Optimization of the Acetone Plant Using CHEMCAD

Chung, Hodgson Hodgson, Roundy Points Points Possible earned 2 2 1 1 5 2 2 2 10 10 15 15 70 70 15 15 15 15 5 5 5 5 5 5 150 147 0 150 147

Grades for Project 2‐Optimization of an Existing Chemical Process Facility

Report Section Letter of Transmittal Title Page Executive Summary Table of Contents Introduction Methods Process Design Results Discussion Conclusion References Appendix‐Calculations Gross Total Adjustments for Report Appearance, Spelling, Grammar, etc. Net Total

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i

Optimization of the

Acetone Production Facility

Texas Tech University Ch E 4555 February 26, 2007 Yongchul Chung Cade Hodgson Sam Roundy


ii

Executive Summary The objective of this project is to design an improvement to an existing chemical processing facility which produces acetone by the dehydrogenation of isopropyl alcohol. It is recommended that a series of three compressors and two heat exchangers be used to purify a hydrogen product stream to industrial grade and recover more acetone. This design was optimized using a CHEMCAD simulation to determine the conditions necessary to provide adequate separation. The incremental revenue from the extra acetone and improved hydrogen is considered against the capital cost of equipment and operating costs for that equipment. Capital costs were obtained from CHEMCAD and CAPCOST while utility costs and revenues were calculated manually. The resulting cash flows and rates of return are presented in the table below. The proposed expansion is extremely profitable, and the project team strongly recommends that this proposal be implemented as soon as practical. Table 1: Economic summary Capital Investment Required Incremental Annual Revenue Incremental Annual Operating Cost (COMd) Average Annual Net Profit Before Tax Rate of Return After Tax Rate of Return After Tax Rate of Return on Net Profit

$1,460,100 $8,814,200 $2,805,500 $3,053,800 159% 123% 117%


iii

Table of Contents Introduction ………..…………………………………………………………………….. 1 Methods ……………………………………………………………………………..…… 2 Process Design ………...………………………………………………………………… 4 Results …………………………………………………………………………………… 8 Discussion ……………………………………………………………………………… 10 Conclusion ...…………………………………………………………………………… 12 References ……………………………………………………………………………… 13 Appendix ………………………………………………...………………………………14

List of Tables Table 1 ……………………………………………………………………….………….. ii Table 2 ……………………………………………………………………………….….. 5 Table 3 ……………………………………………………………………….………….. 6 Table 4 ……………………………………………………………….………………….. 6 Table 5 …………………………………………………………….…………………….. 8

List of Figures Figure 1 ………………………………………………………………………..………… 4 Figure 2 …..……………………………………………………………………………… 5 Figure 3 ………………………………………………………….………………………. 8 Figure 4 ……………………………………………………….…………………………. 9


1

Introduction The purpose of the project is to improve the profitability of the acetone production facility given in Turton et al. [6]. This facility makes acetone via the dehydrogenation of isopropyl alcohol (IPA). A complete description and economic analysis of the existing facility is available in a previous report [3]. This optimization is performed as if the current facility is already built and operating, so none of the economics of the base condition are relevant. The four main costs of the facility are raw materials, utilities, waste treatment, and operating labor. Unless significant equipment changes are made, little or no savings can be made from operating labor. Waste treatment accounts for an insignificant amount of the total cost, so it was also discarded as an optimization area. Utilities are a larger cost than waste treatment or operating labor, but utilities can only be optimized to a certain extent, saving only a small percentage. Raw material costs are by far the largest operating cost. The current process achieves a 99.9% overall conversion of IPA to acetone, which leaves no room for significant improvement. However, some of the acetone product is lost in the hydrogen byproduct. The dehydrogenation process produces hydrogen which is separated out and sold as fuel gas because it contains significant amounts of acetone. By improving the separation, the process could recover additional acetone and thereby increase the revenue from the main product. At the same time the hydrogen could be purified to industrial grade, which would greatly increase its value. CHEMCAD was used to simulate the compressors, exchangers, and flash vessel for the proposed expansion, and an optimum design was identified.


2

Methods In this project, it quickly became obvious that at the current pressure of the hydrogen stream (stream 7 in Figure 1), cooling alone would not separate out the acetone. For this reason, compressors were added to raise the pressure before cooling. Simulating the current process in CHEMCAD suggested that for a temperature of -15 °C, the stream would need to be compressed to somewhere between 50 and 100 bar. Using this information, the plant was designed more rigorously in the CHEMCAD simulation according to the following procedures, assumptions, and guidelines. Compressor Design Compressors are applied in series to achieve the desired final pressure. The compression ratio for each compressor is limited to 4:1 based on heuristics for single stage compression [6]. Theoretical power for each compressor is calculated by CHEMCAD based on the outlet pressures specified. The efficiency for all compressors was assumed to be 80% [6]. CHEMCAD provided the installed cost for each compressor based on compressor size and type. All compressors are reciprocating compressors because CHEMCAD does not allow the user to select a rotary compressor which is probably more reasonable for the compressor sizes considered. All other cost factors in CHEMCAD were left at default settings. Heat Exchanger Design Heat exchangers cool the compressed gas stream. Cooling water is available at 30 oC and low temperature refrigerant is available at -20 oC. Because the plant already uses refrigerated water, it was assumed that low temperature refrigerant would be available. Very low temperature refrigerant (-50 oC) was not considered because the existing plant utilities probably cannot supply it. The inlet temperature to the first heat exchanger is calculated by CHEMCAD based on adiabatic compressor operation. The lowest temperature achievable with cooling water is 41 oC if an 11 oC approach is maintained with counter-current flow. The refrigerant can cool down to 15 oC while maintaining a 5 oC approach. These approach values were specified based on heat exchanger heuristics [6]. CHEMCAD calculates duty from the specified outlet temperature for each heat exchanger. The heat transfer area of each exchanger is generated manually from the log mean temperature difference, CHEMCAD’s calculated duty, and a heat transfer coefficient of 150 Btu/hr ft2 oF based on heuristics for condensers [6]. CHEMCAD generated installed costs for the heat exchangers based on the heat transfer area. For cost estimating purposes, the heat exchanger using cooling water was designated a shell and tube heat exchanger with U tubes, and the heat exchanger using refrigerant was designated a refrigeration exchanger. Costs were compared for installing two heat exchangers versus one heat exchanger at each pressure. Costs Two types of costs are associated with this project: utility and capital costs. The utility for the compressors is electricity, and the power consumption is obtained directly from CHEMCAD. The price of electricity is obtained from the Energy Information Administration website [1]. The utilities for the heat exchangers are cooling water and refrigerant. The costs for heat exchanger utilities were obtained from Turton [6] then indexed with the electricity price from the Energy Information Administration website [1] because most costs for cooling water are based on


3 electricity (compression and pumping). For all utilities, price is multiplied by hourly usage to calculate hourly cost. The annual cost is calculated from the hourly cost based on a stream factor of 0.9. As discussed above, the capital costs of each heat exchanger and compressor is obtained from CHEMCAD based on current values of the CEPCI. Capital cost of the vessel is obtained from CAPCOST using the same dimensions as vessel V-402 in the existing process [3]. Revenues The revenue of this project comes from two sources: recovering more acetone into the saleable product and purifying the hydrogen stream to increase its value. It is assumed that currently, the hydrogen can only be used as fuel gas. A price of $0.208/100 scf was applied to this hydrogen based on the value of the fuel gas [1] with a heating value equivalent to the hydrogen stream [1],[4]. The purified hydrogen is worth $3.50/100 scf [2]. The value of acetone is $1.46/kg [3]. CHEMCAD was used to determine the flow rates and purities resulting from the proposed process improvements. Flow rates were multiplied by unit prices to calculate hourly revenue. Annual revenue is calculated from hourly revenue with a stream factor of 0.9. Optimization Method The proposed addition was simulated in CHEMCAD in order to find the optimal design. The goal of the optimization performed was to minimize the objective function. The objective function was defined as the present cost of the capital expenditure and annual utility costs. It was assumed that the revenue produced would be constant as would the cost of the single vessel. Present values of the utility costs are based on a 10 years life at 15% interest with an additional year for construction. The objective function is given in qualitative form in Equation 1. Capital Cost Present Value Present Value    Capital Cost                Objective   of Heat Exchangers of Compressor s of Compressor Utility of Exchanger Utility        

(1)

CHEMCAD was treated as a black box function, and the optimization was performed according to the following procedure: 1. Select an outlet pressure for each compressor. 2. Vary the outlet temperature from the heat exchangers to achieve 99.95% purity in the final hydrogen product. 3. Evaluate the objective function. 4. Go back to step 1 until an optimum is identified. The specifications supplied for heat exchangers and compressors in step 1 and 2 were subject to the design limitations discussed above. Cash Flows This project is considered an addition to an existing process. No economics or cash flows related to the current operations are considered, only incremental investments, costs, and revenues. The tax rate is assumed to be 35%, and capital was depreciated using MACRS with a 9.5 year class life and a 6 year recovery period. Since this project involves only installation of a few pieces of equipment, it is assumed that construction can be completed in one year. The project life is taken as 10 years with no working capital and no salvage value.


4

Process Design Current Process The current configuration comes from Turton et al. [6] and is presented in Figure 1. It is discussed in detail in the previously submitted report [3]. An azeotropic mixture of IPA and water is supplied to the process and combined with the recycle from the final separation column. This mixture is vaporized then reacted in packed tubes. The reactor is heated by molten salt which passes through a fired furnace. The exiting product mixture is cooled then refrigerated.

Figure 1: Process as currently installed. Reflux pumps, reboiler pumps, and accumulators are not shown on the distillation columns. Data boxes show temperature in °C and pressure in bar. This refrigerated product mixture is flashed in a vessel. The overhead product passes through a packed column which strips most of the acetone and IPA out of the gas with water, leaving a 90% hydrogen stream. This hydrogen stream exits as a product stream. The bottom product of the stripping column joins the liquid from the flash vessel and enters a second column for distillation. The overhead product of this second column is condensed and leaves as pharmaceutical grade liquid acetone. The bottom product of the second column is mostly IPA and water. This mixture is sent to a third and final column where it is distilled again. The overhead product of the third column is an azeotropic mixture of IPA and water which is condensed and recycled back to the beginning of the process. The bottoms product of the third column is waste water potentially containing trace amounts of organic solvents. Improved Process The new design proposes keeping all streams and equipment in the current process with the addition of three compressors and two heat exchangers. The added equipment purifies the hydrogen stream to 99.95% while recovering the acetone and water in this stream to be separated by the second and third columns in the separation train. Figure 2 shows the proposed configuration with the additions outlined in red. The hydrogen stream passes through three compressors (C-401, C-402, C-403) in series to achieve a total compression ratio of 47:1. Compressors C-401 and C-402 each have compression ratios of 4:1, and C-403 has a compression ratio of 2.9:1 so that the largest compression ratio occurs at the lowest pressures. This is less expensive than three equally-sized compressors. All three compressors are


5 reciprocating compressors with efficiencies of 80% as per heuristics [6]. In order to maintain a stream factor of 0.9, a rotating spare compressor is included in the proposal. This spare is sized to meet the demands of the largest compressor (C-403), so it can also do the work of either of the smaller compressors. All the equipment in this expansion is made of carbon steel for consistency with the existing equipment. Table 2 summarizes the specifications for the new compressors and vessel V-405.

Figure 2: Proposed process configuration. The added equipment is outlined in a bold red rectangle. All other process equipment and configuration is the same. Data boxes show pressure in bar and temperature in °C. Table 2: Equipment specifications for new compressors and vessel. C-401 C-402 C-403 A/B Equipment Carbon Steel Carbon Steel Carbon Steel MOC 54.5 84.7 93.3 Fluid Power (kW) 80% 80% 80% Efficiency Reciprocating Reciprocating Reciprocating Type 185 406 650 Temperature (°C) 1.5 6 24 Pressure In (bar) 6 24 70 Pressure Out (bar) ― ― ― Diameter (m) ― ― ― Height (m) ― ― ― Orientation ― ― ― Internals ― ― ― Pressure (bar)

V-405 Carbon Steel ― ― ― ― ― ― 0.75 2.25 Vertical SS Demister 70

In order to achieve the final purity of 99.95% hydrogen, the compressed vapor stream must be cooled to −15 °C. First, a small heat exchanger of 2.0 m2 and a duty of -967 MJ/h cools the vapor to 41 °C using cooling water. This is the lowest temperature attainable with regular cooling water while maintaining a 10 °C approach for counter-current flow. To achieve the final


6 temperature of -15 °C, a second heat exchanger with an area of 4.8 m2 and a duty of -87 MJ/h uses low temperature refrigerant supplied at -20 °C. A 5 °C approach with counter-current flow is maintained as per heuristics for refrigeration [6]. Both these heat exchangers have very small heat transfer areas because this is not a large stream. Table 3 summarizes the specifications for the new exchangers. Table 3: Specifications for new heat exchangers. Equipment E-409 E-410 Duty (MJ/h) -967 -87 2 Area (m ) 2.0 4.8 Shell Side Max Temp (°C) 650 41 Pressure (barg) 70 70 Phase Condensing Vapor Condensing Vapor MOC Carbon Steel Carbon Steel Tube Side Max Temp (°C) 45 35 Pressure (bar) 4 4 Liquid Liquid Phase (cooling water) (refrigerant) MOC Carbon Steel Carbon Steel

The cooled effluent from the second heat exchanger enters a flash vessel identical to vessel V402 in the existing plant. The vapor from this vessel is 70.1 kg/h of 99.95% hydrogen. The liquid effluent of the vessel contains 2.5 kmol/h of acetone and 1.3 kmol/h of water. The acetone and water mixture is throttled down adiabatically and mixed into the feed (stream 9) to tower T402 of the existing separation train. It is assumed that T-402 will require a proportional increase in reboiler duty to separate this small incremental flow (less than a 6% increase in the feed rate), but the column continues to recover all (99.6%) of this additional acetone into the condensed overhead product. The increase in waste water flow will be negligible (approximately 3%). Compared to the existing process, only streams in the separation train are changed, and these changes are small (less than 10%). Table 4 shows stream tables with current operating conditions and projected operating conditions for those streams which will be affected by the expansion.


7 Table 4: Stream conditions for streams in Figures 1 (current) and 2 (new). Only streams which are affected by the proposed expansion are shown. 7 9 11 12 Stream Number Current New Current New Current New Current New Temperature (°C) 33 33 22 19.5 61 61 90 90 Pressure (bar) 1.5 1.5 1.63 1.63 1.5 1.5 1.4 1.4 Vapor Fraction 1.0 1.0 0.0 0.0 0.0 0.0 0.0 0.0 Total Flow 38.60 38.60 74.02 77.83 32.29 34.78 41.73 43.04 (kmol/h) Hydrogen Flow 34.78 34.78 0.00 0.00 ― ― ― ― (kmol/h) Acetone Flow 2.51 2.51 32.43 34.92 32.27 34.78 0.16 0.16 (kmol/h) IPA Flow 0.02 0.02 3.84 3.86 0.02 0.02 3.82 3.84 (kmol/h) Water Flow 1.29 1.29 37.75 39.04 ― ― 37.75 39.04 (kmol/h) Table 4 (continued) Stream Number Temperature (°C) Pressure (bar) Vapor Fraction Total Flow (kmol/h) Hydrogen Flow (kmol/h) Acetone Flow (kmol/h) IPA Flow (kmol/h) Water Flow (kmol/h)

15

16

17

Current 109 1.4 0.0 35.85

New 109 1.4 0.0 37.14

Current NA NA NA NA

New -15 70 0.0 3.81

Current 33 1.5 1.0 38.60

New -15 70 1.0 34.79

―

―

NA

0.00

34.78

34.78

―

―

NA

2.50

2.51

0.01

―

―

NA

0.02

0.02

―

35.85

37.14

NA

1.29

1.29

―


8

Results This addition has an after-tax rate of return of 117% on the net profit. The total capital investment of $1,460,100 and annual cost of manufacturing without depreciation (COMd) of $2,805,500 result in an average revenue increase of $8,814,200 per year. This project will have a payback period of less than 4.5 months after startup. The Appendix has the full spreadsheet calculations from which these values and cash flows were obtained. The following figures and tables summarize the economic results of the proposed expansion. Only incremental investment, costs, and revenue are considered. Table 5: Economic summary of major cash flows and rates of return. Capital Investment Required $1,460,100 Incremental Annual Revenue $8,814,200 Incremental Annual Operating $2,805,500 Cost (COMd) Average Annual Net Profit $3,053,800 Before Tax Rate of Return 159% After Tax Rate of Return 123% After Tax Rate of Return on 117% Net Profit $7 $6

Before Tax After Tax

Cash Flow (Millions)

$5 $4 $3 $2 $1 $0 0

1

2

3

4

5

6

7

8

9

10

11

-$1 -$2

Year

Figure 3: Discrete before and after tax cash flows for the life of the proposed project.


9

$70

Before Tax After Tax

$60

Cash Flow (Millions)

$50

$40

$30

$20

$10

$0 0 -$10

1

2

3

4

5

6

7

8

9

10

11

Year

Figure 4: Cumulative before and after tax cash flows for the life of the proposed project.


10

Discussion Profitability can be improved in two ways: reducing cost and increasing revenue. In this project, the team chose to increase revenue. This decision was made based on an evaluation of the potential profitability of each approach. In this case, the primary motivation was the lack of incentive to reduce operating costs. Of the four major operating costs (raw materials, utilities, waste treatment, and operating labor), raw material costs make up 92% of cost in the existing design [3]. Because of effective separation and recycle, the overall mass balance of the system has 99.9% conversion of IPA to acetone. This leaves no way to reduce raw material costs, so only waste treatment, utilities, and operating labor can be optimized. Of these, waste treatment is insignificant (less than $500/yr), and operating labor is essentially invariant without a major redesign of the entire plant. Reducing utility usage leaves less than $1.4 million on the table for improvement. Because even the best designed plant will obviously require some utilities, only a fraction of this $1.4 million can be saved. By contrast, 2.51 kmol/h of acetone is lost in the hydrogen stream. This acetone is potentially worth $1.7 million per year. Additionally, this acetone contaminates the hydrogen stream so the hydrogen can only be used as fuel gas. Hydrogen for heating purposes is worth 6% of industrial grade hydrogen feedstocks. Preliminary calculations showed that if the hydrogen could be made 99.95% pure and sold as industrial hydrogen, the gas stream would be worth an additional $7.1 million annually. The combined potential for revenue from recovered acetone and purified hydrogen is thus $8.8 million annually. It was believed that this entire potential revenue could be realized, and the subsequent design work accomplishes this. By comparison, any fractional reduction in utility costs is inconsequential. For this reason, no utility reduction projects were explored. In optimizing using the objective function, it became clear that for this project, compression requires much larger capital and operating costs than refrigeration. The optimal design applies the maximum cooling possible with just enough compression to achieve the required separation. The cooling is limited by the temperature of the refrigerant, and this in turn sets the conditions for the optimum design. Future Projects Although no utility reduction projects were explored, several opportunities for optimization were identified. The design team suggests the following for future optimization projects:  Use the reactor effluent to heat the reactor affluent. Exchangers E-401 and E-402 provide similar heat duties to heat then cool the process stream at a total utility cost of $339,000/yr. It is unclear if one of these exchangers could be used or if a new exchanger would be required. This straight-forward change was not pursued because of the small saving potential.  Use the furnace to provide heat to exchanger E-401 or possibly E-403. Direct energy from natural gas is cheaper than steam, and these exchangers have an annual utility cost of $585,000. To avoid using molten salt, it would probably be most economical if the process fluid could pass directly through tubes inside the furnace. This was not pursued because no information was available about the cost of such a change.


11 





Use a cheaper catalyst and/or molten salt. Little information was available about these costs, but both must be replaced annually. If a cheaper material with equivalent functionality could be found, this would be the simplest change possible. Use the gas stream leaving compressor C-403 to heat the reactor affluent. This could save duty from both exchanger E-401 and E-409. This should probably be considered before E-409 is installed, but the gas stream is small and could only supply a fraction of the duty E-401 requires. Recover the waste water and recycle it as process water. The waste water contains only trace process contaminants, and may be clean enough for recycle. This is a simple and obvious change, but the total cost of fresh process water and waste treatment is less than $1,000 annually [3]. The piping and pumping costs probably outweigh the benefits.


12

Conclusions This report proposes an expansion to the existing acetone production facility. This expansion purifies hydrogen and recovers more acetone. This separation is accomplished by compressing and refrigerating the gaseous product stream. The most economic design involves minimizing compression and maximizing refrigeration. Although there are many other opportunities for optimization, the proposal in this report represents the largest, most profitable improvement available. This expansion is strongly recommended because the after tax net profit of $3.05 million per year produces a 117% rate of return on a capital investment of $1.46 million. This project should be undertaken immediately.


13

References 1. Energy Information Administration, US Department of Energy, February 19 (2007): http://eia.doe.gov/ . 2. Forrest, W, “Prices, supply getting more stable,” Purchasing.com, April 6 (2006): http://www.purchasing.com/article/CA6320292.html . 3. Frantz, E et al., Production of Acetone via the Dehydrogenation of Isopropyl Alcohol, February 2007. 4. Hydrogen Analysis Resource Center, US Department of Energy, February 19 (2006): http://hydrogen.pnl.gov/cocoon/morf/hydrogen . 5. Pyle, Walt. “Hydrogen Purification.” Home Power #67, October/November (1998). 6. Turton, R., et al., Analysis, Synthesis, and Design of Chemical Processes, 2nd ed. (Upper Saddle River, New Jersey 2003).


14

Appendix Table A.1: Capital investment Capacity

Heat Exchanger E-409 2 E-410 4.8 Compressor C-401 56.5 C-402 84.7 C-403 93.3 Vessel V-405 0.75 x 2.25 Bare Module Cost

Bare Module Price $11,800 $97,700

CHEMCAD CHEMCAD

$213,800 $275,600 $293,000

CHEMCAD CHEMCAD CHEMCAD

$52,500

CAPCOST

Source

$1,237,400

Total Module Cost

$1,460,100

Table A.2: COMd, annual utility costs Utilities Utility

Utility Duty

Utility Unit

Utility Price

Utility Cost ($/yr)

E-409

Cooling water

967

MJ/h

$0.0004136

$3,150

E-410

Refrigerant

87

MJ/h

$0.00922

$6,320

C-401

Electricity

56.5

$0.059

$26,280

C-402

Electricity

84.7

$0.059

$39,400

C-403

Electricity

93.3

$0.059

$43,400

T-402

Low pressure steam

210

$0.0155

$25,610

Heat Exchanger

Compressor kW (actual power) kW (actual power) kW (actual power)

Tower MJ/h

Total


$144,160

15 Table A.3: COMd, incremental annual operating labor Operating Labor Current Design

Improved Design

0

0

13

15

3.05

3.12

14

14

Pieces of Particulate Handling Equipment Pieces of Fluid Handling Equipment Number of Operators per Shift Number of Operators Total Incremental Labor Cost

$0

Table A.4: COMd, annual incremental waste treatment costs Waste Treatment Extra Waste Water Flow (m3/h) Tertiary Treatment Cost ($/m3)

0.0234 0.0942

Incremental Cost

$17

Table A.5: COMd, incremental raw material costs Raw Materials Additional Raw Materials Incremental Cost

None $0

Table A.6: COMd, summary of total cost of manufacturing without depreciation COMd Raw Materials Utilities Waste Treatment Operating Labor FCI

$0 $144,160 $17 $0 $1,460,100

COMd

$2,805,518


16 Table. A.7: Incremental revenue from more acetone recovery and purified hydrogen Current Improved Flow Incremental Chemical Unit Unit Rate Revenue Value Value Acetone 145 $0 $1.46 $1,669,043 (kg) Hydrogen 27530 $0.208 $3.50 $7,145,171 (scf) Total

$8,814,214

Table A.8: Cash flows Year 0 1 2 3 4 5 6 7 8 9 10 11

Cost $1,460,100 $0 $2,805,518 $2,805,518 $2,805,518 $2,805,518 $2,805,518 $2,805,518 $2,805,518 $2,805,518 $2,805,518 $2,805,518

Revenue

Depreciation

$0 $0 $8,814,000 $8,814,000 $8,814,000 $8,814,000 $8,814,000 $8,814,000 $8,814,000 $8,814,000 $8,814,000 $8,814,000

$0 $0 $292,020 $467,232 $280,339 $168,204 $168,204 $84,102 $0 $0 $0 $0

Tax $0 $0 $2,000,762 $1,939,437 $2,004,850 $2,044,097 $2,044,097 $2,073,533 $2,102,969 $2,102,969 $2,102,969 $2,102,969

Before Tax Discrete -$1,460,100 $0 $6,008,482 $6,008,482 $6,008,482 $6,008,482 $6,008,482 $6,008,482 $6,008,482 $6,008,482 $6,008,482 $6,008,482

Table A.8 (continued) Year 0 1 2 3 4 5 6 7 8 9 10 11

After Tax Net Profit -$1,460,100 $0 $3,715,700 $3,601,812 $3,723,293 $3,796,181 $3,796,181 $3,850,847 $3,905,513 $3,905,513 $3,905,513 $3,905,513

Discrete -$1,460,100 $0 $4,007,720 $4,069,044 $4,003,632 $3,964,384 $3,964,384 $3,934,949 $3,905,513 $3,905,513 $3,905,513 $3,905,513

Cumulative -$1,460,100 -$1,460,100 $2,547,620 $6,616,665 $10,620,296 $14,584,681 $18,549,065 $22,484,014 $26,389,527 $30,295,040 $34,200,554 $38,106,067


Cumulative -$1,460,100 -$1,460,100 $4,548,382 $10,556,864 $16,565,345 $22,573,827 $28,582,309 $34,590,791 $40,599,273 $46,607,755 $52,616,236 $58,624,718

17 Table A.9: Calculation used to find unit value of current hydrogen stream as fuel gas Flow of Heating Value Equivalent Value as Value as Price of Fuel Hydrogen of Hydrogen Amount of Fuel Fuel Gas Fuel Gas Gas ($/100scf) (scfh) (MJ/h) Gas (scfh) ($/h) ($/100scf) 27530 8516 8212 0.696 57.15552 0.20761177 Table A.10: Table used to evaluate the object function. Outlet pressures from the final compressor (green) and outlet temperature from the second exchanger (orange) were specified. All other values are calculated by CHEMCAD or in the spreadsheet based on heuristics. The present value of costs (the result of the objective function) is highlighted in blue. Table A.10a: Compressor outlet pressure of 75 bar with 2 heat exchangers E-1 E-2 C-1 C-2 Temperature 665 -14 Pressure Out (bar) 6 24 (oC) Theoretical Power Duty (MJ/h) -991 -85 45.2 67.73 (kW) TLM (oC) 151 6.00 Efficiency 0.8 0.8 2 Area (m ) 2.1 4.63 Power Used (kW) 56.5 84.6625 Utility Price Utility Price 0.414 5.18 0.059 0.059 ($/GJ) ($/kWh) Utility Cost $3,231 $3,468 Utility Cost (yr-1) $26,281 $39,381 (yr-1) Present Value Present Value of $16,217 $17,407 $131,900 $197,645 of Cost Utility Cost Capital $11,929 $96,495 Capital $213,778 $275,581 Total Cost Table A.10b: Compressor outlet pressure of 100 bar with two heat exchangers E-1 E-2 C-1 C-2 Temperature 729 -9 Pressure Out (bar) 6 24 (oC) Theoretical Power Duty (MJ/h) -1099 -77 45.2 67.73 (kW) TLM (oC) 163 8.25 Efficiency 0.8 0.8 2 Area (m ) 2.2 3.05 Power Used (kW) 56.5 84.6625 Utility Price Utility Price 0.414 5.18 0.059 0.059 ($/GJ) ($/kWh) Utility Cost $3,584 $3,142 Utility Cost (yr-1) $26,281 $39,381 (yr-1) Present Value Present Value of $17,985 $15,769 $131,900 $197,645 of Cost Utility Cost Capital $12,087 $66,715 Capital $213,778 $275,881 Total Cost


C-3 75 80.1 0.8 100.125 0.059 $46,574 $233,743 $306,346 $1,501,042

C-3 100 103.82 0.8 129.775 0.059 $60,366 $302,961 $360,966 $1,595,687

18 Table A.10c: Compressor outlet pressure of 85 bar and two heat exchangers E-1 E-2 C-1 C-2 Temperature 692 -12 Pressure Out (bar) 6 24 (oC) Theoretical Power Duty (MJ/h) -1038 -82 45.2 67.73 (kW) TLM (oC) 156 6.95 Efficiency 0.8 0.8 2 Area (m ) 2.2 3.85 Power Used (kW) 56.5 84.6625 Utility Price Utility Price 0.414 5.18 0.059 0.059 ($/GJ) ($/kWh) Utility Cost $3,385 $3,346 Utility Cost (yr-1) $26,281 $39,381 (yr-1) Present Value Present Value of $16,987 $16,793 $131,900 $197,645 of Cost Utility Cost Capital $12,087 $94,044 Capital $213,778 $275,881 Total Cost

C-3 85 90.23 0.8 112.7875 0.059 $52,464 $263,304 $330,279 $1,552,697

Table A.10d: Compressor outlet pressure of 70 bar with two heat exchangers (optimal) E-1 E-2 C-1 C-2 C-3 Temperature 650 -15 Pressure Out (bar) 6 24 70 (oC) Theoretical Power Duty (MJ/h) -967 -87 45.2 67.73 74.65 (kW) TLM (oC) 148 5.48 Efficiency 0.8 0.8 0.8 2 Area (m ) 2.1 5.18 Power Used (kW) 56.5 84.6625 93.3125 Utility Price Utility Price 0.414 5.18 0.059 0.059 0.059 ($/GJ) ($/kWh) Utility Cost $3,153 $3,550 Utility Cost (yr-1) $26,281 $39,381 $43,405 -1 (yr ) Present Value Present Value of $15,825 $17,817 $131,900 $197,645 $217,839 of Cost Utility Cost Capital $12,087 $97,724 Capital $213,778 $275,881 $293,002 Total Cost


$1,473,498

19 Table A.10e: Compressor outlet pressure of 70 bar with one heat exchanger E-1 E-2 C-1 C-2 Temperature 651 NA Pressure Out (bar) 6 24 (oC) Theoretical Power Duty (MJ/h) -1052 NA 45.2 67.7 (kW) TLM (oC) 127 NA Efficiency 0.8 0.8 2 Area (m ) 2.7 NA Power Used (kW) 56.5 84.625 Utility Price Utility Price 5.180 NA 0.059 0.059 ($/GJ) ($/kWh) Utility Cost $42,963 NA Utility Cost (yr-1) $26,281 $39,364 (yr-1) Present Value Present Value of $215,620 NA $131,900 $197,558 of Cost Utility Cost Capital $494,573 NA Capital $213,778 $275,881 Total Cost


C-3 70 74.7 0.8 93.375 0.059 $43,434 $217,985 $293,002 $2,040,297

Optimization of the Acetone Plant Using CHEMCAD

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