol and acetone at the condenser pressure of 64.00 psia is shown
in Figure 21. This plot shows that almost all of the methanol is able to be removed from the
top of the column.
Figure 20: The reflux ratio does not change with number of trays for Column 600

42
Feed Condition
Figure 21: XY diagram for methanol and acetone at 64.00 psia

43
From Figure 23, the optimal feed tray was tray 26.
Figure 22: Knee of the curve analysis to find the optimum number of trays for Column 700

44
Shell and Tube Reactor Key Variables
The reactor design was dependent on information provided by our research team. The inlet
pressure and total pressure drop through the reactor were specified as 50 and 20 psi,
respectively since equilibrium is favored by low pressure. Selectivity and conversion are
temperature dependent, which makes it important to keep the reactor isothermal so that a
consistent product purity is maintained. The reactor also needed to be designed large
enough to hold the catalyst given its dimensions and its weight hourly space velocity
(WHSV).
The desired reaction is the dehydrogenation of IPOH to acetone and hydrogen, shown
below. Two major side reactions were accounted for that IPOH could participate in: IPOH
can participate in an aldol condensation reaction to form mesityl oxide, water, and
hydrogen, and IPOH can undergo a dehydration reaction to form propylene and water, also
shown below.
The desired reaction is an equilibrium reaction, so the reactor was designed at a high
temperature and low pressure to drive the process the reaction in the forward direction. The
plug-flow characteristics of the reactor also help to drive the reaction to equilibrium by
avoiding uniform mixing of the reaction. Removing acetone will help to shift the reaction
towards completion based on Le Châtelier's principle. The plug-flow properties of the shell
and tube reactor is favorable for this because the concentration of acetone starts very low,

45
and ends at the outlet concentration. If a CSTR type reactor was chosen, the reactor would
always be run at the outlet acetone concentration, decreasing acetone production.
The production of side products was minimized by choosing the appropriate reaction
temperature. The selectivity for mesityl oxide and propylene increased with temperature, as
did the conversion of IPOH. The tradeoff between selectivity and conversion was
considered and optimized.
Reactor Choice
A shell and tube reactor was chosen based on the volume needed for the heterogeneous
catalyst, and the surface area needed for heat transfer to keep the reactor near isothermal
operation. A direct fired heater reactor will not be used because a fixed bed or shell and
tube reactor in combination with available utilities can accommodate the temperatures that
are needed; and it is a more expensive alternative.
For an assumed flow rate of 9,800 lb/hr into the reactor, a heat input of 3.55 MMBTU/hr
yielded a process temperature change of 52.9o
F. Because the reactor needs to be run at
650o
F, hot oil at 750o
F must be used as the heat transfer fluid since the temperature
approach is 100-200o
F when heating above 600o
F. Dowtherm oil was chosen as the heat
transfer fluid, and it has an overall heat transfer coefficient of approximately 15 BTU (hr ft2
o
F)-1
[24]. The desired outlet temperature of the hot oil needed to be about 730o
F to achieve
a ∆TLM less than 90o
F (730o
F gives ∆TLM=81o
F), and to stay above the 100o
F temperature
approach. The minimum area required for heat transfer was determined to be 2,895 ft2
. The
surface area of the packed bed reactor was assumed to be the same as the heat transfer area.
For a fixed bed reactor, a length over diameter ratio of 3 was used to find a diameter of 17ft
and a resulting reactor volume equal to 11,575ft3
. Based on the catalyst’s WHSV, the
volume needed to accommodate the catalyst with a void fraction of 0.3 is 1,254ft3
. Because
of the factor of 10 difference in reactor volume needed for the catalyst versus the volume
needed for heat transfer using a cylindrical packed bed, a reactor with a higher area of heat
transfer to volume ratio will be needed, such as a shell and tube heat exchanger design.
General Optimization Technique
The research group that developed our catalyst specified that the reactor feed needed to be
50 psi, with a maximum pressure drop of 20 psi. The tube diameter is set to 1” to hold the
catalyst with a maximum linear length (L) of 40 ft, and the volume for the catalyst was set
by its WHSV of 0.25 (lb feed/hr)/(lb catalyst). The number of tubes (N) was calculated
using the required catalyst volume and individual tube volume at a specific length. The
maximum N per reactor was specified to be 10,000. The Ergun Equation (Appendix,
Equipment sizing calculation methodologies) was used to determine the pressure drop
through the tubes. The reactor was sized by iterating the linear length to get a pressure drop
below 20 psi, and fewer than 10,000 tubes. The area for heat transfer was not a constraint
because the required area is 2,895 ft2
when using 750o
F hot oil, and the surface area of
10,000 tubes is on the order of 60,000 ft2
. Because equilibrium is favored by low pressure,
we chose to design to the maximum pressure drop of 20 psi, which also gave the minimum

46
number of tubes, helping to make catalyst replacement easier. The design is: N=6967, L=33
ft, and ∆P=19.9 psi.
Reactor (IPOH)
Purpose
The reactor converts a feed stream of 0.80 mass fraction IPOH and 0.0086 mass fraction
acetone into a stream of 0.090 mass fraction IPOH and 0.71 mass fraction acetone. It needs
to provide enough volume to hold the catalyst, and enough surface area for heat transfer to
maintain a nearly isothermal reactor.
Conversion and Selectivity Analysis
The optimal operating temperature for the reactor was determined by finding the knee of the
curve for conversion versus selectivity. This fell between two data points, which
corresponded to 600 and 700 o
F. An operating temperature of 650 o
F was chosen. The
values of conversion at 600 and 700 o
F were averaged to find the conversion of 0.90 at 650
o
F. The conversion versus selectivity for the side products was also plotted to make sure
there were no significant differences in mesityl oxide or polypropylene selectivity between
600 and 700 o
F. At an IPOH conversion of 0.935, the selectivity of the side products fell in
a near-vertical region, meaning those variables are not sensitive to temperature changes
between 600 and 700 o
F, and do not need to be further considered. A conversion of 90%
was used in calculations and the HYSYS model to account for the temperature variations
within the reactor because it is not perfectly isothermal.
Figure 24: Knee of the curve optimization of conversion and selectivity

47
Hot Oil Heating Loop
Purpose
A utility needed to be provided to keep the reactor running isothermally. Based on the heat
of reaction and the moles of IPOH reacted, it was determined that 3.55 MMBTU/hr of heat
needs to be provided to the reactor to keep it near isothermal. The most economical utility
that could supply heat to reach a reaction temperature of 650 o
F was hot oil heated to 750o
F.
Optimization of Utility Stream Used to Heat the Hot Oil
There were seven waste streams in our process that could be used in the direct fire heater to
heat the hot oil. The utility in each stream was determined by multiplying the lower heating
value by the flow rate of the stream. The stream has to provide 3.55 MMBTU/hr, and no
single stream provided enough heat without providing ≥100% more than necessary. By
combining the bottoms of T-100 and distillate of T-700, a total of 5.646 MMBTU/hr can be
supplied to the direct fire heater, which can transfer 3.67 MMBTU/hr based on a 65%
thermal efficiency [18]. This provides enough heat to the reactor, with a safety factor of
1.03, and allows us to use a waste stream directly in the process.
Reactor and Catalyst Maintenance
Regeneration Process
Purpose
The copper on alumina catalyst experiences losses in activity (a function of the rate constant
and conversion) over time. This is likely due to coke forming on the surface as the
hydrocarbons pass over it at high temperature. Coke formation is known to happen during
dehydrogenation reactions, and has specifically been seen on a copper on alumina catalyst
[11, 14].
Regeneration Process
The catalyst must be regenerated every 6 months, and the entire regeneration cycle takes 7
days. Because Eurecat is the company supplying our catalyst and has a location in the US
Gulf Coast, we will be using their expert catalyst regeneration services rather than designing
and operating the process in-house.
Detailed Equipment Lists
Inside Battery Limit (IBL)
The IBL contains all of the essential equipment to meet our plant capacity and hard
specifications. A summary of the number of each piece of process equipment for the base

48
case is given below. For details on the sizing of each, see the Appendix, Alternative Case
Capital and Cash Flow Sheets.
Table 6: Summary of all process equipment required for the acetone retrofit plant
Outside Battery Limit (OBL)
The OBL contains all of the auxiliary support equipment for our process. This is existing
infrastructure from the existing process. This includes the utility systems, which
encompasses refrigerant, cooling water, low low pressure steam, low pressure steam,
medium pressure steam, and high pressure steam. The steam system contains boilers that are
currently fed by the streams that would become feed streams to this acetone retrofit process,
pressurizing equipment, liquid and gas fuel storage tanks, and the steam distribution system.
There is also a substation to provide electricity for the process needs such as pumps and
compressors. The OBL will house the product holdup tank which can store the acetone
product for 14 days, and emergency flares for system leaks or when rupture disks break.

Section Four: Alternative
Cases

Alternative Studies
Acetone-Methanol Separation
Currently, the designed base case acetone recovery plant has 7 distillation columns, 1
reactor and 2 carbon beds. The challenge to minimize cost came with the acetone-methanol
separation. Three main types of distillation were designed and tested to separate acetone and
methanol. These systems were vacuum distillation (base case), extractive distillation, and
pressure-swing distillation.
The base case features two vacuum columns to separate acetone and methanol. These were
Columns 300 and 500. Column 300 features a length of 198 feet and a diameter of 15 feet.
Column 500 features a height of 139 feet and a diameter of 13 feet. These columns are at the
maximum possible diameter that allow the columns to be prefabricated and shipped to the
plant location. The low pressures in these columns cause the azeotrope to disappear and
allows for nearly pure acetone to be obtained in the distillate of each column.
Extractive distillation is used for mixtures with low relative volatility and those that form an
azeotrope. Extractive distillation uses an entrainer as a separation solvent. The entrainer is
miscible in the mixture and has a higher boiling point. The entrainer is added to enhance the
separation between the acetone and methanol while avoiding the formation of an azeotrope.
In this case water was used as the entrainer to separate acetone and methanol. These
columns were modeled in HYSYS and then sized. The first column has a height of 109 ft
and a diameter of 5 ft, while the second column has a height of 149 ft and a diameter of 15.6
ft. The cost compared to the vacuum case can be seen in Table 7. A major disadvantage
with extractive distillation is the large duty of the feed pump due to the requirement of
feeding 20,000 lb/hr of entrainer to the columns. The extractive distillation system is shown
in the image below.
Pressure Swing Distillation is another method that breaks the acetone methanol azeotrope to
produce a pure stream of acetone. The HYSYS schematic shown below is the pressure
swing system. The theory behind this separation technique is to operate the first tower at
low pressure and then the second column at high pressure creating the pressure swing. This
Figure 25: HYSYS simulation snip of the extractive distillation system

51
breaks the azeotrope by removing the acetone as the bottoms product of the first column.
The distillate goes through the high pressure column to produce a methanol stream out the
bottoms of the second column. The distillate of the second column gets recycled back and
fed to the first column to conserve as much acetone as possible. This separation technique
was able to meet the desired production of acetone, but it came at a very high utility and
capital cost derived from the extreme recycle flow rate and column diameters.
When attempting to size the two columns for the pressure swing distillation system, neither
of the flow rates for the liquid and vapor allowed for the Glitsch Method plot to be used. For
the first column, which operates at high pressure (approximately 50 psi), the Glitsch Method
plot was able to be extrapolated to account for the high flows in the column. This gave an
estimated diameter of 25.5 feet and a height of 57.5 ft. For the second column, which
operates at vacuum, the flows were so high that an extrapolation of the Glitsch method plot
could not be obtained. It is estimated that the diameter of the column would have to have
been at least 50 feet. The minimum column diameter for the column to be prefabricated and
shipped to the Gulf Coast location is 15 feet. Therefore, each column would have to be
fabricated on site, and the capital cost of the large diameter columns plus the construction
cost would be astronomical. The pumps required to move the extremely high flows in
pressure swing distillation system (due to the large recycle stream) would also require a very
large amount of energy. Thus the conclusion was drawn that pressure swing distillation was
an unfeasible solution for acetone methanol separation.
Economic Analysis of the Acetone-Methanol Separation Techniques
As previously explained, the pressure-swing system featured such large flows that the
capital cost would have been exorbitantly high and thus that system was not analyzed
further.
Table 7 shows the capital cost and utility cost associated with the base case and extractive
distillation alternative case. The capital costs are fairly similar, but the total operating cost
per year is approximately 2.5 times higher for extractive distillation system.
Figure 26: HYSYS simulation snip of the pressure swing distillation system

52
Table 7: Total capital cost and operating cost per year for the base case and two alternative cases
Separation Method Capital Cost ($) Total Operating Cost ($/yr)
Vacuum Distillation (Base
Case)
194,136,000 40,000,000
Extractive Distillation 193,867,000 113,000,000
Pressure Swing Distillation N/A N/A
Further economic analysis was performed on the base case and extractive distillation
systems. The extractive distillation system featured a raw material cost associated with
adding 20,000 lb/hr of water to the system as an entrainer. The cost of this water stream was
determined to be 17.95 ¢/lb acetone. This stream alone made this process economically
unfeasible. On top of that, there are high refrigeration costs (22.37 ¢/lb acetone) associated
with the condenser of the second column in the extractive distillation system, which
contributes to the operating cost of about $40 million per year. For the extractive distillation
system to reach the ATROR hurdle rate of 20%, the price of acetone would have to raise to
93.5 ¢/lb, which is more than double its current price.
The base case requires a slightly higher capital cost due to the presence of the vacuum
system and the large size of the vacuum column. The operating cost is much lower,
however, due to the fact that the condenser in Column 300 uses a lower cost refrigerant than
the second column in the extractive distillation system and because there is no required
entrainer stream. The acetone price required for the base case plant to reach the ATROR
hurdle rate of 20% is 54.4 ¢/lb, which is approximately 14 cents higher than its current
price. Based on this number, it can be concluded that the base case is a more economical
option than both the extractive distillation and pressure-swing distillation alternative cases.
Additional Separations and Containments:
An additional separation that was considered was the separation of hydrogen from
propylene. This would produce two alternative product streams for additional revenue. The
hydrogen-propylene product in the base case is used to fuel the fire heater, saving the cost of
natural gas that would otherwise be needed to fuel the fire heater. The addition of this
separator also adds a heat exchanger, a pump and a compressor. This is depicted in Figure
27.

53
The costs associated with the hydrogen-propylene separation system are summarized in
Table 8.
Table 8: Summary of the additional capital and utility costs associated with the addition of the hydrogen-
propylene separation system
Separation Method Capital Cost Utility Cost Total
Separator (PPE-H2) $1,909,000 $315,133
This process produces 246.3 lb/hr of 99.74 wt.% hydrogen which can be sold for 81 ¢/lb, as
well as 441.5 lb/hr of 100 wt.% propylene which can be sold for 41 ¢/lb. These two
additional sources of revenue increase the ATROR of the project from 11.51% to 13.55%,
based on an acetone price of 40 ¢/lb. The entire case flow sheet for the base case plus this
hydrogen-propylene system can be found in the Appendix.
In conclusion, the best technique for separating acetone and methanol is vacuum distillation.
This technique employs an expensive vacuum and refrigeration system to achieve the
separation, but avoids the extremely high utility costs associated with the extremely large
flows in the pressure-swing and extractive distillation systems. While the base case appears
to be the most effective system, the addition of a hydrogen-propylene separation to the base
case plant gives a better return on investment. Although the company does not want to get
into new product lines, the production of hydrogen and propylene as products increases the
ATROR by approximately 2%, making it a viable option to consider in addition to the base
case.
Figure 27: HYSYS simulation snip of the hydrogen-propylene separation system

Section Five: Outside
Battery Limit

Outside Battery Limit
The OBL is located 1 mile from the plant. It contains the equipment needed to produce all
of the utilities including electricity, product storage, and flares for product leaks. It does not
include refrigeration or hot oil systems, which are included in the IBL.
Quotes from external contractors for various elements of the capital cost of constructing the
OBL were provided from previous years. Table 9 shows these bids. The costs given in the
quotes were scaled to present-day costs using provided correlations. Table 9 also shows the
summary of the present-day OBL capital costs.
Table 9: Summary of previous OBL bids and the present-day OBL capital costs

Section Six: Environmental,
Safety and Special Design
Considerations

Environmental/Safety Information
Chemical Information
Hydrogen
Hydrogen is a gas at room temperature and is typically the product from the reactions in the
process. It is a side product from the oxidation reaction from isopropanol to acetone and
from the reaction that converts isopropanol to mesityl oxide. It is flammable (even at low
concentrations) and usually travels with propylene throughout the entire process due to the
similar boiling points.
Propylene
Propylene is a product that is produced from a dehydration reaction of isopropanol.
Propylene is also a gas at room temperature that has a high flammability NFPA category of
4. It is highly flammable and oxidants were avoided to explosive behavior. Similarly,
contact of cold liquid propylene with water was also avoided due to the large temperature
difference.
Methanol
Methanol is found in both of the initial waste streams. Methanol is completely soluble in
water and is a liquid at room temperature. It is a flammable liquid and it is toxic orally.
Since it was soluble in water, there was an azeotrope between the two compounds in the
separation.
Mesityl Oxide
Mesityl oxide is the main product of a side reaction of multiple isopropanol forming mesityl
oxide, hydrogen and water. Mesityl oxide is a liquid at room temperature with a low
solubility in water. It is also a very flammable compound that is also toxic. It is not very
reactive but is a side product that reduces the purity of the product stream and ideally goes
to fuel along with multiple other components.
Acetic Acid
Acetic acid is only present initially in the acetone waste stream. Acetic acid is a liquid at
room temperature while being completely soluble in water. It is a somewhat flammable
liquid with a NFPA category rating of 2. It also is toxic orally, and dermally. Since, the
process was designed in a way to eliminate the acetic acid as quickly as possible from the
system due to its corrosive nature, incompatible materials like oxidizing agents, hydroxides
and some metals were not a primary concern in the end products of the design. Therefore
part of the process had a stainless steel component to avoid corrosion.

58
Formaldehyde
Formaldehyde is found as a trace product from the main and side reactions. It is a liquid at
room temperature. Formaldehyde is somewhat flammable with a NFPA category of 2 and is
very toxic if ingested and hazardous through skin contact, eye contact or inhalation. It is
reactive with many components like anhydrides, carbonyl compounds, oxides and
peroxides. Polymerization can be inhibited by adding methanol or stabilizers such as methyl
cellulose.
Isopropanol
Isopropanol is a liquid at room temperature and is the reactant that produces acetone. It
comes in as large quantities through incoming waste streams. It is a liquid at room
temperature and has a very high flammability with a NFPA category of 3. It is completely
soluble in water, reacts violently with hydrogen, oxidants, and is incompatible with many
acids, alkali metals, Isopropanol reacts with metallic aluminum at high temperatures and
attacks some plastics, rubber, and coatings. Isopropanol can also be peroxidized.
It undergoes an oxidation in the main reaction to produce acetone and in the side reaction to
produce the mesityl oxide. While, in the last side reaction it undergoes a dehydrogenation
reaction to produce propylene. Considering the reactivity of all the components in the
streams, many holding tanks were constructed out of nickel.
Acetone
Acetone is the end desired product of the system. It is a liquid at room temperature and is
completely soluble in water. It is also very flammable with a NFPA category of 3.
Additionally, it is toxic orally and dermally. It undergoes explosive reactions with
chloroform and base and reacts violently with some acids.
Table 10: Summary of physical and chemical properties for each of the chemicals present in the plant
Chemical Molecular
Weight
(g/mol)
Boiling
Point (C)
Freezing
Point
(C)
Flash
Point (C)
Toxicity Flammability
(UFL/LFL) by
volume
Reactivity
Hydrogen 2.016 -252.8 -259.2 -149.99 Simple
asphyxiant
4%/74.2% Highly
flammable.
Strong reducing
agent
Propylene 42.08 -47.7 -94 -107.990 Nontoxic 2.4%/11.0% Highly
flammable.
Methanol 32.04 64.7 -98.0 9.7 LD50 Oral %/36% Acid chlorides,
acid anhydrides,
oxidizing agents,
alkali metals

59
Mesityl Oxide 98.15 130 -41.5 31 Acute
toxicity
1.4%/7.2% None
Acetic Acid 60.05 117.5 16.2 40 LD50 Oral,
LC50
Inhalation,
LC50
Dermal
4%/19.9% Oxidizing
agents,
hydroxides,
Water 18.016 100 0 N/A Nontoxic Nonflammab
le
Water reactive
substances
Formaldehyde 30.031 98 -15 50 Ingestion,
skin
contact, eye
contact
hazard
6%/ 36.5% Incompatible
with carbonyl
compounds,
oxides
Isopropanol 60.10 82 -89.5 12.0 Inhalation/
Oral
2%/12.7% Reacts violently
with hydrogen
Acetone 58.08 132.8 -94 -17 Oral
(LD50),
Inhalation
(LC50),
Dermal
(LD50)
Highly
flammable,
NFPA
Category 3
2%/13%
Explosive with
chloroform and
base; reacts
violently with
nitric acid
Waste Considerations
The only stream going to waste is the bottoms of T-600. It is 92.6% water by mass, and 7%
formaldehyde by mass. This stream will go to industrial wastewater treatment outside of the
process. All other streams are burned as plant fuel or are sold.
Safety Precautions
Maintenance workers, engineers and other employees working in the system should be
wearing the proper protective equipment to ensure safety in the plant from high pressure,
high temperature and corrosive environments that are prevalent in the system.

Process Hazard Analysis (PHA)
Process Unit Hazard Effects Severity Likelihood Risk Current Control Verifications
Column High
pressure
buildup,
Leak
Shock,
Explosion,
Leak
Major Possible High Rupture cap to
prevent pressure
overload
Test, analysis and inspection
and training for employee
Compressor High power
and high
pressure
Shock,
Leak/Explosion
Major Possible High Shut off switch,
metal components
grounded/guarded
Test, analysis and
inspection, certification,
maintenance
Carbon bed High
pressure
buildup
Shock,
Leak/Explosion
Major Possible High Rupture cap to
prevent pressure
overload
Test, analysis and inspection
with maintenance
Pumps/Mixer Excessive
Pressure
Pipe rupture, Major Possible High Pressure vessels
leak before burst
Shutoff Activated
automatically if
fire is detected
Station attendant
trained in
inspection
Maintenance
System tests
Regular system training

61
Active
Electrical
Components
Electric Shock
Burns
Heart Problems
Minor Unlikely Moderate Metal
components
grounded and
insulated.
Station attendant
trained in
inspection
Active charge
components
covered
Fence
surrounding
system
Maintenance
Regular system training for
employees
Holding
Tanks
Excessive
Pressure due
to Vapor
Expansion
Vapor Release
-Hazardous if
inhaled or
absorbed
Major Unlikely Moderate Shutoff activated
automatically.
Inspection
procedure
Tank
Degradation
Chemical
contamination
Hazardous
exhaust fumes
emitted
Harmful if
inhaled
Minor Rare Low Material chosen
that is resistant to
corrosion from
most materials in
system.
Attendant trained
in inspection
Regular system training for
employee.
Maintenance.
System Tests.

62
Cooler High Cold
Temperature
Burns Major Unlikely Moderate Pipes insulated to
extreme
temperatures
Test, analyze, get pipes
certified and quality control
Reactor Fire Hazard Spontaneous
Combustion
Major Possible High Shutoff is
activated
automatically.
Pipes and
pressure vessels
insulated
System Check and
Maintenance
Health
Hazard
Dust Inhalation Major Possible Moderate Weekly
Inspection
Maintenance
Figure 28: Process hazard analysis table of components in the acetone retrofit system

Figure 29: Process hazard analysis matrix to determine risk
Discussion of the Process Hazard Analysis:
The process hazard analysis summary above shows how the different components of the
system are in terms of safety. The risk was obtained by using the matrix above between the
likelihood of the event and the impact of the event.
The high risk conditions were high pressure systems and a fire hazard from the reactor. The
high risks were calculated from the probability and severity of the accidents by using Figure
29. The accumulation of high pressures and high temperature could lead to pipe and system
ruptures. Therefore, rupture disks were added as means to remedy and reduce the risk of
accidents happening. Attendants and inspection training would be provided to insure proper
functioning of the columns, compressor, carbon bed, and pumps/mixer. These modifications
to the plant would save the company money from not having to pay for repairs that are
much more drastic than a blown rupture disk. A complete system shutdown or malfunctions
in the system would be more expensive than adding these safety measures.
Similarly, a heating problem with the reactor could burn workers. The use of a temperature
control system should prevent any major temperature overloads in the reactor.
The other risks are not as high but are still as serious need to be considered. For example, a
leak from the cooler could cause burns and vapor/fumes inhaled from the reactor could
cause major health problems. That is why simple pipe insulation could reduce likelihood of
malfunctioning and weekly inspection of the reactor should drastically reduce chances of
vapor evaporation/leaks from the reactor. Insulation and the grounding of the metal
components of the pumps and mixers may circumvent the problem of electric shocks and
burns for workers. In addition to this, a trained station attendant should be inspecting the
system regularly. However, pipe and pressure vessel insulation is only insulated to 150 o
F,
which is still a hazard to workers. Additional measures could later be implemented if these
safety modifications prove not to be enough. It is in the company’s best interest to protect

64
the lives and well-being of its workers even if it means at a slightly higher price. This
ensures a safe working place for workers. It also avoids any economic and public
repercussion that may occur if there is an equipment malfunction or worker injury due to
poor safety design. For every injury prevented, the company saves itself from being
responsible for the injury, having poor publicity, and providing medical compensation.
Standard Operating Procedure (Startup & Shutdown
Procedure):
Startup Procedure
This startup operating procedure will start with the Column 100 system.
(1) We will initially open the control valve prior to the column to allow the stream to
flow.
(2) Turn on the feed pump which is fed initially by a water reservoir. Then allow the
column to fill up to 3 feet.
(3) Turn the feed pump off.
(4) Turn the reboiler on and feed the utility to the condenser so any vapors are
condensed.
(5) Once the reboiler reaches a temperature of 228 o
F, turn on the reflux pump and
operate at total reflux until the trays reach the specified temperatures:
Table 11: Temperatures that each of the trays in Tower 100 must reach during startup at total reflux
Trays Temperature (o
F)
Condenser 132.7
1 148.8
2 151.3
3 155.1
4 170.7
5 203.8
Reboiler 228.5
(6) Once trays reach temperature within 5 o
F, use a utility stream instead of a process
stream for startup; turn on the feed heat exchanger.
(7) Turn the feed pump on.
(8) Turn on the distillate and bottoms product pumps.

65
(9) Check for leaks.
(10) Repeat steps 1 through 9 for the acetone waste feed stream.
(11) Monitor the column until it reaches steady state.
Shutdown Procedure
This shutdown operating procedure will start with the Column 100 system.
(1) Turn the feed pump off and let the column empty.
(2) Turn the control valve of the process stream that acts as a utility for the heat
exchanger.
(3) Close the feed control valve
(4) Let remaining process run until liquid holdup in the column is emptied.
(5) Turn off the condenser and reboiler
(6) Turn off the reflux, distillate and bottoms pump.
(7) Check for maintenance problems.
Process Control Strategies
Reactor
Temperature control is used to control the amount of high pressure steam entering the shell
side of the final heat exchanger before the reactor. This ensures that the temperature in the
reactor stays approximately constant at the design temperature so that the selectivity and
conversion which were used to model the reactor are valid. The reactor level was also
controlled via a valve in the line before the reactor so that the reactor does not overflow and
leak out flammable materials. Temperature control is used for the reactor to prevent the
reactor contents from getting too hot. A thermocouple will be couple to a valve in the hot oil
line which will cause the valve to close if the temperature in the reactor exceeds 700 o
F.
Fire Heater for Hot Oil System
Temperature control is used to control the amount of fuel being fed to the fire heater so that
the oil temperature is at a desired set point. The temperature control adjusts a valve that
controls the amount of fuel being fed to the heater.
Column System (including reboiler and condenser)
For the feed streams to each of the separators, pressure and temperature control were used to
ensure that the feed was entering the column at a pressure and temperature similar to the
feed tray (as designed) to ensure that the columns work efficiently and consistently meet
specifications. Temperature control on the bottom tray of the column was used to control the
flow of the utility in the reboiler to ensure that the columns have a high enough vapor flow
to achieve the separation. Each reactor also has level control at the bottom of the column to

66
ensure that there is the minimum of 3 feet of liquid holdup so that the column does not run
dry. The level sensor controls the bottoms product valve and closes this valve if the liquid
level in the tank drops below the minimum 3 feet.
Reflux Drums
Level control is used for each of the reflux drums to ensure that the reflux ratio stays
approximately constant at the value for which the column was designed. The level sensor
measurement is used to control the valve of the distillate product stream. The reflux pump
will have a flow controller with it to ensure constant flow back to the column. Therefore, the
level controller was chosen to be connected to the distillate product flow to ensure that the
reflux drum level stays constant. A pressure and level alarm should also be installed for each
of the reflux drums in the case of excess vapor or liquid accumulation.
Storage/Holding Tanks
Pressure and level alarms should be installed for each holding tank in case of excess vapor
or liquid accumulation.
Mixing Points
Each of the streams entering the three mixing points were designed to be at the same
pressure. At each mixing point, the valve controlling the pressure of one of the streams was
controlled by the pressure measurement of the other stream in order to ensure that the
pressures entering each mixing point were the same. For example, if the measured pressure
of stream 8 is lower than the pressure of stream 9, then the controller would decrease the
opening of the valve for stream 9 to drop it to the same pressure as stream 8.
Reflux Pumps
All reflux pumps have a flow controller that controls the valve after the reflux pump to
ensure that the flows of the reflux stream are constant to each of the columns. This ensures
that the column has high enough liquid flows to ensure the column will meet specifications.
Check valves should be installed after each reflux pump to prevent backflow into the pump.
Distillate Pumps
The valve before the distillate pump is controlled by the level in the reflux drum to prevent
liquid accumulation in the drums. Because all of the distillate pumps are centrifugal pumps,
there needs to be a low level alarm on each pump to ensure that the pumps do not run dry.
Check valves should be installed after each distillate pump to prevent backflow into the
pump.

67
Bottoms Pumps
The valve after each of the bottoms pumps is controlled by the level in bottom of the
column to ensure that the liquid holdup in the tank is at or above the minimum. Check
valves should be installed after each bottoms pump to prevent backflow into the pump.
Feed Pumps
Each of the feed pumps are flow controlled to ensure constant flow into each of the
columns. Check valves should be installed after each feed pump to prevent backflow into
the pump.
Process Stream HEX
The temperature of the process stream leaving a heat exchanger is used to control the flow
of the utility to the heat exchanger. This temperature control scheme cannot be used for
process-process heat exchangers because the flow of the process streams need to remain
constant. However, the temperatures of the process streams entering and leaving the
process-process heat exchangers will be closely monitored.
Compressor
The compressor after Column 400 is used to compress the hydrogen-propylene mixture
leaving the column. A valve before the compressor will be connected to a level controller
with the reflux drum to control the flow to the compressor. If the liquid level high in the
reflux drum is high, the valve with open more to prevent backup of the vapor in the system.
Carbon Beds
There will be a thermocouple placed before the carbon bed system to ensure that the acetone
stream being fed to the bed is cool enough so that there is no disruption in the adsorption of
the impurities to the activated carbon. If the stream is too hot, there will be poor adsorption,
the beds will be ineffective, and the product being sent to the consumers will not be up to
standards. There will be pressure control on the feed stream in order to be able to drop the
pressure and thus drop the temperature of the stream if it is entering the system at a high
temperature.
Special Design Considerations
● Nickel was used for the holding tanks because it seemed to be the cheapest material
that was still resistant to the corrosive behavior of the components in the system like
acetone, mesityl oxide and formaldehyde. Carbon steel was used for most of the
process while stainless steel was used for the trays of each of the distillation columns
and for the components of the design that had contact with acetic acid. Since most
materials could not be stand the corrosiveness of acetic acid, stainless steel was
implemented to circumvent the problem.
● To avoid stainless steel equipment, acetic acid was removed from the acetone waste
feed stream with the first distillation column (T-100) so that no acetic acid made it to

68
the rest of the process. The acetic acid waste stream was sent to the fired heater to
be used as fuel. This avoided having to store corrosive and potential toxic waste.
● Chemicals like acetone, mesityl oxide and other components in the streams are toxic
and have environmental consequences if the system does have leaks. That is why
control valves and holdup tanks are designed in locations where there is a possibility
of a severe malfunction or leak contributing to a health hazard. Implementation of
holdup tanks were placed in areas where if a malfunction or breakdown of a column
ahead or behind could cause the entire system to fail; a quick check was to see if a
distillation column were to fail, what would happen and where (if any) would a
holding tank be placed to avoid the crisis.
● Process streams were used in as many places as possible rather than using utilities
when designing heat exchangers. By crossing streams, roughly $750,000/yr were
saved in utility costs. This also means that energy was saved in generating utility
streams, which are heated using a fossil fuel or carbon based chemical like the waste
streams feeding the acetone retrofit process. By saving energy and conserving fuel,
the acetone retrofit process will have a smaller environmental impact related to
greenhouse gas emissions compared to if process streams were not crossed.
● Heat exchangers were designed to withstand 50 psi if they used LLPS so that they
could use LPS at 50 psi if hotter temperatures are needed during operation. It adds
flexibility to the temperatures that the system can be run at so that corrections for
heat loss can be made.

Section Seven: Capital
Estimate

Capital Estimate
Basis
After evaluation, the total plant capital cost (including IBL and OBL) totaled $194 million
as seen in Table 12. This is based on a plant capacity of 189 MMlb acetone/year. The
Aspen Economic Software was used to estimate the capital cost for all equipment except the
two vacuum systems and the refrigeration system. This includes all towers, holding tanks,
reflux drums, heat exchangers, pumps, and the compress. For the vacuum and refrigeration
systems, externally-provided data was used to estimate the capital cost. An escalation of 3%
per year over three years and a project contingency of 20% was assumed in order to make
the capital cost estimation.

71
Summary of Capital Cost Calculations
ICARUS List of Assumptions
Below is a list of assumptions made in ICARUS to serve as the basis for economic analysis
when pricing equipment. This is repeated in the Appendix, Economic Calculation
Methodologies (ICARUS Inputs) along with the assumptions made in each equipment
specification sheet.
Table 12: Capital cost summary sheet for the base case plant design

72
General Specs
● Process description: Proven process (none of the information is proprietary, and all
of the separations have been done before)
● Process complexity: Typical (Azeotropes are common- this was the only major
problem we faced. Our process used mostly standard distillation procedures for
separations)
● Process control: Digital (We will not have manual control processes)
● Plant addition: Adjacent to existing (There is an existent plan that produces our feed
streams adjacent to this plant)
● Estimated start date: Jan 18, 2016 (Assumed to be the beginning of semester)
● Soil conditions: Sand/clay [19]
● Pressure Vessel Design Code: ASME (specified)
● Vessel diameter: ID (specified)
● P and I design level: Full (specified)
Investment Parameters
● Capital escalation: 0 (This will be specified and added into the capital estimate
which includes ICARUS and other capital costs)
● Facility type: chemical process facility (acetone is not a specialty chemical,
pharmaceutical, or food product)
● Operating mode: 24 hrs/day (assumed)
● Length of start-up period: 20 weeks (specified as default)
Discussion
The two most expensive systems in terms of capital are Tower 300 (Acetone/Methanol
Vacuum Tower) and the IPOH Reactor.
The Tower 300 system is expensive mainly because of the large size of the tower, which is
required for the difficult acetone-methanol separation. This cost is much smaller than it
would be if the originally-planned pressure-swing distillation system was installed, because
that system would have required two extremely large towers. Tower 300 requires a vacuum
system to be installed; but even with this cost, the system is cheaper than the pressure-swing
system. A possible alternative to lower the capital cost for Tower 300 would be to use an
extractive distillation column. Further research would need to be done to see if the capital
savings for using extractive distillation would make up for the fact that a fresh stream of an
entrainer would need to be purchased for the system. .Another large capital expense
associated with the Tower 300 system is the holding tank. However, a large holding tank is
necessary for safety reasons.
The IPOH reactor system features a large capital cost because the reactor also acts as a heat
exchanger. The reactor required a large area for the proper heat transfer to occur so that the
reactor stays at the optimal temperature for conversion and selectivity reasons. A large
holding tank is also included in the IPOH reactor system, which greatly increases the capital
cost but is necessary for safety reasons. The final reason the IPOH capital cost is high is due

73
to the presence of a fired heater system, which is required to provide heat to the endothermic
reaction.
The capital estimate also took into account indirect costs, which included engineering costs,
field management/representatives, rack/sewers, tools, temporary structures, rentals, and
surplus materials. The rack/sewers were estimated to be 20% of the total direct cost of each
piece of equipment. ICARUS provided a lump sum indirect cost, which accounted for
installation and engineering costs. The final contribution to the indirect cost was the
difference in ICARUS between the direct total and IBL direct total costs, which represents
costs not accounted for in equipment capital and installation. These three values were
proportionally distributed across all pieces of process equipment, and contributed $26
million to the capital estimate.

Section Eight: Operating
Costs

Acetone Retrofit_Senior Design

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