192
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.

Acetone Retrofit_Senior Design

Embed Size (px)

Citation preview

Page 1: 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.

Page 2: Acetone Retrofit_Senior Design

1

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

Page 3: Acetone Retrofit_Senior Design

2

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

Page 4: Acetone Retrofit_Senior Design

3

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

Page 5: Acetone Retrofit_Senior Design

4

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

Page 6: Acetone Retrofit_Senior Design

5

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

Page 7: Acetone Retrofit_Senior Design

Section One: Background

Page 8: Acetone Retrofit_Senior Design

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

Page 9: Acetone Retrofit_Senior Design

8

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]

Page 10: Acetone Retrofit_Senior Design

Section Two: Process

Description

Page 11: Acetone Retrofit_Senior Design

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.

Page 12: Acetone Retrofit_Senior Design

11

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).

Page 13: Acetone Retrofit_Senior Design

12

Block and Process Flow Diagrams

Figure 4: The block flow diagram for the acetone retrofit plant

Page 14: Acetone Retrofit_Senior Design
Page 15: Acetone Retrofit_Senior Design

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.

Page 16: Acetone Retrofit_Senior Design

15

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

distillation column contains 6 actual trays. Following a flow meter and a control valve, the

feed stream (14.9 psia, 176.9 °F) enters the column at tray 4. The reflux ratio is set to 1.002,

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)

Page 17: Acetone Retrofit_Senior Design

16

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

distillation column contains 43 actual trays. Following a flow meter and a control valve, the

feed stream (25.2 psia, 83.4 °F) enters the column at tray 25. The reflux ratio is set to 107.2,

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

Page 18: Acetone Retrofit_Senior Design

17

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

Page 19: Acetone Retrofit_Senior Design

18

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

Page 20: Acetone Retrofit_Senior Design

19

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

Page 21: Acetone Retrofit_Senior Design

Mass Balance

Page 22: Acetone Retrofit_Senior Design

21

Mass Balances (Continued)

Page 23: Acetone Retrofit_Senior Design

22

Mass Balance Continued

Page 24: Acetone Retrofit_Senior Design

23

Energy Balance

Page 25: Acetone Retrofit_Senior Design

24

Page 26: Acetone Retrofit_Senior Design

25

Page 27: Acetone Retrofit_Senior Design

Section Three: Process and

Equipment Design

Page 28: Acetone Retrofit_Senior 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

Page 29: Acetone Retrofit_Senior Design

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.

Page 30: Acetone Retrofit_Senior Design

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

Page 31: Acetone Retrofit_Senior Design

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

Page 32: Acetone Retrofit_Senior Design

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 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 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

Page 33: Acetone Retrofit_Senior Design

32

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 9.

Using the knee of the curve method, the optimum number of theoretical trays was found to

be 19. With the number of theoretical trays determined, the next thing to be determined was

the feed tray location. Figure 10 shows a plot of reflux ratio versus feed tray location, which

was used to determine the optimum feed tray location.

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

Figure 10: Determining the feed tray location for Column 200

Page 34: Acetone Retrofit_Senior Design

33

From Figure 10, the optimal feed tray was found to be tray 9, as this minimizes the reflux

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

Page 35: Acetone Retrofit_Senior Design

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.

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 13.

Figure 12: XY diagram for acetone and methanol at 1.5 psia

Page 36: Acetone Retrofit_Senior Design

35

Using the knee of the curve method, the optimum number of theoretical trays was found to

be 22. With the number of theoretical trays determined, the next thing to be determined was

the feed tray location. Figure 14 shows a plot of reflux ratio versus feed tray location, which

was used to determine the optimum feed tray location.

From Figure 14, the optimal feed tray was found to be tray 18, as this minimizes the reflux

ratio.

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

Figure 14: Determining the feed tray location for Column 300

Page 37: Acetone Retrofit_Senior Design

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 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 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

Page 38: Acetone Retrofit_Senior Design

37

Column Sizing and Feed Tray Determination

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

Page 39: Acetone Retrofit_Senior Design

38

Feed Condition

The pressure at the condenser was set to 2.00 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 8.10 psia and a temperature of 111.5 °F, with a vapor fraction of 0.0852.

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 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.

Using the knee of the curve method, the optimum number of theoretical trays was found to

be 28. With the number of theoretical trays determined, the next thing to be determined was

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

Page 40: Acetone Retrofit_Senior Design

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.

Figure 18: Determining the feed tray location for Column 500

Page 41: Acetone Retrofit_Senior Design

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

The pressure at the condenser was set to 14.70 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 14.90 psia and a temperature of 176.8 °F, with a vapor fraction of 0.0409.

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 20.

Figure 19: XY diagram for isopropanol and water at atmospheric pressure

Page 42: Acetone Retrofit_Senior Design

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

Page 43: Acetone Retrofit_Senior Design

42

Feed Condition

The pressure at the condenser was set to 64.00 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 73.40 psia and a temperature of 233.8 °F, with a vapor fraction of 0.0043.

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 22.

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

Page 44: Acetone Retrofit_Senior Design

43

Using the knee of the curve method, the optimum number of theoretical trays was found to

be 36. With the number of theoretical trays determined, the next thing to be determined was

the feed tray location. Figure 23 shows a plot of reflux ratio versus feed tray location, which

was used to determine the optimum feed tray location.

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

Figure 23: Determining the feed tray location for Column 700

Page 45: Acetone Retrofit_Senior Design

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,

Page 46: Acetone Retrofit_Senior Design

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 52.9oF. Because the reactor needs to be run at

650oF, hot oil at 750oF must be used as the heat transfer fluid since the temperature

approach is 100-200oF when heating above 600oF. Dowtherm oil was chosen as the heat

transfer fluid, and it has an overall heat transfer coefficient of approximately 15 BTU (hr ft2 oF)-1 [24]. The desired outlet temperature of the hot oil needed to be about 730oF to achieve

a ∆TLM less than 90oF (730oF gives ∆TLM=81oF), and to stay above the 100oF 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 750oF 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

Page 47: Acetone Retrofit_Senior Design

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 oF. An operating temperature of 650 oF was chosen. The

values of conversion at 600 and 700 oF were averaged to find the conversion of 0.90 at 650 oF. 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 oF. 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 oF, 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

Page 48: Acetone Retrofit_Senior Design

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 oF was hot oil heated to 750oF.

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

Page 49: Acetone Retrofit_Senior Design

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.

Page 50: Acetone Retrofit_Senior Design

Section Four: Alternative

Cases

Page 51: Acetone Retrofit_Senior Design

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

Page 52: Acetone Retrofit_Senior Design

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

Page 53: Acetone Retrofit_Senior Design

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.

Page 54: Acetone Retrofit_Senior Design

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

Page 55: Acetone Retrofit_Senior Design

Section Five: Outside

Battery Limit

Page 56: Acetone Retrofit_Senior Design

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

Page 57: Acetone Retrofit_Senior Design

Section Six: Environmental,

Safety and Special Design

Considerations

Page 58: Acetone Retrofit_Senior Design

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.

Page 59: Acetone Retrofit_Senior Design

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

Page 60: Acetone Retrofit_Senior Design

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.

Page 61: Acetone Retrofit_Senior Design

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

Page 62: Acetone Retrofit_Senior Design

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.

Page 63: Acetone Retrofit_Senior Design

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

Page 64: Acetone Retrofit_Senior Design

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 oF,

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

Page 65: Acetone Retrofit_Senior Design

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 oF, 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 (oF)

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 oF, 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.

Page 66: Acetone Retrofit_Senior Design

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 oF.

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

Page 67: Acetone Retrofit_Senior Design

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.

Page 68: Acetone Retrofit_Senior Design

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

Page 69: Acetone Retrofit_Senior Design

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.

Page 70: Acetone Retrofit_Senior Design

Section Seven: Capital

Estimate

Page 71: Acetone Retrofit_Senior Design

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.

Page 72: Acetone Retrofit_Senior Design

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

Page 73: Acetone Retrofit_Senior 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

Page 74: Acetone Retrofit_Senior Design

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.

Page 75: Acetone Retrofit_Senior Design

Section Eight: Operating

Costs

Page 76: Acetone Retrofit_Senior Design

Overview

Operating costs for this project include:

1) Raw Materials - copper on alumina catalyst (feed streams are assumed to have no

value)

2) Byproducts - none of the byproducts are being sold in the base case, so their values

were set to zero.

3) Utilities - Hot Oil, HPS, MPS, LPS, LLPS, CW, Refrigerant, power

4) Fixed Costs - labor, overhead, repair and maintenance, property tax, insurance

The relative price of each operating cost is shown below:

Table 13: Summary of operating costs by type

Operating Cost

Type

Cost (¢/lb acetone product)

Raw Materials 0.08

Byproducts 0

Utilities 9.73

Fixed Costs 11.37

Raw Materials

The only raw material cost in the process is the catalyst. Over the life of the project, 6

reactor volumes of catalyst will be required, and its current cost is $10/lb.

The raw materials are an isopropanol and an acetone waste stream, and are valued as a fuel

source for the existing process. The raw materials were considered to have no value in the

economic analysis, but they will be considered as a different opportunity for the company,

where it is valued as a utility.

Fixed Costs

Labor costs were based off of 60 employees, covering two shifts. This was an estimate

based on having 9 unit operations in the system. This encompasses people in charge of

plant maintenance, one operator per unit operation, and people to take samples and analyze

the data.

Overhead costs were calculated by multiplying the labor cost by a factor of 1.5. This covers

soft costs associated with project support, such as marketing, administration, and support

staff.

Page 77: Acetone Retrofit_Senior Design

76

Repair and maintenance costs were assumed to be 1.5% of the capital costs. Property taxes

were accounted for as a fixed cost while sales and income taxes were accounted for in the

cash flow sheets. Property tax was 2.5% of the capital cost. Insurance was assumed to be

0.5% of the capital costs. This covers casualties, property damage, and pollution liability.

Utilities

The power utilities were used to power the process pumps and compressors. The

heating/cooling utilities available spanned from -150 to 750 oF. The assumed price of each

utility is shown in Appendix, Economic Calculation Methodologies (ICARUS Inputs).

These were calculated based on a natural gas price of $2.5/MMBTU. Although the current

price of natural gas is $1.86/MMBTU, $2.5/MMBTU was used as a safety factor. The

historical natural gas prices show price fluctuations of at least $0.5/MMBTU in any single

year. The natural gas futures predict that the price will increase to $6/MMBTU by 2025

[31]. The combination of these two factors resulted in us estimating the price of natural gas

to be higher than the current value.

The heating/cooling utilities were used for the reactor, process heat exchangers, reboilers,

and condensers. Each utility has a different cost, which allowed us to optimize the use of

each by minimizing price. Heat exchangers with large duties can be split into multiple heat

exchangers in series to reduce the use of expensive utilities. This is done by adding another

heat exchanger using a different process once the process stream is heated/cooled to the

range of a less expensive utility. Utility costs were also minimized by crossing process

streams where one needed to be heated, and the other needed to be cooled. The largest

energy savings seen by crossing streams was from cooling the reactor effluent by preheating

the stream going into T-700. This saved 4.34MMBU/hr. The cheapest utility was always

chosen while satisfying the temperature approach limits for each temperature range.

Page 78: Acetone Retrofit_Senior Design

77

Table 14: Summary of the power usage for the compressor and each of the pumps in the plant

Page 79: Acetone Retrofit_Senior Design

Table 15: Summary of the energy usage per hour of all reboilers, condensers, and heat exchangers in the plant

Page 80: Acetone Retrofit_Senior Design

Section Nine: Economic

Evaluation

Page 81: Acetone Retrofit_Senior Design

Basis

The basis used for the economic evaluation is shown below. The values were provided from

the company’s management. The targets for designing a process that is competitive with

other ones the company is considering are: ATROR ≥ 20% and NPV ≥ $30MM.

Plant Economics

● Plant Operating Time: 8400 hours per year

● Plant Startup: January 2019

● Project Start: March 2016

● Project Life: 15 years from start-up

● Capital Spending: 15%/35%/50%

● Market Build: 40%, 75%, 100% in 1st 3 years of production

● SG&A: 2% of sales

● Income Tax Rate: 38%

● Working Capital: 10% of revenues

● Depreciation: MACRA

● Project Discount Rate: 9%

● Escalation (Inflation): 3% per year

Fixed Costs

● Plant Labor Salary + Benefits: $85,000/man-yr

● Plant Overhead: 1.5 x Plant Labor

● Repairs and maintenance: 1.5% of capital

● Property tax: 2.5% of capital

● Insurance: 0.5% of capital

Future Prospects for the Acetone Market

The consumption of acetone is expected to increase in the future based on the demand of the

products that it is used for, predominantly BPA, methyl methacrylate, and solvents. In

2014, BPA accounted for 31% of all global acetone production, and it is expected to

continue to surpass all other categories by 2020 [25]. The largest market players in the

acetone industry are: Dow chemical company, BASF, INEOS Phenol GmBH, CEPSA

QUIMICA, Shell Chemicals, Mitsui Chemicals, Reliance Chemicals, Honeywell and LG. A

company in Shanghai China recently started up the world's largest phenol and acetone plant,

meaning China will be importing less acetone than historically.

Page 82: Acetone Retrofit_Senior Design

81

The estimated trend in consumption and year-over-year growth both indicate that the

demand for acetone will be consistently increasing in the future. Because acetone is used as

a precursor to BPA, the demand for acetone will likely not drop significantly over the 18

year period that this plant will be in operation [25]. Seventy-three percent of BPA is used to

make polycarbonate resins. Twenty percent of these resins are used in the automotive

industry, and another 20% is used in making media products like CDs and DVDs. Because

of the market shift away from tangible media products, this is not an area of consumption

that should be depended on for demand. However, the automotive industry is increasingly

replacing glass with polycarbonate plastic, which is projected to increase in the future [25,

32]. According to a Research and Markets publication, the BPA market is projected to grow

between 2014 and 2019 [25].

Chemical Commodity Historical and Future Pricing

The acetone price increases at a higher rate than if its price increased only due to inflation

[28]. A line was fit to the historical acetone price to project what it will be in the future

because no commodity was found to correlate to acetone prices to use for future predictions

(crude oil correlation R=0.349, natural gas correlation R=0.252, and cumene correlation

R=0.448. Data shown in Appendix, Price Correlation Curves). We estimate that its price

will increase in the future for two main reasons:

Figure 30: Global acetone consumption forecast through year 2020 [25]

Page 83: Acetone Retrofit_Senior Design

82

1) Acetone is a byproduct of cumene, which is petroleum based. The historical

prices correlate with a linear fit having an R=0.95 (data shown in Appendix, Price

Correlation Curves). The futures for crude oil predict that the price will drop by about 15%

by the year 2025 [31]. However, the cost of cumene was not seen to correlate with

acetone’s price. This means that the price of acetone will likely be driven by supply/demand

more than the cumene price.

2) The demand for acetone will likely grow, based on the analysis given in the

previous section. Combined with the decrease in acetone production, this will increase the

global price of acetone.

Figure 31: Prediction of future acetone prices based on historical data

Page 84: Acetone Retrofit_Senior Design

83

The historical prices for hydrogen and PPE are also shown, and future projections are

shown. Hydrogen and PPE are considered as additional products to sell in an alternative

case. The double dashed lines show the price increasing by the inflation prices each year,

and the single dashed lines are linear fits to the data showing how the prices increase over

time. In 2019, when the process starts up, the price projection for PPE is lower than

hydrogen. However, by the end of the 15 years of operation, hydrogen is less valuable than

PPE. This illustrates that hydrogen’s price does not follow the domestic inflation rate, and

may not be as valuable as PPE because our mass flow of PPE is two times larger than our

flow of hydrogen.

Figure 32: Prediction of future hydrogen and propylene prices based on historical data

Page 85: Acetone Retrofit_Senior Design

84

Basis for Utility Costs

Figure 33: Historical natural gas prices along with forecasted prices up to 2025 [29]

The plot of natural gas prices shows that it peaked in 2005, and has been on a steady decline

since then because of technological improvements of hydraulic fracturing. The inset zooms

in on the price trend since 2014, which also shows a 50% decrease in price over the two

year period. The current price for natural gas is $1.96/MMBTU, so utility prices were based

off of this price. The price has been decreasing for the past decade, but World Bank predicts

that it will triple within the next 10 years.

The historical or future price of isopropanol (the major component in the feed stream other

than acetone) was not factored into the economic evaluation. This is because its current

value is derived from being used as a fuel, rather than as a product to sell. Its heat of

combustion will not change over time like price would, so its value is based upon the price

of natural gas that it saves from being used. The price of natural gas is predicted to increase

from the current price of $2/MMBTU, to $6/MMBTU in 2025 [31]. This means that the

feed streams to the acetone retrofit process will become more valuable with time. The

avoided costs of using the feed streams as fuel instead of natural gas to produce high

pressure steam will be larger in the future, and will be taken into consideration when

comparing the company’s options, but will not be part of the acetone retrofit cash flow.

Page 86: Acetone Retrofit_Senior Design

85

Base Case Economic Analysis

The total plant capital cost is $194,136,000. The operating costs for the plant are

$40,000,000 per year. Figure 34 shows the summary of the cash flow sheet for the base

case at an acetone price of 40 ¢/lb.

Figure 34: Cash flow summary sheet for the base case design

As can be seen from Figure 34, the net present worth of the plant is $28.3 million, which

gives an After Tax Rate of Return (ATROR) of 11.51%. Management has informed the

plant design team that capital spending is constrained due to a major acquisition and that

there are currently a number of other attractive projects with NPVs of over $30 MM and

ATRORs better than 20%. Since the plant design does not meet this ATROR rate of 20%,

the design team does not recommend going forward with the construction of this acetone

Page 87: Acetone Retrofit_Senior Design

86

retrofit plant. The large capital cost of the plant, which can only increase with construction

issues and delays, along with the low ATROR makes this plant risky and economically

unfeasible for the company. Therefore, the design team had to decide between the two other

methods for dealing with the two waste streams- burning them to produce high pressure

steam, or selling them to WasteCo., a company who has offered to pay for the two waste

streams.

The value of the acetone and IPOH waste streams when they are burned to produce high

pressure stream for the current plant is approximately $10 MM per year. A summary of the

calculation of this value is shown in Table 16.

Table 16: Determination of the potential revenue associated with burning the waste streams for fuel

The lower heating value for the combined waste streams and their mass flow rates were

used to determine the MMBTU/hr of energy that can be obtained from the streams when

they are burned. Assuming a burning efficiency of 60% in the boiler and a price of high

pressure stream of 600 cents/MMBTU (based on a natural gas price of $2.50/MMBTU), the

value of the streams when they are burned was determined to be $10 MM/year. However,

this value does not take into account the fact that the reboiler system would need to be

upgraded in order to burn these streams (which would not currently meet environmental

standards). Assuming a boiler upgrade would cost 10% of the acetone retrofit in capital

costs, and would last for 15 years, the new profit from burning the streams becomes $8.7

MM/year.

WasteCo. is offering 15 ¢/lb and 12 ¢/lb for the waste acetone stream and waste IPOH

stream, respectively. As shown in Table 17, the company can make approximately 32.8

MM$/year by selling the acetone and IPOH waste streams to WasteCo. Since this option

does not include upgrading or constructing any additional infrastructure, no significant

deductions need to be taken out of this revenue.

Based on these numbers, selling the waste streams to WasteCo. is recommended over

continuing to burn these streams for fuel. In addition, based on the WasteCo. option revenue

of 32.8 MM$/year, the NPW of the acetone retrofit plant would have to be 540 MM$ to

Table 17: Summary of the revenue associated with selling the waste streams to WasteCo.

Page 88: Acetone Retrofit_Senior Design

87

compete with this option. An NPW of 540 MM$ corresponds to an acetone price of

$1.01/lb, which is 2.5 times the current price of acetone.

Sensitivity Analysis

Case 1: Not all of the product can be sold - 30 MMlb/yr surplus

This investigates the outcome of not being able to sell all of the acetone at 189 MMlb/yr

capacity. Acetone sales are assumed to be under yearly contract, meaning 30 MMlb/yr

cannot be sold for an entire year. This study will investigate what should be done with the

excess product, and how it would affect the profitability of the acetone retrofit process.

How to Handle the Surplus

There are two options:

1) Store the acetone product until a buyer is identified and a contract is signed

2) Use the acetone product as fuel in the meantime to reduce operating costs

The process cannot be shut down as an Option 3 because it is essentially a waste treatment

facility. Shutting down the acetone retrofit process means that the existing process

producing the waste streams will be out of compliance with environmental regulations.

Option 1: Storage of 30 MMlb from one year of surplus operation.

Holding tank dimensions: H= 122 ft, D=80 ft

Holding tank cost: $144,590,000

Option 2: Use the surplus as fuel instead of storing it, and reduce utility costs.

Rather than spending the energy to separate acetone beyond what can be sold, 16% of the

feed stream will be diverted to the fuel boilers to produce high pressure steam, and the

utility requirements are also reduced by 16%. Over the course of a year, 2.23x107 lbs of the

acetone stream and 1.57x107 lbs of the IPOH stream will be sent directly to the fuel boilers

to produce 4.44x105 MMBTU of energy over the course of the year. This equates to $0.888

million saved from not purchasing natural gas but costs $10.5 million in lost product

revenue. The ATROR becomes 8.27%. The cash flow sheets for Option 1 and 2 are in the

Appendix, Sensitivity Analysis.

Discussion

The capital cost for the acetone retrofit process is $194,000,000, so adding the holding tank

nearly doubles the capital cost for the process. This changes the ATROR to 3.23%, and it

was originally 11.51%. It does provide the option of selling the product at a later date,

which would require 14 years to pay off the holding tank on a nominal dollar basis. The

second option produces an ATROR of 8.27%, and is less risky. Option 2 does neglect the

Page 89: Acetone Retrofit_Senior Design

88

possibility of violating environmental regulations and facing fines. The boiler plant in the

OBL does not need to be upgraded if the acetone retrofit is built, so sending the feed streams

back to the boilers violates the new environmental regulations that were driving the acetone

retrofit project in the first place. By diverting the stream, rather than sending it through the

process for storage, all of the equipment is running well below capacity, and at lower

efficiencies.

Recommendation

It is recommended that the feed stream is diverted from the system if the product cannot be

sold, and a partial boiler upgrade is completed when the acetone retrofit is being installed.

The boiler system can then safely burn a percentage of the feed streams in an emergency.

This will add cost, but keeps the company in compliance if all of the product cannot be sold.

Case 2: The Price of Acetone Changes

Based on the acetone pricing data from ICIS, the price of acetone used for the economic

analysis was 40¢/lb. However, the price of acetone could change for various reasons. First,

the demand for acetone could change. This could be caused by a government regulation on

BPA, which is formed by the condensation of acetone. The demand could also change due

to the development of a new technology that decreases production costs. The effect of

acetone price on ATROR is shown in Figure 35. In order for the plant to reach the ATROR

hurdle rate of 20%, the price of acetone would have to increase to about 55 ¢/lb.

Case 3: Capital Costs Increase

The plant could experience an increase in capital costs as a result of delays in construction

or any other unforeseen costs that are incurred during the construction process. Figure 36 is

Figure 35: Effect of acetone price on the ATROR (capacity and capital costs constant)

Page 90: Acetone Retrofit_Senior Design

89

a plot showing the effect of capital cost on ATROR. It can be seen from the plot that the

capital cost would need to be reduced by approximately $79 million (to $115 million) in

order for the plant to eclipse the ATROR hurdle rate of 20%. Since the sizing of all of the

capital equipment was optimized to reduce costs, achieving such a cut in capital cost is

unfeasible.

Case 4: The Price of Natural Gas Changes

The price of natural gas could continue to fluctuate in future years. The price of natural gas

has reached a 10-year low due to technology improvements in hydraulic fracturing. Natural

gas is a non-renewable resource that will become increasingly expensive to drill for as the

supplies decrease. The current domestic supplies are expected to last 84 years [30], but that

does not imply that the cost will remain at $2/MMBTU until the economical reserves are

drained. According to the World Bank, natural gas will reach $6/MMBTU by 2025 [31],

which needs to be considered for the future of the acetone retrofit process. Table 18 shows

the utility prices at two different prices of natural gas. Linear interpolation was used to

obtain utility prices at two additional prices of natural gas. Then, a plot of ATROR as a

function of natural gas price was produced. If the price of natural gas increases to

$6/MMBTU by 2025, the ATROR of the plant would drop to below 10%.

Figure 36: Effect of capital cost on ATROR (acetone price and capacity held constant)

Page 91: Acetone Retrofit_Senior Design

90

Costs Not Included in the Cash Flow sheets used in the Previous Section:

The OBL operating costs, as well as the fixed cost of outsourcing the catalyst regeneration

was not included in the cash flow sheets. One volume of catalyst would be regenerated

every six months by sending the catalyst to Eurecat Group’s Gulf Coast location for a cost

Table 18: Utility costs at two natural gas prices. Linear interpolation was used to obtain utility costs at

other natural gas prices.

Figure 37: Effect of natural gas price on ATROR (acetone price and capital cost held constant)

Page 92: Acetone Retrofit_Senior Design

91

of $1.50 to $6.00 per pound of catalyst. The OBL operating costs, as well as the catalyst

regeneration cost, would only make the economics of the base case worse and therefore not

including these costs did not have an effect on the final decision that the group made

regarding the acetone retrofit plant.

Page 93: Acetone Retrofit_Senior Design

Section Ten: PDRI

Page 94: Acetone Retrofit_Senior Design

PDRI Discussion

The PDRI or the Project Definition Rating Index is a measure of how completely or

thoroughly the project was addressed. In terms of the project, multiple PDRIs were

calculated to monitor the progress of the project after each gate. Each of the various

categories were given a different weighting system to calculate the overall PDRI.

Unfortunately, many of the categories or criterion in the PDRI were outside of the scope of

this project and thus the final score never really approached the optimal score of

approximately 250.

A couple of the categories were either not well-defined or within the scope of the project. A

category that was not within the scope of the project is Civil, Structural and Architectural

and Electrical components. Similarly, there were many miscellaneous components like

training requirements and CADD/Model requirements that are outside the scope of this

project. Other categories like Site Information, Infrastructure, Procurement Strategy and

Project Execution Plan were broadly introduced in the description of the project but were

not well-defined.

The figure below shows the progression of PDRI values as the project progressed

throughout the semester.

Figure 38: PDRI values throughout the design of the plant

Page 95: Acetone Retrofit_Senior Design

94

The following figure shows the reduction of the PDRI as the project progressed:

Figure 39: The breakdown of the PDRI into its three categories for each PDRI completed

Specific categories where the score was a ‘Level 5’ include:

PDRI Section I

Maintenance philosophy

Operating philosophy

Future Expansion Considerations

Expected project life cycle

Site Characteristics Available vs. Required

Dismantling and demolition requirements

Design for constructability analysis (Level 5)

PDRI Section II

Surveys and Soil Tests

Permit requirements

Process Safety Management

Utility Flow Diagrams

Piping System Requirements

Plot Plan

Line List

Page 96: Acetone Retrofit_Senior Design

95

Tie-In List

Piping Specialty Items List

Instrument Index

Equipment Status

Equipment Location Drawings

Civil/Structural Requirements

Architectural Requirements

Transportation Requirements

Logic Diagrams

Electrical Area Classifications

Substation Requirements/Power Sources Identified

Electric Single Line Diagrams

Instrument & Electrical Specifications

PDRI Section III

Procurement procedures and plans

Procurement Responsibility Matrix

CADD/Model Requirements

Project Accounting Requirements

Engineering/Construction Plan & Approach

Pre-Commissioning Turnover Sequence Requirements

Startup requirements

Training requirements

Page 97: Acetone Retrofit_Senior Design

Section Eleven: Outstanding

Issues

Page 98: Acetone Retrofit_Senior Design

Technical

1. Column 400 was difficult to converge in HYSYS when the XY diagram looks like it

should be an easy separation. The feed stream was not binary, so an XY diagram is

not a perfect way to predict the separation, and there may have been hidden,

complex interactions between molecules. It may have to do with the temperature

profile of the column spanning from about -100oF to 100oF.

2. The calculation to determine the value of burning the waste streams was a simple

calculation that assumed a boiler heat efficiency. Knowing the efficiency of the

boiler at the existing OBL will provide a more accurate value.

Economical

1. The project is too risky based on the ATROR of 11.51%, with a hurdle ATROR of

20%. However, the company was approached by WasteCo., who wanted to buy the

waste. This indicates that there is an economical process for handling the waste,

meaning that our design could be improved to become economical (assuming

WasteCo. also upgrades the acetone to product-grade).

2. No information was provided on the economics of the boiler upgrade project.

Without knowing the NPW or ATROR of the boiler project, it was difficult to make

an informed recommendation to management on which project to pursue.

3. The refrigeration costs were 6-20 times the cost of any other single utility in the

process. This should be used as a driver for the design knowing its negative impact

on project economics. This process was not designed to intentionally avoid

refrigeration, although the price of refrigeration was minimized in the heat

exchanger sizing calculations.

4. The cost of refrigeration for the gas separation was extrapolated from the data given

for refrigeration down to -150 F. A vendor would need to be contacted in order to

confirm that the pricing is similar to that modeled in the alternative case.

Environmental/Safety

1. Equipment improvements could be made to help prevent possible leakage of

material resulting from corrosion of process vessels. If needed, stainless steel could

replace carbon steel in more distillation columns.

2. Extra holding tanks could be implemented to prevent product accumulation or spills

in the event of a process malfunction.

3. Hydrogen gas is very flammable, and so is acetone. A worker safety training

program should be put in place to make all workers aware of the dangers of the

Page 99: Acetone Retrofit_Senior Design

98

flammable materials in the process. This should include training on what to do in

the event of a material leak, and how to respond to fires.

Page 100: Acetone Retrofit_Senior Design

Section Twelve: Conclusion

and Recommendations

Page 101: Acetone Retrofit_Senior Design

Conclusions

A summary of the economic analysis of the three possible choices- plant construction, boiler

upgrade to continue burning the waste, and selling to WasteCo., can be seen in Figure 40.

As previously described, due to the plethora of other worthy investments that are available

to the company, the acetone retrofit plant had to meet an ATROR of 20% and NPW of $30

MM to be considered. The acetone retrofit plant ATROR of 11.51% and NPW of $28.3 MM

does not meet these criteria. Sending the waste streams to fuel would save 10 $MM/yr in

utility costs for the current plant. However, this option requires capital costs associated with

a reboiler upgrade in order to meet environmental specifications. WasteCo. made an offer of

15 ¢/lb for the acetone stream and 12 ¢/lb for the IPOH stream. This offer would generate

31.5 $MM/yr in revenue for the plant with no large capital investments, which is why

selling to WasteCo. is the final recommendation.

Figure 40: The net profit per year for each of the three options for handling the waste streams

Upon the completion of the economic evaluation for the base case, it was concluded that

constructing the acetone retrofit plant is not the best option for the company.

Unless the price of natural gas stays at its current value of $2/MMBTU and the price of

acetone increases by 35.5%, selling the acetone and IPOH waste streams to WasteCo. is the

best option going forward. One area of the plant design that could be improved to make

producing acetone more profitable is reducing the total utility usage of the plant. Utility

costs account for over 37% of the total plant cost as can be seen in Figure 41.

Page 102: Acetone Retrofit_Senior Design

101

Specifically, reducing the amount of refrigeration that is used in the system would make the

plant more profitable. As can be seen from Figure 42, refrigeration accounts for

approximately 71% of the total utility costs.

Figure 41: Breakdown of Total Plant Costs (Capital + Operating) by type

Figure 42: Breakdown of utility costs by type

Page 103: Acetone Retrofit_Senior Design

102

The design team did not have pricing for refrigeration until the plant design was completed.

Therefore, the design was focused on designing the plant distillation columns to run at

atmospheric pressures or higher to eliminate the costs associated with vacuum systems.

However, the pricing for both vacuum and refrigeration systems shows that avoiding low

temperatures in the columns was more important than avoiding vacuum to produce an

economically feasible design.

Recommendations

The acetone retrofit should not be considered for construction unless the following

conditions are met:

Based on the sensitivity analysis

1. With capital costs and plant capacity held at the current values, the price of acetone

must increase to at least 55 ¢/lb over the entire plant lifespan for the plant to meet

the 20% hurdle rate

2. With plant capacity and acetone price fixed at the current values, the capital cost

must be reduced by 41% in order to meet the 20% hurdle rate.

3. An increase in natural gas prices by $4/MMBTU decreases the ATROR from

11.51% to 9.56%. Therefore, further studies to severely cut utility costs must be

completed.

4. A partial boiler upgrade must accompany the construction of the project to protect

against times when all of the product cannot be sold. This will enable the feed

streams to be temporarily burnt as fuel and remain in compliance with new

environmental regulations.

Based on the alternative case studies

4. Use vacuum distillation unless water prices fall drastically and utility costs are no

longer dependent on natural gas.

5. Sell hydrogen and PPE to increase the ATROR of the project, making the investment

less risky.

Based on the economic analysis:

6. Sell to WasteCo. unless acetone price reaches 101 ¢/lb, at which point the acetone

retrofit design will generate a NPW (divided over 18 years of the project life) equal

to the profits gained from WasteCo.

Page 104: Acetone Retrofit_Senior Design

103

Supporting Information for Recommendations:

Sensitivity Analysis

A sensitivity analysis, summarized in Table 19, was performed to find the conditions under

which plant construction would meet the ATROR hurdle rate of 20%.

Alternative Cases

Three different alternative cases were investigated. A total of three different methods for

separating acetone and methanol were analyzed (one of which was the base case), along

with the option of separating and selling hydrogen and propylene as side products.

The three ways to break the acetone-methanol azeotrope were vacuum, extractive, and

pressure swing distillation. Of these, the vacuum distillation method was the most

economical.

Pressure Swing

The pressure swing design required two columns and a large recycle loop to achieve the

desired purity of acetone. This resulted in towers that were sized to be larger than what

could be prefabricated and shipped to the plant. The high recycle flows also required higher

utilities to heat and condense the large flow through the towers than the other two cases.

These two design issues made the process unreasonable compared to the other two, and it

was not considered to be a viable option.

Table 19: Sensitivity Analysis summary

Page 105: Acetone Retrofit_Senior Design

104

Extractive

Extractive distillation required two columns and a new feed stream of 20,000 lb/hr of fresh

water. This additional flow rate increased the size of the columns, and added $34 million/yr

of operating costs, while the capital costs were essentially the same as the base case.

Vacuum

The vacuum distillation option required a single tower, and an additional vacuum pump

system to produce the low pressures needed for the system. It was used as the base case

because it offered the highest ATROR, had fewer unit operations to maintain, and did not

require an additional feed stream into the system. It posed the fewest technical and

economic risks.

Below are the theoretical acetone prices needed to reach a 20% ATROR with each

separation technique. Vacuum distillation requires nearly half of the price of acetone

compared to the next best case. Based on the acetone market and historical price trends,

reaching an acetone price of 54.4 ¢/lb is conceivable, but reaching the price of 93.5¢/lb is

not reasonable to expect within

the lifetime of the project.

Gas Separation to sell hydrogen

and propylene

An additional separator with

auxiliary equipment can be added

to the base case to purify

hydrogen from PPE and produce

two product-grade byproducts.

The price of hydrogen is 81 ¢/lb,

and PPE is 41 ¢/lb, both of which

are more valuable than the

acetone product. Assuming all of

the hydrogen and PPE is sold, the

additional profits increase the

ATROR to 13.55% from 11.51%.

Although it does not achieve an ATROR of 20%, it makes the investment less risky.

Figure 4: Acetone price required to meet hurdle rate for the base

case and two alternative cases

Page 106: Acetone Retrofit_Senior Design

Section Thirteen: References

Page 107: Acetone Retrofit_Senior Design

References

GATE 1 References

1. Bailey, William A., and Sumner H. Mcallister. Separation of By-Products from

Isophorone. Shell Dev, assignee. Patent 2,351,352. 13 June 1944. Print.

2. Weber, Markus, and Oto Schnurr. Continuous Process for Recovering Acetone from

a Waste Stream Resulting from Acetone Purification. Ineos Phenol Gmbh & Co. Kg,

assignee. Patent 7,416,645. 3 July 2003. Print.

3. Shell International Corporation. Production of Aliphatic or Cyclo-aliphatiic

Carboxylic Compounds. Patent 938,854. 9 Oct. 1963. Print.

4. J.A. Kent, “Handbook of Industrial Chemistry and Biotechnology,” vol 1, Edition

12, 314-390.

5. A. Seidel and M. Bickford, “Kirk-Othmer Encyclopedia of Chemical Technology,”

Acetone, Jan 14, 2011.

6. C. Maldqnado, J.L.G. Fierro, G. Birke, E. Martinez, P. Reyes, Conversion of

Methanol to formaldehyde on TiO2 supported Ag Nanoparticles,J. Chil. Chem Soc.,

55, 2010.

GATE 2 References

7.Luyben, William L., Comparison of Extractive Distillation and Pressure-Swing

Distillation for Acetone−Methanol Separation, Industrial & Engineering Chemistry

Research Ind. Eng. Chem. Res. 47.8 (2008) 2696-707. Web.

8. Gil, Ivan D., Diana C. Botia, Pablo Ortiz, and Oscar F. Sanchez. "Extractive

Distillation of Acetone/Methanol Mixture Using Water as Entrainer." Industrial &

Engineering Chemistry Research Ind. Eng. Chem. Res. 48.10 (2009): 4858-865. Web.

9. “Mass Transfer in VOC Adsorption on Zeolite:  Experimental and Theoretical

Breakthrough Curves,” Stephan Brosillon, Marie-Helene Manero, and, and Jean-Noel

Foussard Environmental Science & Technology 2001 35 (17), 3571-3575

10. “Fixed Bed Adsorption of Acetone and Ammonia onto Oxidized Activated Carbon

Fibers,” Christian L. Mangun, Richard D. Braatz,, James Economy,and, and Allen J.

Hall Industrial & Engineering Chemistry Research 1999 38 (9), 3499-3504

GATE 3 References

11. F. Pepe, C. Angeletti, S. De Rossi, and M. Lo Jacono, Catalytic Behavior and

Surface Chemistry of Copper/Alumina Catalysts for Isopropanol Decomposition,

1984.

Page 108: Acetone Retrofit_Senior Design

107

12. Allan, Lister, and Porter H. Thomas. Regeneration of Catalysts. British Petroleum

Co, assignee. Patent 3,041,290. 26 June 1962. Print.

13. Engelhard Material Data Sheet

14. M.D. Argyle and C.H. Bartholomew, Heterogeneous Catalyst Deactivation and

Regeneration, Catalysts, 2015, 145-269.

15. Huang, T. "ICIS Pricing, Acetone (Asia Pacific)." Chemical Industry News &

Chemical Market Intelligence. ICIS, Jan. 2014. Web. 19 Apr. 2016.

16. Jardin, Frederic, Eurecat. Message to the author. N.d. Email.

GATE 4 References

17. Engineers Ede. "Overall Heat Transfer Coefficient Table Charts and Equation |

Engineers Edge | Www.engineersedge.com." Overall Heat Transfer Coefficient

Table Charts and Equation | Engineers Edge | Www.engineersedge.com. Engineers

Edge, 2016. Web. 02 Apr. 2016.

18. Nedwick, Robert. Chemical Engineering Blue Book. 2016

GATE 5 References

19. Almanac, Texas. "Texas Almanac - The Source For All Things Texan Since 1857."

Soils of Texas. Texas Almanac, 2010. Web. 19 Apr. 2016.

20. Wintek Corporation. "Rotary Vane vs Oil Sealed Liquid Ring." Wintek -. Wintek

Corporation, 2014. Web. 19 Apr. 2016.

21. Stainless Steel Coroporation. "Stainless Steel Grades." - SS Material Grades.

Stainless Steel Corporation, 2015. Web. 19 Apr. 2016.

22. Elliot Group. "Single Stage Centrifugal Compressors■ ■." (n.d.): n. pag. Elliot

Group. Elliot Group, 2013. Web.

23. Moore, Richard L. "Implementation of DOWTHERM A Properties into RELAP5-

3D/ATHENA." Idaho National Laboratory (2011): n. pag. Web.

24. Dow. "Engineering and Operating Guide for DOWTHERM." A-to-Z Guide to

Thermodynamics, Heat and Mass Transfer, and Fluids Engineering AtoZ (2008): n.

pag. Web.

25. FMI. "Acetone Market: Global Industry Analysis and Opportunity Assessment, 2014

- 2020." FMI - Future Market Insights. FMI - Future Market Insights Market

Research Report, 2014. Web. 19 Apr. 2016.

26. ICIS. "Chemical Profile." Chemical Industry News & Chemical Market Intelligence.

ICIS, 2015. Web. 19 Apr. 2016.

27. Research and Markets. "Global Bisphenol A Market 2015-2019 - Rising Demand for

Polycarbonate Plastics with Bayer Material Science, Dow Chemical, LG Chemical,

Mitsubishi Chemical Holdings & Mitsui Chemicals Dominating." PR Newswire. PR

Newswire, 15 May 2015. Web.

28. Multpl. "US Inflation Rate by Year." Multpl. Multpl, 2015. Web. 19 Apr. 2016.

29. US EIA. "Natural Gas Weekly Update." U.S. Energy Information Administration

(EIA). U.S. Energy Information Administration, 14 Apr. 2016. Web. 19 Apr. 2016.

Page 109: Acetone Retrofit_Senior Design

108

30. US EIA. "Frequently Asked Questions." U.S. Energy Information Administration.

U.S. Energy Information Administration, 18 Nov. 2015. Web. 19 Apr. 2016.

31. World Bank. "Crude Oil Spot Prices." (n.d.): n. pag. World Bank Commodities Price

Forecast. World Bank, 20 July 2015. Web.

32. ICIS Chemical Business. “Chemical Profile Bisphenol-A.” ICIS. ICIS, 08 Dec.

2015. Web. 15 Apr. 2016.

Page 110: Acetone Retrofit_Senior Design

Section Fourteen: Appendix

Page 111: Acetone Retrofit_Senior Design

Equipment Sizing Calculation Methodologies

Distillation Columns

Methodology

1) Input tray parameters from HYSYS including temperature, pressure, vapor and liquid

flow rates, vapor and liquid densities, key light and heavy K values, liquid viscosity

2) Determine vapor and liquid design flow rates

-Determine which stage has highest total flow

-Input vapor and liquid flows and densities of corresponding stage

-Convert liquid volume flow to liquid mass flow using density

-Calculate Vapor load using following equation:

-Calculate liquid GPM (flow in gallons per minute) using following equation:

3) Determine necessary correction factors

-Input tray spacing (assumed 24 inches), corresponding tray spacing factor, non-foaming

system, corresponding system factor, flood value (assume 70)

-Calculate Flooding Factor (80/Flood)

Calculate corrected vapor and liquid loads using following equation:

Determine tower diameter using Glitsch Method:

Page 112: Acetone Retrofit_Senior Design

111

- Draw line connecting corrected vapor load (left y axis) to corrected liquid load (right

y axis)

- Line intersects single pass tray and double pass tray lines at value of diameter

- To determine diameter for four pass tray, divide corrected vapor and liquid loads by

2 and repeat Glitcsh Method.

- Determine where line intersects double pass tray and multiply this diameter by 21/2

4) Determine actual stages and length

-Calculate alpha (sum light K's/sum heavy K's)

-Calculate average viscosity

-Calculate efficiency using following equation:

-Input theoretical number of stages from HYSYS

-Actual number of trays = theoretical/efficiency

-Input number of feed locations (usually 1)

-Input liquid level time (assume 15 min)

-Input bottoms product rate and density from HYSYS

-Calculate height of liquid level (Volume = Volumetric flow * Time)

-Use volume to get height

Page 113: Acetone Retrofit_Senior Design

112

List of Heights

-Assume reflux = 3 ft

-Height of trays and feed = (# of trays * tray spacing) + (# of feeds * (2*tray spacing))

-Height of reboiler = 3 + (2*tray spacing)

-Height of liquid level = calculated value If value is less than 3 ft, use 3ft (minimum

value)

-Assume tower bottom = 0.5 ft

-Height of manyways: 3 ft per manway, must have one at top and bottom and 1 for every

20 trays in tower TOTAL HEIGHT = SUM OF ALL HEIGHTS CALCULATED

5) Test width vs height. If L/D > 30, switch to two diameter tower (never implemented in

this design)

Assumptions

Most towers did not require liquid level heights of >3ft. However, 3 ft was added to the

bottom of each column for liquid holdup since this was defined as the minimum liquid level

height.

Towers which featured high vapor and liquid flows often did not allow for the Glitsch

method to be applied for single-pass and two-pass columns. The high flows only allowed for

the calculation of the diameter of a four-pass column in these situations. Therefore, it was

assumed that towers with high flows could not have single-pass or two-pass trays.

Reflux Drums

Methodology

The reflux drums were designed to be able to hold up to 15 minutes of flow in case a

column following the reflux drum goes down and flow gets backed up in the system. The

following details the calculations to determine the volume of these drums:

1) Obtain condenser volumetric flow rate from HYSYS (in ft3/s).

2) Multiply the flow by the residence time you wish to achieve (15 min) in order to obtain

the volume.

3) Assume an L/D of 3, and use V= (Pi/4)D4 L to obtain the length and diameter of the

reflux drum.

Page 114: Acetone Retrofit_Senior Design

113

Assumptions:

The reflux drums are assumed to be horizontal, so we set the L/D to the minimum of 3 to

prevent sagging of the drum due to high static head pressure.

Heat Exchangers

Methodology

Desired stream temperatures were obtained from the process model and used to determine

the utility type, area, and flow needed for a heat exchanger. Excel's solver function was

used to minimize the utility cost when more than one utility was in the correct temperature

range.

Calculations:

Two equations are used to solve for any combination of two unknown variables, allowing

the area to be solved for:

The values for U provided on p.5-31 of the Blue Book were used to account for fouling and

estimate the overall heat transfer unit. An additional safety factor of 1.1x the calculated area

was used.

The F factor for counter current heat exchangers was used, and is summarized below:

Page 115: Acetone Retrofit_Senior Design

114

Assumptions

The heat exchangers were assumed to be perfectly insulated from the environment such that

no heat is lost from the shells. The temperature approach of 20oF was used for temperatures

below between 0-300oF, and 100oF for temperatures above 600oF. A correction factor “F”

was used to correct the log mean temperature difference to account for imperfectly

countercurrent, multi-tube heat exchangers. The U values provided in the Blue Book were

used to take into account fouling and nucleate/film boiling.

Page 116: Acetone Retrofit_Senior Design

115

Shell and Tube Reactor

Methodology

Sizing Calculation Defined: inlet pressure, pressure drop, maximum number of tubes,

maximum tube length, reactor diameter, and tube diameter.

1) A tube length was chosen, and the number of tubes needed to hold the volume of

catalyst was calculated assuming the tubes were cylinders. The pressure drop across

each tube was calculated using the single phase Ergun Equation because it is a vapor

phase reaction:

2) Once a pressure drop was found, the length of tube was updated, and the iteration

continued until the pressure drop equaled 20 psi. The largest pressure drop led to the

smallest number of tubes, which reduces capital cost. The reactor diameter was estimated by

assuming the tubes were square packed into a square shape: where N is the number of tubes.

Once the number of tubes, pressure drop, and reactor diameter were within the defined

limits, the reactor was sized.

Heat Calculations: The area needed for heat transfer was calculated using an overall heat

transfer coefficient: The log mean temperature difference had a defined limit of 90 F, Q was

determined by the heat of reaction and conversion (taken from HYSYS), and U was taken

from reference 16.

Assumptions:

Dowtherm oil is used as the heat transfer fluid, and it has an overall heat transfer unit of 15

BTU/(hr ft2 oF) [17] when used in a shell and tube reactor. We assumed that there was only

a single vapor phase flowing through the tubes, making the single phase Ergun Equation

valid. The Ergun Equation assumes a uniform packing density, no wall effects, that the

Page 117: Acetone Retrofit_Senior Design

116

catalyst shape/size do not change over time, and the flow patterns are uniform throughout

the tube. It was also assumed that a single shell reactor could be built with nearly 7,000

tubes. The area for heat transfer was calculated assuming there is space between every tube

so the entire surface area of all of the tubes comprises the area for heat transfer. It was also

assumed to have a perfectly insulated shell that does not lose heat to the surroundings. The

temperature change across the reactor is estimated as a countercurrent exchanger, where the

log mean temperature difference is a good approximation across the length.

Hot Oil System

Methodology

1) The amount of duty needed from the system was determined, and a thermal

efficiency was factored in to determine the absolute amount of duty to provide to the

system:

Actual Duty = Theoretical Duty/n, where n=efficiency

2) The mass flow rate of hot oil was calculated from a known Q, Cp, and ᐃT using the

following equation, assuming no vaporization of the oil:

3) The mass flow rate, volume of the reactor, assuming no tubes and a safety factor of

1.5 (equation below), plus the volume in the length of the pipes was used to

determine the volume of hot oil in the system. An exchange of two volumes per

hour was used, giving a final volume of oil in the system.

4) A surge tank was designed to hold all of the oil in the system, and used an L/D of 3.

Assumptions

No heat is lost to the environment as the oil is piped through the system. The system piping

is 10m long, and a value of 129.06L oil/m pipe was used to determine the volume in the

pipes. The pressure drop across the system was assumed to be 5 psi across the fire heater, 5

psi through the surge tank, 10 psi through the purge valve, 20 psi across the reactor, and 5

psi through the piping because of the oil’s high viscosity. A safety factor of 1.5 was used

when calculating the surge tank volume to account for oil expansion with temperature.

Page 118: Acetone Retrofit_Senior Design

117

Pumps

Methodology

1) Flow rate obtained from bottoms stream for the bottoms pump, the distillate stream for

the distillate pump, and from the reflux rate for the reflux pump. The flow was multiplied by

the stream density to calculate the flow in GPM. The design flow that we used was 1.25

times this value (for safety considerations).

2) Determining Psctn: The source pressure was obtained from the reboiler or condenser in

HYSYS. The static head on the suction side of the pump was calculated based on the

reboiler being at the skirt height of 10 ft, and the condenser being at the top of the column

(used total column height + skirt height). A 1 psi drop was assumed for the suction line.

3) Determine Pdsch: First we determined the destination pressure from the feed tray of the

destination column for the bottoms and distillate, or the first tray in the column for the

reflux pump. The pressure drops due to static head, the heat exchanger, the flow meter, and

the control valve were added on to the destination pressure to obtain Pdsch. Pressure drop

across HEX was assumed to be 5 psi, across flow meters was assumed to be 3 psi, and

across control valves was assumed to be 10 psi. The discharge line pressure was assumed to

be 10 psi.

4) The head on the pump in feet was calculated as Pdsch-Psctn.

5) This head was used along with the flow rate in GPM to determine the NPSH using the

following plot. If the required NPSH was greater than Psctn, there would be cavitation in the

pump, so we made sure all of our suction pressures were above the NPSH.

Page 119: Acetone Retrofit_Senior Design

118

6) The pump efficiency was then obtained from the following plot using the flow rate of the

stream being pumped (in GPM).

Page 120: Acetone Retrofit_Senior Design

119

6) The brake horsepower was obtained via the following equation, where w is the mass flow

rate in lb/min and head is the pressure head in feet:

7) A motor efficiency of 90% was assumed. The motor horsepower was determined using

the following equation:

Assumptions

Reboilers serve as source for bottoms pumps

Reflux drums serve as source for distillate and reflux pumps

Page 121: Acetone Retrofit_Senior Design

120

Compressors

Methodology

1) Input incoming fluid properties including molecular weight, compressibility factor,

K ratio (Cp/Cv) from HYSYS

2) Input polytropic efficiency of compressor from HYSYS

3) Input inlet and outlet temperatures from HYSYS

4) Calculate discharge temperature (T2) actual using following equation:

5) Compare T2 calculated to T2 from HYSYS. If T2 calculated >> T2 HYSYS, design

multiple compressors in series with intermittent cooling.

6) Calculate polytropic head using following equation:

6) Calculate power per stage using following equation:

Assumptions

Motor efficiency= 90%.

Page 122: Acetone Retrofit_Senior Design

121

Holdup Tanks

Methodology

1) Decide where to put holding tank in the process by assuming malfunction of a unit

operations, type of material being transported and the effectiveness of putting a hold up tank

in a specific location.

2) Sizing Calculations

a. Obtain volumetric flow rate of the stream.

b. Decide on a hold up time specified by industry guidelines or necessary operation

holding time.

c. Multiply holding time and volumetric flow rate to get tank of volume required.

d. Apply assumption of oversizing tank based on given volume (i.e. assuming all tanks are

90% full)

e. Decide on what dimensional size you want the tanks to be (i.e. tall and skinny tanks or

fat and short tanks)

Volumetric flow rate * hold up time = Tank of Volume

3.1415 * radius2

* length = Volume

3) To conserve space, tall and skinny tanks were implemented except for the 14 day hold up

time

4) Floating roof design was used to prevent vapor loss or otherwise fixed roof design was

implemented

5) API Standard 650 standard tanks were chosen and API 620 for very low pressure

6) Material considerations were used by checking the corrosiveness and reactivity of

materials in the stream with the multiple materials used for construction; nickel came out to

be the cheapest and most resistant.

Assumptions

Design Recommendations from the Blue Book:

Product: 14 days

Raw material: 7 days

Chemical for process: 7 days

Intermediate Tankage: 7 days

Intermediate Holdup Time 8 hours

The intermediate holdup time value was used since it is more practical and similar to

industry. We assumed the tanks would be filled to 90 % capacity.

Page 123: Acetone Retrofit_Senior Design

122

Carbon Beds

Methodology

The following method was used to calculate the height, diameter, and volume of the two

carbon beds in the system:

1) We are given that we need 1 lb of carbon for every lb/hr of acetone product flow.

2) We divided the required mass of carbon by the density of the carbon (55 g/ft3) to get the

volume of the carbon section of the bed.

3) We used an L/D of 4 to size the carbon section of the bed. This gave us a diameter and

length for the carbon section of the bed since volume was known.

4) 6 feet at the bottom of the column and 6 feet at the top of the column were added as

manways for maintenance. These heights were added to the height of the carbon section to

obtain the total height.

Assumptions

The carbon adsorbs the same over time - assumes no loss of activity. An L/D of 4 was

assumed because it is the standard ratio used for columns.

Deciding Where to Place Holding Tanks

Before the reactor

The reactor has a large flow and if the reactor goes down, no fluid should be able to pass

because the reactants will not be converted to products, and the product downstream will

therefore not meet specifications.

If separators 400 and/or 500 go down, the fluid in the recycle can go to the holdup tank

before the reactor, and fluid coming to the reactor from separator 200 will also be held since

the reactor products cannot be purified.

If the reactor goes down, the holdup before the reactor can hold the volume that was

in that column, and the feed to the reactor will also be held-up since the reactor products

cannot go to the recycle stream until separator 500 is operational again.

Page 124: Acetone Retrofit_Senior Design

123

Before Separator 300

This also has a high flow rate, and a column equipment malfunction would mean a large

amount of fluid that cannot be processed. This is needed so the flow from both of the feeds

does not need to be stopped, and the reaction can continue, and product can continue to be

made. Any liquid after separator 300 will be sent into a mixing point to another holding

tank.

Before Separator 500

So if separator 500 fails, the stream goes on to separator 600 and proportionally a large

amount of the 9000 lb/hr stream going into separator 500 would end up potentially going to

waste. It also can act as a control of the stream going into mixer E and is the major portion

of the bottom recycle that can potentially go awry. And if separator 400 goes down, it can

easily be stored into a holdup tank prior to separator 500 that can be recycled again. From

separator 400, about 300 barrels/day goes up the top and about 771 barrels/day go from the

bottom. So having a tank of about 1000 barrels/day should be safe to hold most of it and

have the excess balanced in the tank right before the reactor (in the recycle).

Before and After Separator 700

This is because if our final separator fails, our acetone purity will drop and we will not be

able to meet the specifications so it is necessary to put a holdup tank before separator 700.

Additionally, there is a need to have a holdup tank after the separator to hold the product

and keep the stream in case there is a need to put it back in the process if there was a prior

failure.

Material and Type of Holding Tank Consideration

The material of the tanks should be made out of nickel since it is the most resistant to

acetone, methanol, propanol, formaldehyde and the other chemical components. It is also

cheaper than monel, tantalum, titanium and the other possibilities for resistant materials for

holding tank materials.

The type of holding tanks were designed based on the pressure of the stream so if they are

less than 2.5 psi, they are put in an API Standard 650 and for greater pressure values, we

used API standard 620. For streams with small proportions of vapors, we used a floating

rooftop tank to prevent less vapor loss while for 100% liquid streams, we used the fixed

rooftop tank.

Page 125: Acetone Retrofit_Senior Design

124

Equipment Specification Sheets

Distillation Column

Page 126: Acetone Retrofit_Senior Design

125

Page 127: Acetone Retrofit_Senior Design

126

Heat Exchanger

Page 128: Acetone Retrofit_Senior Design

127

Page 129: Acetone Retrofit_Senior Design

128

Pumps

Page 130: Acetone Retrofit_Senior Design

129

Page 131: Acetone Retrofit_Senior Design

130

Page 132: Acetone Retrofit_Senior Design

131

Compressor

Page 133: Acetone Retrofit_Senior Design

132

Page 134: Acetone Retrofit_Senior Design

133

Equipment Sizing Calculations by Unit Operation

Tower 100

Page 135: Acetone Retrofit_Senior Design

134

Page 136: Acetone Retrofit_Senior Design

135

Page 137: Acetone Retrofit_Senior Design

136

Page 138: Acetone Retrofit_Senior Design

137

Tower 200

Page 139: Acetone Retrofit_Senior Design

138

Page 140: Acetone Retrofit_Senior Design

139

Page 141: Acetone Retrofit_Senior Design

140

Page 142: Acetone Retrofit_Senior Design

141

Tower 300

Page 143: Acetone Retrofit_Senior Design

142

Page 144: Acetone Retrofit_Senior Design

143

Page 145: Acetone Retrofit_Senior Design

144

Page 146: Acetone Retrofit_Senior Design

145

Page 147: Acetone Retrofit_Senior Design

146

Tower 400

Page 148: Acetone Retrofit_Senior Design

147

Page 149: Acetone Retrofit_Senior Design

148

Page 150: Acetone Retrofit_Senior Design

149

Page 151: Acetone Retrofit_Senior Design

150

Page 152: Acetone Retrofit_Senior Design

151

Tower 500

Page 153: Acetone Retrofit_Senior Design

152

Page 154: Acetone Retrofit_Senior Design

153

Page 155: Acetone Retrofit_Senior Design

154

Page 156: Acetone Retrofit_Senior Design

155

Page 157: Acetone Retrofit_Senior Design

156

Tower 600

Page 158: Acetone Retrofit_Senior Design

157

Page 159: Acetone Retrofit_Senior Design

158

Page 160: Acetone Retrofit_Senior Design

159

Page 161: Acetone Retrofit_Senior Design

160

Tower 700

Page 162: Acetone Retrofit_Senior Design

161

Page 163: Acetone Retrofit_Senior Design

162

Page 164: Acetone Retrofit_Senior Design

163

Page 165: Acetone Retrofit_Senior Design

164

Page 166: Acetone Retrofit_Senior Design

165

Reactor

Page 167: Acetone Retrofit_Senior Design

166

Page 168: Acetone Retrofit_Senior Design

167

Hot Oil System

Page 169: Acetone Retrofit_Senior Design

168

Economic Calculation Methodologies (ICARUS Inputs):

Assumptions

General Process Data

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)

Page 170: Acetone Retrofit_Senior Design

169

Project Area

The footprint area was left as 50x50 ft, except for Tower 700 where the holdup tank had a

40 ft diameter. The footprint was increased to 65x50 ft.

**All equipment was specified as 1 identical unless otherwise stated**

Distillation Columns

Assume all columns have valve trays

Shell material: Tower 100 - stainless steel (ss304)

All other towers- carbon steel (A515)

Tray material: ss410 for all (avoids tray corrosion, trays are thin)

Kept corrosion allowance to ⅛” because we have a corrosive material but we compensate by

changing to stainless steel

Design gauge pressure was set to 30 psig if our calculated value was below that. Vacuum

design gauge pressure was set to -14.7 psig.

Pdesign psig = Poperation psia + 30 psia - 14.7 psia

Reflux Drums

Specified to withstand full vacuum (-14.7 psig) and the pressure seen at the condenser.

Condensers

All assumed to be fixed tube-sheet heat exchangers

Reboilers

All assumed to be thermosiphon heat exchangers

Pumps

Each pump was duplicated (2 identical items).

All pumps have double mechanical seals

All assumed to be API610-centrifugal pumps or ANSI if GPM<300 or head<500ft

Page 171: Acetone Retrofit_Senior Design

170

Pumps for tower 100 use ss304 on all parts to protect against corrosion. ss316 is not

necessary because it is designed for chloride resistance, which is not in our system.

Reflux and distillate pumps were designed to the temperature of the condenser. Bottoms

pumps were designed for the temperature of the reboiler.

Heat Exchangers

Fixed tube exchangers (sizings were based on delta T log-mean difference and a safety

factor of 1.1 times calculated area)

The tube side of the heat exchanger feeding Tower 100 was specified to be ss321

($157,800) because it offered the corrosion resistant properties that are needed for heating

the stream containing acetic acid, but was cheaper than ss347 ($158,200), which offered

similar corrosion resistant properties. ss321 was also specified for the reboiler tubes because

the reboiler is in contact with the high acetic acid content of the bottoms product.

Pdesign was calculated assuming the pressure into the heat exchanger is the same as the

operating Pdsch from the pump before it, if applicable. This means that the control valve in

the line is set to keep the pressure specified to Pdsch, not another pressure.

For reboiler pumps, the designed pressure and temperatures were calculated using the

pressure and temperature at the condenser. The same method was used for the reboiler.

Carbon steel was chosen to be the tube and shell material for all heat exchangers. One

exception was the tube material for the heat exchangers before and part of Tower 100

because acetic acid is corrosive.

The pressure of refrigerant systems was assumed to be 50 psi, and the pressure of LLPS was

assumed to be 50 psi instead of 25 psi.

Vacuum Pump

Using oil-sealed vacuum pump because water-sealed requires an extra cooler.

Reactor

Number of identical items - 2

Modeled as heat exchanger

Holdup Tanks

Vertical vessels

Page 172: Acetone Retrofit_Senior Design

171

MOC - Nickel

Skirt height is 10 ft

Compressors

Centrifugal

Tdesign =200oF because this is the minimum allowed in ICARUS. Companies build

compressors for T>500oF [21]

Carbon Beds

2 identical items

Sizing Inputs

Tower 100

Page 173: Acetone Retrofit_Senior Design

172

Tower 200

Tower 300

Page 174: Acetone Retrofit_Senior Design

173

Tower 400

Page 175: Acetone Retrofit_Senior Design

174

Tower 500

Tower 600

Page 176: Acetone Retrofit_Senior Design

175

Tower 700

IPOH Reactor

Page 177: Acetone Retrofit_Senior Design

176

Carbon Bed

Recycle Pumps

Alternate Cases:

PPE/H2 Separator

Page 178: Acetone Retrofit_Senior Design

177

Extractive Distillation

ICARUS Individual Equipment Prices

Base Case

Page 179: Acetone Retrofit_Senior Design

178

Page 180: Acetone Retrofit_Senior Design

179

Price Correlation Curves

Page 181: Acetone Retrofit_Senior Design

180

Page 182: Acetone Retrofit_Senior Design

181

Alternative Case Capital and Cash Flow Sheets

Extractive Distillation

Page 183: Acetone Retrofit_Senior Design

182

Page 184: Acetone Retrofit_Senior Design

183

Hydrogen-Propylene Separator

Page 185: Acetone Retrofit_Senior Design

184

Page 186: Acetone Retrofit_Senior Design

185

Sensitivity Analysis:

Case 1: 30MM lb Surplus of Acetone

Option 1: Add holding tank to store surplus, sell eventually

Page 187: Acetone Retrofit_Senior Design

186

Option 2: Use surplus for fuel to reduce utility costs

Page 188: Acetone Retrofit_Senior Design

187

Case 2: Acetone Price Changes

Page 189: Acetone Retrofit_Senior Design

188

Acetone Price vs. ATROR

Case 3: Capital Cost changes

Capital Cost vs. ATROR

Page 190: Acetone Retrofit_Senior Design

189

Page 191: Acetone Retrofit_Senior Design

190

Case 4: Natural Gas Price Changes

Linear Interpolation of Utility Prices with Natural Gas Prices

Natural gas cost vs. ATROR

Page 192: Acetone Retrofit_Senior Design

HYSYS Model