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Ethyl Acetate Design Project

University of California Santa Barbara

Omid Borjian

Executive Summary

The catalytic conversion of ethanol to produce ethyl acetate has shown to be a profitable market. We

designed a plant to optimize the production of ethyl acetate, while minimizing operating and production

costs and undesired side products. We found that the reaction was best suited for a 41m3 plug flow

reactor operated at 285 oC and 1 atm. Feeding 120 MM kg/yr of ethanol into the reactor produced 100

MM kg/yr of ethyl acetate to be sold along with 5 MM kg/yr of hydrogen for a total profit before taxes

of a total of $43.6 MM$/yr. To start the plant a total capitalized investment (TCI) was approximated to

be $28.8 MM with a net present value of the project (NPVproj) to be $102 MM and net present value

percent (NPV%) of 32.7%. The internal rate of return (IRR) was found to be 113.5%. The discussed

numbers are approximations, and flexible approach should be considered when plant production

commences. The plant design accounted for market fluctuations, and the process control was purposely

designed simplistically. They are, however, a good basis to gain an understanding of the plant’s general

function and expectations.

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Goals and Introduction

With an increasing industrial demand for ethyl acetate, many have found successful ways to create a

marketable business for the production and distribution of ethyl acetate. This increasing demand has

also initiated industry to develop commercial processes, such as that by DAVY Process Technology, for

large production of ethyl acetate. The GSI Process Feasibility Group has developed a plant that will be in

direct competition with Davy Process Technology. In this plant, ethyl acetate will be synthesized via the

interaction of ethanol with a catalyst consisting of 94% copper oxide, 5% cobalt oxide, and 1% chromium

oxide. Unfortunately, under these conditions ethanol can react to form ether acetaldehyde or diethyl

ether. Diethyl ether is a side product that is of lesser importance and may not be profitably sold. While

diethyl ether does not need to be disposed of and can be burned, in this initial profitability design and

analysis, heat exchange interactions were not taken into account and any credits able to be obtained

from burning the diethyl ether were not accounted for. In addition to ethyl acetate and diethyl ether,

this system of reactions will also produce hydrogen and water. Through the use of a flash drum, the

hydrogen will be separated from the system and be sold for further profit.

The primary challenge is to create an optimally profitable amount of ethyl acetate, while working

around an azeotropic solution involving the ethyl acetate, ethanol, and water. Since in this system the

selectivity is constant over reactor conversion, a higher conversion was able to be chosen without loss of

selectivity. Once a specific conversion is selected, a separation system is able to be designed and the

equipment and streams of the system are able to be cost and implemented into a cost diagram. To

ensure profitability, an economic analysis will be run on the five most important economic parameters.

Conceptual Design

Various factors were taken into consideration when making design decisions to optimize the plant

profitability. These factors consisted of reactor volume, reactor temperature and pressure, along with

other equipment constrictions. The system was found to be optimized at a reactor conversion of 90%

with a recycle stream to the reactor.

Using Douglas’s Conceptual Design hierarchy (Douglas, 2011), ideal stoichiometric mole balances were

developed to find the flow rates of the inlet, outlet, and ideal recycle streams. Using the kinetic data

provided from the GSI technical data sheet, a graph of reactor volume versus reactor conversion was

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constructed for varying temperatures and pressures (Doherty, 2011). Analysis of the chart provided a

minimal reactor volume, which facilitated the selection of optimal operating conditions. It essential that

a minimal reactor volume is shown for cost analysis shows reactor cost grows exponentially as a

function of reactor volume. The reactor optimally operated at 285 oC and 1 atm.

The reaction was run in a heat exchanger with circulating heating fluid because in order to run the

endothermic reaction isothermally. A shell-and-tube heat exchanger was utilized to combine the

costing of the reactor and heat exchanger. Maintaining the heating fluid at the desired temperature

was the primary factor regarding the reactor operating cost. To approximate the heat produced in the

reactor, and thus cost the reactor, the heat capacities were assumed to be constant with respect to

temperature.

The separation consisted of a split block that separated out the diethyl ether, a flash drum that

separated out the hydrogen, two distillation columns that separated out ethyl acetate, and one

distillation column that removed water in the process, purifying the ethanol recycle stream. The flash

drum was optimized at 1 atm and 255 K, allowing for approximately 100% of the hydrogen to exit the

column in the vapor stream. It was particularly challenging to separate the ethanol, ethyl acetate, and

water because they contain azeotropes that prevent separation of individual species. Each distillation

column was designed at specific pressures and temperatures that avoid azeotropes by analysis of

ternary maps as shown in Appendix B. Flash drum calculations were designed in the attached MATLAB

code (Appendix D) and the three distillation columns were designed in ASPEN as shown in Appendix B

To solve the distillation systems multiple trials were run in ASPEN to determine which design was most

effective. Firstly, processes with conversions of 70%, 80%, and 90% were simulated in ASPEN and

compared. Conversions of 80% and 90% showed to be far more profitable than that of 70%. Reactor

size and cost drastically grow exponentially as conversion surpasses 90%. Thus, we set 90% conversion

as a maximum possible conversion for the reactor. Secondly, purge stream and waste disposal analysis

was run. Since the remaining unwanted products were easily burned and did not require an additional

disposal cost, running the process with three distillation columns and a recycle stream or two distillation

columns and disposing the remaining products were both viable options. As depicted in the process

flow diagram (PFD), the former design proved to be more profitable as shown in the economic section.

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Azeotrope Conceptual Design

Designing a system that separates an azeotropic mixture is particularly challenging because the

conventional methods, Gilliland, Fenske and Underwood equations, were not adequate to calculate the

number of stages, V/F ratio, and the distillation feed stage for a specified recovery. This interaction

between the components makes complete separation impossible unless the mixture is operated at a

different pressure (pressure-swing distillation) or added another component (entrainer) to break the

azeotrope. In our design, the first method was sufficient.

Using ASPEN PLUS, a ternary map of ethanol, water, and acetyl acetate was acquired for the system as

shown in Figure 1 below. Depending on the location of feed compositions, we adjusted the pressure to

obtain the maximum separation distillation boundary. In the first and second column, the high

concentration of ethyl acetate was removed by running the distillation at highest possible pressure, 15

atm. We were able to conceptually extract ethyl acetate with 99.99% purity from bottoms of the

columns. In the third column water was separated out the bottom of the column at atmospheric

pressure. Knowing the behavior of the equilibrium curves and tie lines, we specified reflux ration,

distillate, and bottoms compositions for the column in such way that the rectifying and stripping curves

cross each other simultaneously when the system converges. We used ASPEN PLUS to design and

calculate the total number of stages, feed stage, and V/F ratio to obtain the required separation.

Figure 1 Ternary map of ethanol, water, and ethyl acetate at 12 atm.

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Process Control

In the design of the process control we used the standard feedback controllers used in the flash unit and

distillation units to adjust the pressure, temperature, liquid level, reflux ratio, and stream composition.

Our reactor operates isothermally, which required the temperature controller to adjust the temperature

of the feed stream. Fluctuations may occur as the species are reacting.

The recycle stream from the third distillation column enters a recycle surge tank, which is regulated by

signals from the composition controller on the products stream. A valve on the purge stream interacts

with the recycle surge tank’s liquid level controller, which opens to prevent an over flow in the surge

tank. The pressure inside the flash unit and the first two distillation units are controlled by adjusting a

valve on the top vapor stream. The liquids are driven inside the units via pumps to ensure a steady drive

of flows in and out of the system. Lastly the ethanol feed flow rate is controlled by another flow

controller based on the production rate. Figure 2 shows the process control diagram with every

controller listed in Table A.4 with the corresponding controlled and manipulated variables. A larger view

of Figure 2 is referenced in Appendix B.

Flash

Unit

Distillation

Unit1

Distillation

Unit2

Distillation

Unit3Recycle Surge Tank

Mixer

Split Block

Wastes Stream

Purge

Stream

LC

23

Cooling water

TC

24

Bot3

LC

19

FC

1

E stream, w1

Cooling Water

TC

3

Cooling water

LC

5

PC

4

Steam

TC

2

Reflux3

AC

20

Steam

AC

18

Cooling Water

Condenser3

Top3

LC

17

AC

13

Condenser2

Reflux2

LC

12

Top2

Cooling Water

Condenser1

Reflux1

LC

7

Top1

AC

8

Hydrogen

PC

11

PC

6

AC

22

Reactor

Bot2

LC

15

AC

16

Steam

Bot1

Steam

LC

9

AC

10

EA Stream, w2

Products

W Stream

Waste water

FC

21

TC

13

Cooling water

Figure 2 Process Control Flowsheet

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Economic Design and Analysis

While the basis of most of the plant design decisions were products of the conceptual design, the

reactor conversion and reactor temperature were decided based on the final economic analysis run on

the conceptual design of the system. This analysis was further justified by the reactor volume, pressure,

and temperature relationship.

The economic analysis run on the conceptual design consisted of graphing the net present value of the

project (NPVproj), the net present value at year zero (NPVzero), the risk associated with the project (NPV%),

the return on investment before taxes (ROIbt) based off of the total investment (TI), and the total

capitalized investment (TCI) against the reactor conversion of 80% and 90% for two different situation as

seen in Figures 3 through 4. Other economic figures are located in Appendix C. (While Figures 3

through 4 only show situations at 80% and 90%, it should be noted that an initial analysis was done for

70%, 80%, and 90% which showed 80% and 90% to be the more profitable reactor conversions). The

two different situations analyzed for the system at each reactor conversion revolved around the recycle

stream. One design analyzed the profitability to have less distillation columns and purge the potential

recycle stream, while the other proved that the design of a distillation column with a recycle stream was

more profitable. (The first situation is indicated on the figures by a 0.05 addition to the conversion). The

trends on these figures were then analyzed to find the most profitable reactor temperature and

conversion. The most profitable combination was based off of the two parameters NPVproj and NPV%.

Economically, the most desirable combination would maximize both of these quantities leading to the

highest net plant worth at the time of project conception with the least amount of risk associated with

the project. For this project a risk of less than 15% was not acceptable since ethyl acetate is a

commodity chemical.

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Figure 3. Net present value of the project versus the reactor conversion with the operating point highlighted.

Figure 4. The net present value percent of the project versus reactor conversion with the operating point highlighted.

Since it was found that for all alternatives analyzed the NPV% was much higher than 15%, process design

was chosen by optimizing NPVproj (Mellichamp, 2011). By optimizing NPVproj, this allowed for the plant

worth to be maximized. This optimization was found to be consistent with the temperature of 285 oC,

and the separation system that consisted of three distillation columns and a recycle stream to the

reactor which corresponds to 0.9 on the figures presented previously.

Operating Point:

T = 285 oC and P = 1

atm

Operating Point:

T = 285 oC and P = 1

atm

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Economic Cash Flow Analysis

The conceptual and economic design executed in HYSYS provided a framework to calculate the fixed

capital and the fraction of working capital as well as the predicted profit before taxes utilizing the

conceptual design MATLAB program (Appendix D). Discounted cash flow analysis as well as a sensitivity

and fluctuation analysis was then performed using the previously mentioned results.

The economic model used in this analysis was created on the basis that the finance and construction

interest rates, the fractions of startup capital and salvage value, the amount of fixed capital spent during

the construction years, the fraction of profit before taxes made in the ramp up years, and the operating

and fixed capital costs were all reasonable approximations. The finance interest rate was assigned a

value based on colloquial knowledge that normal finance rates range from three to five percent but can

be as high as ten percent. With this knowledge, an eight percent finance rate was chosen in order for

the calculations of project value to be conservative.

An average construction rate was taken to be approximately 7% based on data and examples from

Evaluating Plant Profitability in a Risk-Return Context (Mellichamp, 2011). Again to keep the plant value

calculations conservative a construction rate of 10% was chosen for the cash flow analyses. This will

reduce the risk of calculating financial data that would indicate an exaggerated profit. The fractions of

startup capital and salvage value were based on the fact that the startup capital would be only a small

portion of the fixed capital and the amount salvaged from sales after decommissioning of the plant

would be even significantly smaller. Fixed capital spending rate in the construction years was chosen

assuming that less of the fixed capital would be spent in the first year, when the final plant designs are

being finalized and plant construction is minimal. In the second year during plant construction, the

majority of the fixed capital is used.

The ramp-up fractions were chosen assuming that in the first few years, profitability will be lower than

expected. This is believed to be true since ethyl acetate is a commodity chemical and market

competition exists, which is expected to cause low initial profit. Finally, the operating and fixed capital

costs were based on the factors that could not be assumed to be negligible. In both of these

calculations offsite costs, or outside battery limit (OSBL) costs, were assumed to be negligible, while the

onsite costs, or inside battery limit (ISBL) costs, were assumed to greatly affect these two economic

calculations. ISBL costs were considered to be the most important costs but not all of these costs were

consider in the model calculations. Among others, the cost of pumps and mixers were assumed to be

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negligible. Ultimately, the fixed capital costs included the costs for the separation system (three

distillation columns), the heat exchanger reactor and multiple other pieces of heating equipment. The

reactor cost was assumed to be negligible in comparison to the heat exchanger and only the heat

exchanger portion of the reactor was cost.

The operating costs included the cost for the separation system, the cost for additional heating and

cooling operations, and the cost for heating the Dowtherm used to keep the reactor isothermal. These

were assumed to be the main costs that would affect the calculation for the operating cost, and costs

such as the plant electricity were assumed to be negligible. Once these values were all chosen

(Appendix C), a second economic analysis was performed to find the sensitivity of the base case to

variations in specific parameters, and to determine relative maximum finance rate, minimum selling

price of ethyl acetate, and the maximum cost of ethanol before the NPVzero is equal to zero. The results

of this analysis are summarized below in Table 2 with variations listed in Table 1.

Table 1. The variations performed on the base case.

Variation Number Alteration Performed

Variation 1.a/b Increase/decrease the cost of the ethanol

Variation 2.a/b Increase/decrease the value of ethyl acetate

Variation 3.a/b Increase/decrease the construction rate.

Variation 4.a/b Decrease/increase the finance rate.

Abnormal 1 New political leadership drastically lowers tax rate.

Abnormal 2 Competitor enters market in year 5 reducing profits.

Sell Price Drop How low can the selling price of ethyl acetate go?

Raw Material Raise How high can the cost of ethanol go?

IRR At what finance rate does the NPVzero equal zero?

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Table 2. A summary of the results of the sensitivity and dependence analysis run on the base case plant

design produced in the HYSYS and ASPEN simulation.

Variation Change Variable

Variable Originally

Variable Changed To

NPV(0) NPVproj Percent Deviation From Base Case

NPV(0) NPVproj

Base Case NA NA NA 118.08 101.24 NA NA

Variation 1.a

Ethanol Price

$0.70/kg $1.00/kg 3.85 3.30 96.74 96.74

Variation 1.b

Ethanol Price

$0.70/kg $0.40/kg 232.32 199.18 96.74 96.74

Variation 2.a

Ethyl Acetate Value

$1.30/kg $1.50/kg 178.66 153.18 51.31 51.31

Variation 2.b

Ethyl Acetate Value

$1.30/kg $1.10/kg 57.50 49.30 51.31 51.31

Variation 3.a

Construction Rate

10% 15% 117.99 101.16 0.08 0.08

Variation 3.b

Construction Rate

10% 5% 118.17 101.32 0.08 0.08

Variation 4.a

Finance Rate

8% 3% 156.30 147.32 32.36 45.53

Variation 4.b

Finance Rate

8% 13% 91.17 71.40 22.79 29.47

Abnormal 1

Tax Rate 50% 35% 155.65 133.45 31.82 31.82

Abnormal 2

b's (fraction of P_bt made)

b1 = 0.5, b2 = 0.8, b3 = 1

Varying Less Than 1

101.40 86.94 14.12 14.12

Sell Price Drop

Ethyl Acetate Value

130.5 MM $/yr

91.4 MM $/yr

0.00 0.00 NA NA

Raw Material Raise

Ethanol Price

85.1 MM $/yr

124.2 MM $/yr

0.00 0.00 NA NA

IRR Finance Rate

8% 106.8% 0.00 0.00 NA NA

This analysis showed that the two parameters that have the most effect on the economics of the plant

are the price of ethanol and the value of the ethyl acetate, causing deviations from the base case of 97%

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and 51% as seen in Variation 1.a through Variations 2.b. Base case changes in the finance rate,

construction rate, and tax rate produced relatively small percent deviations. However, between the

interest rates the one with the largest effect on the plant profitability was found to be the tax and a

large drop in this interest rate was further analyzed.

The final analysis that was done on the economics was a fluctuation analysis which focused on a

potential drop in ethyl acetate worth and rise in ethanol worth to discover how much fluctuation the

project could withstand before the project risk was too great. While the fluctuation analysis does look

slightly at a raise in ethyl acetate worth and drop in ethanol worth, the main focus is on the ethyl

acetate worth dropping and the ethanol worth increasing because it is a primary concern that with the

addition of this plant to the market that price of ethyl acetate will be forced down since it will be in

greater supply. Also, an increase in ethanol worth is analyzed because this process adds additional value

to ethanol and it is likely that the addition of the plant will cause the demand for ethanol to increase.

After the fluctuation analysis was completed, it was obvious that this project is worth further

investigation if it is projected that the value of ethyl acetate will not drop much below $1.00/kg and the

cost of ethanol will not increase much higher than $0.80/kg. This boundary is indicated in the

fluctuation analysis table (Table 3) by the brown coloring between the green and red.

Table 3. Fluctuation analysis of the risk of the project based on the ethanol and ethyl acetate prices.

Table is focused on drop in ethyl acetate since it is a primary concern that the market will not be able to

withstand the current selling price of ethyl acetate once GSI enters the market.

Sell Price

Raw Material MM $/yr 56.1 61.0 66.3 72.0 78.3 85.1 91.9 99.3 107.2 115.8

56.7 -4.2% -7.9% -11.8% -16.1% -20.8% -25.9% -31.0% -36.5% -42.5% -48.9%

61.6 -0.5% -4.2% -8.1% -12.5% -17.1% -22.2% -27.3% -32.8% -38.8% -45.2%

67.0 3.5% -0.2% -4.1% -8.5% -13.1% -18.2% -23.3% -28.8% -34.8% -41.2%

72.8 7.8% 4.2% 0.2% -4.1% -8.8% -13.9% -19.0% -24.5% -30.4% -36.8%

79.1 12.6% 8.9% 5.0% 0.6% -4.0% -9.1% -14.2% -19.7% -25.7% -32.1%

86.0 17.7% 14.0% 10.1% 5.8% 1.1% -4.0% -9.1% -14.6% -20.5% -26.9%

93.5 23.3% 19.7% 15.7% 11.4% 6.7% 1.6% -3.5% -9.0% -14.9% -21.3%

101.6 29.4% 25.7% 21.8% 17.5% 12.8% 7.7% 2.6% -2.9% -8.9% -15.3%

110.5 36.0% 32.3% 28.4% 24.1% 19.4% 14.3% 9.2% 3.7% -2.2% -8.7%

120.1 43.2% 39.5% 35.6% 31.3% 26.6% 21.5% 16.4% 10.9% 4.9% -1.5%

130.5 51.0% 47.3% 43.4% 39.1% 34.4% 29.3% 24.2% 18.7% 12.8% 6.3%

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Recommendations

This design appears to be highly profitable, but further analysis should be performed to discover how

profitable the plant is when certain assumptions are not made. Before this project is fully invested in,

there are multiple experiments and investigations into assumptions that need to be performed to

determine the effect of these assumptions on profitability.

To begin with, one of the main assumptions made in this project is that there is no pressure drop in the

reactor. This is highly unlikely to be true, however. Even when tested in the laboratory reactor there

was a small pressure drop, which while assumed to be negligible, it is possible that with design scale up

the pressure drop is no longer negligible. Experiments on pressure drop across the reactor must be

done before the final decision to invest in this project is made.

Another assumption made in this design that must be further investigated is that the catalyst is fully

active until the moment that it needs to be changed. Catalyst deactivation over time was not taken into

account in this analysis but is likely to be occurring and may affect the rates of reaction. It is possible

that deactivation over time will cause the rates of reaction to slow as it approaches time to replace the

catalyst. This slowing of the reaction rates could potentially not only cause less product to be made, but

also cause the composition into the distillation columns to drastically differ leading to failure of one of

the columns.

The final main assumption made in this initial design was that the heat exchanger system was not

interconnected. This is something that could likely increase the profitability of the plant since the use of

one fluid to heat or cool another fluid would decrease the use of a utility to perform the same job. By

designing a system to handle this, it is likely that the profitability of the plant will increase.

After analysis of these basic assumptions is performed along with other less relevant assumptions,

further discussion of investing and pursuing the proposal may commence.

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Process Alternative

In our development of a design for ethyl acetate process, basic process decisions included:

Operate at a 90% conversion

To recycle ethanol

Alternatively, changing these factors and others lead to different process designs, which we have

concluded to be not as profitable or marginally more difficult to operate and impractical. Because of the

formation of an azeotrope, we designed several separation systems for various reaction conversions. At

conversions lower than 90% (80% and 70%) we produced less ethyl acetate, which resulted in a drastic

change in the separation system design of the three columns. The first column, with a high pressure of

15atm, separated ethyl acetate with the specified purity (99.96% Ethyl Acetate, 40ppm ethanol and

10ppm water) from the azeotrope mixture. It then entered the second and the third column to have the

water removed in an atmospheric pressure of 1atm. The water from the second and the third columns

would exit from the bottoms to the waste water treatment facility and the top stream, which was rich in

ethanol, would recycle back to the reactor for further conversion. This design was not as profitable as

our high conversion design, according to the NPV percent calculations.

For the 90% conversion, we were able to separate ethyl acetate in the first and second distillation

columns in 15atm form the azeotrope mixture and finally separated most of the water before we

recycled the ethanol rich stream. There were three process alternatives:

1. Break the stream into two streams which would enable us to feed a fraction back to the first

distillation column for further separation and the remaining to the reactor.

2. Burn the mixture as fuel to avoid the operating cost of utilities.

Method (1) showed to be effective and showed a slightly higher NPV% in our calculation. However,

once we adjusted the process control design we realized this would significantly complicate the process

control model. Also, the plant would be very sensitive to disturbances because of the third distillation

column as well as becoming increasingly difficult to operate. Knowing an adequate NPV% was already

obtained, we chose method (1), which is easier to operate and still results in a high profitability. Plan (2)

drastically reduced profitability, clearly eliminating it from our conceptual design.

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Assuming the catalyst degradation and effective pressure drop to the system were negligible, simplified

calculations. However, if we were to account for the pressure drop and the diffusion of the particles in

the porous catalysts, it likely that other conversions may be advantageous. Likewise, in more complex

designs a system must be established to separate the diethyl ether, rather than using a split block.

Furthermore, the use of a decanter may assist in the separation of liquid species. We did not see a

necessity to include a decanter in our system, however. Also, to conserve energy, interconnected heat

exchangers running countercurrent may be implemented to reduce the amount of energy needed in the

heat exchangers.

Recommendations

This design appears to be highly profitable, but further analysis should be performed to discover how

profitable the plant is when certain assumptions are not made. Before this project is fully invested in,

there are multiple experiments and investigations into assumptions that need to be performed to

determine the effect of these assumptions on profitability.

Figure 5 Alternative design with a fraction back to the first distillation column for further separation and the remaining to the reactor

15

To begin with, one of the main assumptions made in this project is that there is no pressure drop in the

reactor. This is highly unlikely to be true, however. Even when tested in the laboratory reactor there

was a small pressure drop, which while assumed to be negligible, it is possible that with design scale up

the pressure drop is no longer negligible. Experiments on pressure drop across the reactor must be

done before the final decision to invest in this project is made.

Another assumption made in this design that must be further investigated is that the catalyst is fully

active until the moment that it needs to be changed. Catalyst deactivation over time was not taken into

account in this analysis but is likely to be occurring and may affect the rates of reaction. It is possible

that deactivation over time will cause the rates of reaction to slow as it approaches time to replace the

catalyst. This slowing of the reaction rates could potentially not only cause less product to be made, but

also cause the composition into the distillation columns to drastically differ leading to failure of one of

the columns.

The final main assumption made in this initial design was that the heat exchanger system was not

interconnected. This is something that could likely increase the profitability of the plant since the use of

one fluid to heat or cool another fluid would decrease the use of a utility to perform the same job. By

designing a system to handle this, it is likely that the profitability of the plant will increase.

After analysis of these basic assumptions is performed along with other less relevant assumptions,

further discussion of investing and pursuing the proposal may commence.

Conclusion

Combining the conceptual, economic, ASPEN, and HYSYS designs into a single plant design that

optimizes profitability has shown to be a challenge because of conflicting calculations, numerous

assumptions, and pragmatic judgment decisions. However, the plant designed with these assumptions

and judgment decisions was found to be highly profitable. Though any decision can be scrutinized,

general facts for this plant are that the reactor will be run at 285 oC at a pressure of 1atm and a

conversion of 90%, the separation system will consist of a flash drum and three distillation columns, and

lower reactor temperatures will lead to less profitability. This design could be greatly improved if further

investigation and experiments are performed on the assumptions laid out in the previous section.

16

Appendix A: Tables

Table A.1. Chemical data on all species involved in the chemical plant.

Name Formula Weight (g/mol)

Density (g/cm3)

Melting Point (oC)

Boiling Point (oC)

Ethanol C2H5OH 46.07 0.789 -114 78

Acetaldehyde CH3CHO 44.05 0.788 -123.5 20.2

Hydrogen H2 2.016 8.5*10-5 -259 -252.8

Ethyl Acetate C4H8O2 88.105 0.897 -83.6 77.1

Diethyl Ether C4H10O 74.12 0.7134 -116.3 34.6

Water H2O 18.02 1 0 99.98

Table A.2. This table shows the operating conditions of each piece of equipment involved in the

chemical plant.

Equipment Temperature Pressure

Reactor/Heat Exchanger 558 K 1 atm

Flash Drum 255 K 1 atm

Distillation Column 1 Feed: 335 Reboiler: 464 Condenser: 441

15 atm


15 atm


1 atm

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Table A.3. This table shows the specifications of each piece of equipment involved in the

chemical plant.

Equipment Diameter & Length

Volume & Area

Number of Trays

Material Operating Cost (MM $/yr)

Installation Cost (MM $)

Energy

Reactor/ Heat Exchanger

0.70 m, 109 m

41.5 m3, 8300 m2

NA Carbon Steel

0.09 3.2 7.6*106 kJ/hr

Process Furnace

NA NA NA Carbon Steel

1.1 0.02 2.4*107 kJ/hr

Heat Exchanger 2 (cooling)

NA 63.6 m2 NA Carbon Steel

0.05 0.19 3.8*106 kJ/hr

Heat Exchanger 3 (cooling)

NA 2.7 m2

NA Carbon Steel

0.04 0.02 9.2*105 kJ/hr

Distillation Column 1

3.9 m, 23.2 m

NA 35.1 Carbon Steel

1.4 0.4 Reboiler: 1.6*107 kJ/hr Condenser: 1.8*107 kJ/hr


3.4 m, 20.6 m


1.3 0.3 Reboiler: 1.2*107 kJ/hr Condenser: 1.8*107

kJ/hr


2.9 m, 17.6 m


1.1 0.5 Reboiler: 1.9*107

kJ/hr Condenser: 1.8*107 kJ/hr

18

Table A.4 Proposed Control System Structure (Control loops) for the Reactor/flash/Distillation Unit Plant

Loop Number

Controller Type Controlled Variable Manipulated Variable/Valve

1 Cascade (Secondary) E stream flowrate,w1 A feed stream,V1 2 Feedback Reactor temperature Steam supply,V2 3 Feedback Flash unit feed temperature, TF Cooling water,V3 4 Feedback Flash unit pressure, PF Hydrogen stream,V4 5 Feedback Flash unit liquid level, HF Distillation unit 1 feed,V5 6 Feedback Distillation unit1 pressure,PD1 Top1 stream,V6 7 Feedback Condesner1 liquid level,HC1 Distillation unit 2 Feed,V7 8 Feedback EA composition in top1 stream, xD1 Reflux1 stream,V8 9 Feedback Distillation unit1 liquid level, HD1 Bot1 stream,V9 10 Feedback EA composition in bot1 stream,XB1 Steam supply,V10 11 Feedback Distillation unit2 pressure,PD2 Top2 stream,V11 12 Feedback Condesner2 liquid level,HC2 Distillation unit 3 Feed,V12 13 Feedback Distillation3 Feed temperature, TF3 Cooling water,V13 14 Feedback EA composition in top2 stream, xD2 Reflux2 stream,V14 15 Feedback Distillation unit2 liquid level, HD2 Bot2 stream,V15 16 Feedback EA composition in bot2 stream,XB2 Steam supply,V16 17 Feedback Condesner3 liquid level,HC3 Recycle stream,V17 18 Feedback W composition in top3 stream, xD3 Reflux3 stream,V18 19 Feedback Distillation unit3 liquid level, HD3 Bot3 stream,V19 20 Feedback W composition in bot3 stream,XB3 Steam supply,V20 21 Cascade(Primary) Plant production rate, w2 Set point for w1 (FC1) 22 Feedback E composition in the product, xP Recycle liquid stream,V21 23 Feedback Recycle surge tank liquid level, HS Purge Stream,V22 24 Feedback Surge Tank feed temperature, Ts Cooling water,V23

19

Appendix B: Figures

Figure B.3 Reactor volume increases with increasing temperature and decreases with increasing pressure

Figure B.4 RTI increases as temperature increases and pressure decreases

20

Figure B. 5 NPVproj increases with temperature and decreases with pressure

21

Figure 6 NPV% is optimized at 285K and 1 atm

Figure B. 7 TCI is optimized at 255K and 10 atm

22

Figure B.8 94.6MMKg/yr are produced at 285K and 10 atm

Appendix B.7: ASPEN Plus Distillation system design

Distillation unit 1

23

Distillation unit2

24

Distillation unit 3

25

26

Appendix B.8:

HYSYS Process

Flow Diagram

27

Fla

sh

Un

it

Dis

tillatio

n

Un

it1

Dis

tillatio

n

Un

it2

Dis

tillatio

n

Un

it3

Mix

er

Sp

lit Blo

ck

Die

thyl e

the

r ou

t: 9 M

M k

g/y

r

Wa

ter o

ut: 2

MM

kg

/yr

Hyd

rog

en

Pa

y: 3

.3 M

M $

/yr

Re

cycle

: 15

MM

kg

/yr

Re

acto

r:

Insta

lled

: 3.2

MM

$

Op

era

ting

: 0.0

9 M

M $

/yr

Eth

yl A

ce

tate

Pa

y: 1

30.5

MM

$/y

r

Pu

rge

: 4 M

M k

g/y

r

He

at E

xch

an

ge

r 1:

Insta

lled

: 0.0

2 M

M $

Op

era

ting

: 1.1

MM

$/y

r

Eth

an

ol C

ost: 8

5.1

MM

$/y

rH

ea

t Exch

an

ge

r 2:

Insta

lled

: 0.1

9 M

M $

Op

era

ting

: 0.0

5 M

M $

/yr

He

at E

xch

an

ge

r 3:

Insta

lled

: 0.0

2 M

M $

Op

era

ting

: 0.0

4 M

M $

/yr

Dis

tillatio

n 1

:

Insta

lled

: 0.4

MM

$

Op

era

ting

: 1.4

MM

$/y

r

Dis

tillatio

n 2

:

Insta

lled

: 0.3

MM

$

Op

era

ting

: 1.3

MM

$/y

r

Dis

tillatio

n 3

:

Insta

lled

: 0.5

MM

$

Op

era

ting

: 1.1

MM

$/y

r

Appendix B.9:

Economic

Process Flow

Diagram

28

Fla

sh

Un

it

Dis

tillatio

n

Un

it1

Dis

tillatio

n

Un

it2

Dis

tillatio

n

Un

it3

Mix

er

S

plit B

lock

Die

thyl e

the

r ou

t: 9 M

M k

g/y

r

Wa

ter o

ut: 2

MM

kg

/yr

Hyd

rog

en

ou

t: 5 M

M k

g/y

r

Re

cycle

: 15

MM

kg

/yr

Re

acto

r

Eth

yl A

ce

tate

ou

t: 10

0 M

M k

g/y

r

Eth

an

ol F

ee

d: 1

20

MM

kg

/yr

Pu

rge

: 4 M

M k

g/y

r

Appendix B.10:

Process Flow

Diagram

29

Appendix C: Discounted Cash Flow Statement

Table C.1. The discounted cash flow sheet for the base case design based off of the HYSYS and ASPEN

analysis.

Fla

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Un

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24

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19

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17

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15

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16

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21

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r

Appendix B.11:

Process

Control

Diagram

30

Table C.2. Discounted cash flow sheet with no numbers but notations to where each variation had a

direct change to the sheet.

Discounted Cash Flow Table

Establishes the ROI_BT Based on TI (also TCI ) and the Annual % Increase of NPV (Normalized by TCI )

Fixed Capital and Profit_BT are the two independent variables.

CR and FR can be chosen independently. WC and SV are converted from inventory to Profit_BT in Year 10.

All dollar amounts in table represent millions of dollars.

Profit_BT = 43.58

Construction Rate 10.0% Tax Rate 50% N_construction 2 using Capitalization =

Finance Rate 8.0% N_operation 10 Yields Tot.Cap.Inv. TI =FC+WC+SU

ROI_BT = 151.2% 152.5%

Fixed Capital 8.4

alpha_Work ing Capital 2.31 a-2 0.00 b_1 0.5 TI =FC+WC+SU 28.6

alpha_Start-Up Capital 0.10 a-1 0.30 b_2 0.8

alpha_Salvage Value 0.03 a0 0.70 b_3 1.0 TCI =Tot.Cap.Inv. 28.8

Capital In (+) Discount Discounted

YearDesignConstruction Periodor Out (-) Factors Cash Flows

-2 Fixed Capital in Y-2 0.0 1.210 0.0

-1 Fixed Capital in Y-1 -2.5 1.100 -2.8

0 Fixed Capital in Y0 -5.9 1.000 -5.9

0 Work ing Capital -19.3 1.000 -19.3

0 Start-Up Capital -0.8 1.000 -0.8

0 Total of Capital Outlays -28.8

(=Sum of Constr. DCFs)

0 Total Capital Investment 28.8

(=Proceeds of Bond Issue)

Profit Bond Depreciation Profit Cash

Operations Period Before Taxes Financing Allowed After Taxes Flows

1 21.8 -2.3 -0.9 9.3 10.2 0.926 9.4

2 34.9 -2.3 -0.9 15.8 16.7 0.857 14.4

3 41.4 -2.3 -0.9 19.1 20.0 0.794 15.9

4 43.6 -2.3 -0.9 20.2 21.1 0.735 15.5

5 43.6 -2.3 -0.9 20.2 21.1 0.681 14.4

6 43.6 -2.3 -0.9 20.2 21.1 0.630 13.3

7 43.6 -2.3 -0.9 20.2 21.1 0.583 12.3

8 43.6 -2.3 -0.9 20.2 21.1 0.540 11.4

9 43.6 -2.3 -0.9 20.2 21.1 0.500 10.6

10 43.6 -2.3 -0.9 20.2 21.1 0.463 9.8

10 Work ing Capital 19.3 9.7 9.7 0.463 4.5

10 Salvage Value 0.3 0.1 0.1 0.463 0.1

10 Pay-Off TCI -28.8 -28.8 0.463 -13.4

WC & SV Total Profit Bond Interest Total Total Profit Total NPV (0) NPV-proj [NPV (0)

All Figures Represent Recovery Before Taxes Payments Depreciation After Taxes Cash Flow Discounted to EOY(-2)]

PV of Operations==> 4.5 263.1 -15.5 6.2 125.2 118.1 118.1 101.2

Bond Total Capital

Repayment Recovery NPV Increase per Year

-13.4 10.7 normalized/annualized

Net Present NPV (0) Avg. NPV_proj Avg.

Value of Bonds Over 10 Years Over 12 Years

0.0 41.0% 29.3%

31

Table C.3. The discounted cash flow sheet for Variation 1.a in which the price of ethanol increases. This

change had the largest percent deviation in the NPVzero and NPVproj from the base case design at 96.74%.


Profit_BT =

Variation 1 &

2, Sell Price

Drop, Raw

Material Raise

Construction Rate Variation 3 Tax Rate Abnormal 1 N_construction 2

Finance Rate

Variation 4,

IRR N_operation 10 Yields Tot.Cap.Inv. TI =FC+WC+SU

ROI_BT = xx xx

Fixed Capital 8.4

alpha_Work ing Capital 2.31 a-2 0.00 b_1 0.5 TI =FC+WC+SU xx


alpha_Salvage Value 0.03 a0 0.70 b_3 1.0 TCI =Tot.Cap.Inv. xx

b_4 Abnormal 2

Capital In (+) b_5 Abnormal 2 Discount Discounted

YearDesignConstruction Periodor Out (-) b_6 Abnormal 2 Factors Cash Flows

b_7 Abnormal 2

-2 Fixed Capital in Y-2 xx b_8 Abnormal 2 xx xx

-1 Fixed Capital in Y-1 xx b_9 Abnormal 2 xx xx

0 Fixed Capital in Y0 xx b_10 Abnormal 2 xx xx

0 Work ing Capital xx xx xx

0 Start-Up Capital xx xx xx

0 Total of Capital Outlays xx


0 Total Capital Investmentxx




1 xx xx xx xx xx xx xx










10 Work ing Capital xx xx xx xx xx

10 Salvage Value xx xx xx xx xx

10 Pay-Off TCI xx xx xx xx



PV of Operations==> xx xx xx xx xx xx xx xx

Bond Total Capital

Repayment Recovery

xx xx



xx xx xx

normalized/annualized





using Capitalization =

NPV Increase per Year

32

Table C.4. The discounted cash flow sheet for Variation 2.b in which the value of the ethyl acetate

decreases. This change had the second largest percent deviation in the NPVzero and NPVproj from the base

case design at 51.31%.


Profit_BT = 7.11

Construction Rate 10.0% Tax Rate 50% N_construction 2


ROI_BT = 19.1% 19.3%

Fixed Capital 8.4









0 Work ing Capital -27.6 1.000 -27.6








1 3.6 -3.0 -0.9 -0.2 0.8 0.926 0.7

2 5.7 -3.0 -0.9 0.9 1.8 0.857 1.6

3 6.8 -3.0 -0.9 1.4 2.4 0.794 1.9

4 7.1 -3.0 -0.9 1.6 2.5 0.735 1.9

5 7.1 -3.0 -0.9 1.6 2.5 0.681 1.7

6 7.1 -3.0 -0.9 1.6 2.5 0.630 1.6

7 7.1 -3.0 -0.9 1.6 2.5 0.583 1.5

8 7.1 -3.0 -0.9 1.6 2.5 0.540 1.4

9 7.1 -3.0 -0.9 1.6 2.5 0.500 1.3

10 7.1 -3.0 -0.9 1.6 2.5 0.463 1.2

10 Work ing Capital 27.6 13.8 13.8 0.463 6.4

10 Salvage Value 0.3 0.1 0.1 0.463 0.1

10 Pay-Off TCI -37.1 -37.1 0.463 -17.2



PV of Operations==> 6.5 42.9 -19.9 6.2 14.8 3.8 3.8 3.3

Bond Total Capital

Repayment Recovery

-17.2 12.6



0.0 1.0% 0.7%








33

Table C.5. The discounted cash flow sheet for Variation 3.a in which the construction rate increases by

5%. This change had the smallest percent deviation in the NPVzero and NPVproj from the base case design

at 0.08%.


Profit_BT = 23.51



ROI_BT = 81.6% 82.3%

Fixed Capital 8.4









0 Work ing Capital -19.3 1.000 -19.3








1 11.8 -2.3 -0.9 4.3 5.2 0.926 4.8

2 18.8 -2.3 -0.9 7.8 8.7 0.857 7.5

3 22.3 -2.3 -0.9 9.6 10.5 0.794 8.3

4 23.5 -2.3 -0.9 10.1 11.1 0.735 8.1

5 23.5 -2.3 -0.9 10.1 11.1 0.681 7.5

6 23.5 -2.3 -0.9 10.1 11.1 0.630 7.0

7 23.5 -2.3 -0.9 10.1 11.1 0.583 6.5

8 23.5 -2.3 -0.9 10.1 11.1 0.540 6.0

9 23.5 -2.3 -0.9 10.1 11.1 0.500 5.5

10 23.5 -2.3 -0.9 10.1 11.1 0.463 5.1

10 Work ing Capital 19.3 9.7 9.7 0.463 4.5

10 Salvage Value 0.3 0.1 0.1 0.463 0.1

10 Pay-Off TCI -28.8 -28.8 0.463 -13.4



PV of Operations==> 4.5 141.9 -15.5 6.2 64.7 57.5 57.5 49.3

Bond Total Capital

Repayment Recovery

-13.4 10.7



0.0 19.9% 14.3%








34


Profit_BT = 43.58



ROI_BT = 150.5% 152.5%

Fixed Capital 8.4









0 Work ing Capital -19.3 1.000 -19.3








1 21.8 -2.3 -0.9 9.3 10.2 0.926 9.4

2 34.9 -2.3 -0.9 15.8 16.7 0.857 14.3

3 41.4 -2.3 -0.9 19.1 20.0 0.794 15.9

4 43.6 -2.3 -0.9 20.2 21.1 0.735 15.5

5 43.6 -2.3 -0.9 20.2 21.1 0.681 14.4

6 43.6 -2.3 -0.9 20.2 21.1 0.630 13.3

7 43.6 -2.3 -0.9 20.2 21.1 0.583 12.3

8 43.6 -2.3 -0.9 20.2 21.1 0.540 11.4

9 43.6 -2.3 -0.9 20.2 21.1 0.500 10.6

10 43.6 -2.3 -0.9 20.2 21.1 0.463 9.8

10 Work ing Capital 19.3 9.7 9.7 0.463 4.5

10 Salvage Value 0.3 0.1 0.1 0.463 0.1

10 Pay-Off TCI -29.0 -29.0 0.463 -13.4



PV of Operations==> 4.5 263.1 -15.5 6.2 125.2 118.0 118.0 101.2

Bond Total Capital

Repayment Recovery

-13.4 10.7



0.0 40.8% 29.1%








35

Appendix D: MATLAB code

Design2Varying clear; clc; %number of plots making PlotNumber = 1; %need to know values %selectivity s = 10/11; %flow rates Pea_kg = 100*10^6; %kg/yr Pea = Pea_kg*1000*(1/88.105); %mol/yr [=] MM kg/yr*1000g/kg*mol/g Ps = 2*Pea; %mol/yr Fa = 2*Pea/s; %mol/yr Pde = Pea*(1/s - 1); %mol/yr Pw = Pea*(1/s - 1); %mol/yr %density in g/L = g/cm^3*(1 cm^3/0.001 L) da = 0.789*(1/0.001); CuOdensity = 6.31*(1/0.001); CoOdensity = 6.44*(1/0.001); Cr2O3density = 5.22*(1/0.001); catdensity = 0.94*CuOdensity+0.05*CoOdensity+0.01*Cr2O3density; %molecular weight in g/mol MWa = 46.07; %molar densities = density*(1/MW) mda = da/MWa; %g/L*(mol/g) = mol/L %costs Cost_pea = 1.30; %ethyl acetate $/kg Cost_a = 0.70; %ethanol $/kg Cost_s = 0.31; %hydrogen $/lb %reactor conversion xa = 0.01:0.01:1; %equipment temperature and pressure reactT = 498:20:558; %K reactP = 1:3:10; %atm

36

%need to know values R = 1.987; %cal mol^-1 K^-1 R2 = 0.082057; %L atm K^-1 mol^-1 for m = 1:1:PlotNumber for p = 1:1:length(reactP) for t = 1:1:length(reactT) for h = 1:1:length(xa) %calculating recycle flowrate = mol/yr Ra(h) = 2*Pea/s*((1-xa(h))/xa(h)); %total molar flow rate leaving reactor = mol/yr TFLR(h) = Pea + Ps + Pde + Pw + Ra(h); %mole fractions leaving reactor za(h) = Ra(h)/TFLR(h); zea(h) = Pea/TFLR(h); zs(h) = Ps/TFLR(h); zde(h) = Pde/TFLR(h); zw(h) = Pw/TFLR(h); %total molar flow rate to separation system = mol/yr %after removing all diether with split block %after removing all hydrogen TFTS(h) = Pea + Pw + Ra(h); %mole fractions zas(h) = Ra(h)/TFTS(h); zeas(h) = Pea/TFTS(h); zws(h) = Pw/TFTS(h); %solving for reactor volume %K's Ka(t) = exp(5890/(R*reactT(t))-6.40); %atm^-1 Ks(t) = exp(6850/(R*reactT(t))-7.18); %atm^-1 k(t) = exp(-16130/(R*reactT(t))+16.25); %mol hr^-1 g-cat^-1 Pain(p,h) = reactP(p); %(mol/L)*(K)*(L atm K^-1 mol^-1) = atm Paout(p,h) = reactP(p)*za(h); %atm step(p,h) = (Pain(p,h) - Paout(p,h))/100; Pa = Paout(p,h):step(p,h):Pain(p,h); %mol/L Psin(p,h) = 0; Psout(p,h) = reactP(p)*zs(h); step2(p,h) = (Psin(p,h) - Psout(p,h))/100; ps = Psout(p,h):step2(p,h):Psin(p,h); for g = 1:length(Pa) %concentration of acetaldehyde is zero funct(h,g) = 20*10*(1+Ka(t)*Pa(g)+Ks(t)*ps(g))^2/(k(t)*Ka(t)*Pa(g));

37

end AreaUnderCurve(p,t,h) = trapz(Pa,funct(h,:)); %hr g-cat/L tau(p,t,h) = AreaUnderCurve(p,t,h); %catalyst weight catW(p,t,h) = Fa*tau(p,t,h)*(1/8765.81277); %mol/yr*(hr g-cat/mol)*conversion = g-cat %reactor volume [=] m^3 Vreact(p,t,h) = 2*catW(p,t,h)*(1/catdensity)*0.001; end end end end for p = 1:1:length(reactP) for t = 1:1:length(reactT) for h = 1:1:length(xa) vol(h) = Vreact(p,t,h); end if p==1 && t==1 plotcolor = 'b'; end if p==2 && t==1 plotcolor = 'g'; end if p==3 && t==1 plotcolor = 'r'; end if p==4 && t==1 plotcolor = 'k'; end if p==1 && t==2 plotcolor = 'b'; end if p==2 && t==2 plotcolor = 'g'; end if p==3 && t==2 plotcolor = 'r'; end if p==4 && t==2 plotcolor = 'k'; end if p==1 && t==3 plotcolor = 'b'; end if p==2 && t==3 plotcolor = 'g'; end

38

if p==3 && t==3 plotcolor = 'r'; end if p==4 && t==3 plotcolor = 'k'; end if p==1 && t==4 plotcolor = 'b'; end if p==2 && t==4 plotcolor = 'g'; end if p==3 && t==4 plotcolor = 'r'; end if p==4 && t==4 plotcolor = 'k'; end hold on; fig1 = figure(1); set(fig1,'Color','white') ylim([0 100]) ylabel('Reactor Volume (m^3)') xlabel('Reactor Conversion') plot(xa,vol,plotcolor) end end Design2Conversion clear; reply = input('Reactor Temperature in Celsius: ', 's'); if reply == '225' reactTC = 225; plotcolor = 'c'; end if reply == '255' reactTC = 255; plotcolor = '--b'; end if reply == '285' reactTC = 285; plotcolor = '--k'; end %number of plots making PlotNumber = 1;

39

%molecular weight in g/mol MWa = 46.07; MWea = 88.105; MWw = 18.02; MWs = 2.016; MWde = 74.12; %need to know values %selectivity s = 10/11; %flow rates Pea_kg = 100*10^6; %kg/yr Pea = Pea_kg*1000*(1/88.105); %mol/yr [=] MM kg/yr*1000g/kg*mol/g Ps = 2*Pea; %mol/yr Fa = 2*Pea/s; %mol/yr Pde = Pea*(1/s - 1); %mol/yr Pw = Pea*(1/s - 1); %mol/yr Fa_kg = Fa*MWa*(1/1000); %kg/yr Pde_kg = Pde*MWde*(1/1000); %kg/yr %density in g/L = g/cm^3*(1 cm^3/0.001 L) da = 0.789*(1/0.001); dea = 0.897*(1/0.001); dw = 1*(1/0.001); ds = 8.5*10^-5*(1/0.001); %density in g/cm^3 CuOdensity = 6.31; CoOdensity = 6.44; Cr2O3density = 5.22; catdensity = 0.94*CuOdensity+0.05*CoOdensity+0.01*Cr2O3density; dagcm = 0.789; %molar densities = density*(1/MW) mda = da/MWa; %g/L*(mol/g) = mol/L mdea = dea/MWea; %mol/L mdw = dw/MWw; %mol/L mds = ds/MWs; %mol/L %costs Cost_pea = 1.30; %ethyl acetate $/kg Cost_a = 0.70; %ethanol $/kg Cost_s = 0.31; %hydrogen $/lb Cost_cat = 4; %catalyst $/kg electricity_cost = 0.06; %$/kWh

40

%operating hours operating_hours = 8000; %hours/yr %reactor conversion %recycle, no recycle, recycle, no recycle xa = [0.8 0.8 0.9 0.9]; %equipment temperature and pressure reactT = reactTC+273; %K reactP = 10; %atm %need to know values R1 = 1.987; %cal mol^-1 K^-1 R2 = 0.082057; %L atm K^-1 mol^-1 T_in = reactT; %K %heat of formation H_R=-170.7; %kJ/mol (gas) {liquid=-196.4} H_EA=-445; %kJ/mol (gas) {liquid=-480} H_A=-235.3; %kJ/mol (gas) {liquid=-277} H_DE=-252.7; %kJ/mol (gas) {liquid=-271.2} H_W=-241.83; %kJ/mol (gas) {liquid=-285.83} dH0_a= 68.9*1000; %J/mol dH0_b= ((H_EA)-(2*H_R))*1000; %J/mol dH0_c = ((H_DE+H_W)-(2*H_A))*1000; %J/mol %specific heat capacity of ethanol Cp_a=87.53; %J mol^-1 K^-1 %reference temperature T_ref=298; %K %Heat of Reaction QrJ = (dH0_a*Fa+dH0_b*Pea+dH0_c*Pde)*10^-6; %MJ/yr Qr = QrJ*(1/315569265); %MW %marshal and swift factor for 210Q1 MS=1461.3; %Reactor %shell material use carbon steel Fm_react = 1.00; %P = 1 atm Fp = 1; P = 5 atm Fp = 1.05; P = 10 atm Fp = 1.15

41

%1 atm = 14.7 psi Fp_react = 1.00; Fc_react = Fm_react + Fp_react; Fm_heat=1.00; %material factor for CS/Ti Fp_heat=0; %pressure factor for up to 150psi Fd_heat=1.0; %type factor Floating head Fh_heat=(Fd_heat+Fp_heat)*Fm_heat; %F factor from gutheric Apendix E.2 of Douglas %Distillation Column 1 dist1P = [14, 14, 15, 14]; %atm z_D1Bea = [1, 1, 1, 1]; z_D1Ba = [0, 0, 0, 0]; z_D1Bw = [0, 0, 0, 0]; z_D1Tea = [0.1, 0.1, 0.107, 0.150]; z_D1Ta = [0.761, 0.7329, 0.633, 0.5836]; z_D1Tw = [0.139, 0.1671 0.260, 0.2664]; D1Bflowkmols = [0.033, 0.0334, 0.034, 0.0334]; D1Bflow = D1Bflowkmols*1000*31556926; %kmol/s*(1000mol/kmol)*(31556926s/yr) = mol/yr D1Tflowkmols = [0.0259, 0.0215, 0.0138, 0.0135]; D1Tflow = D1Tflowkmols*1000*31556926; %mol/yr D1stages = [33.9, 34.1, 36.0, 29.3]; D1cond = [-8.6073*10^6, -7.1807*10^6, -4.6221*10^6, -4.4782*10^6]; %J/s D1reboil = [7.2309*10^6, 6.1347*10^6, 4.1742*10^6, 4.1352*10^6]; %J/s Tcin1 = [344, 344, 345, 345]; %K Tcout1 = [438, 438, 441, 438]; %K Treboil1 = [461, 461, 464, 461]; %K %Distillation Column 2 dist2P = [15, 15, 15, 15]; %atm z_D2Bea = [1, 1, 1, 1]; z_D2Ba = [0, 0, 0, 0]; z_D2Bw = [0, 0, 0, 0]; z_D2Tea = [0.080, 0.080, 0.105, 0.130]; z_D2Ta = [0.7779, 0.7492, 0.6344, 0.5974]; z_D2Tw = [0.1421, 0.1708, 0.2606, 0.2726]; D2Bflowkmols = [5.6317*10^-4, 4.675*10^-4, 3.0869*10^-5, 3.0983*10^-4]; D2Bflow = D2Bflowkmols*1000*31556926; %kmol/s*(1000mol/kmol)*(31556926s/yr) = mol/yr D2Tflowkmols = [0.0253, 0.021, 0.0138, 0.0132]; D2Tflow = D2Tflowkmols*1000*31556926; %mol/yr D2stages = [31.3, 29.9, 31.6, 30.6]; D2cond = [-8.3974*10^6, -7.009*10^6, -4.6165*10^6, -4.3697*10^6]; %J/s D2reboil = [5.9406*10^6, 4.9313*10^6, 3.2247*10^6, 3.0864*10^6]; %J/s Tcin2 = [438, 438, 441, 438]; %K Tcout2 = [441, 441, 441, 441]; %K

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Treboil2 = [464, 464, 464, 464]; %K %Distillation Column 3 dist3P = [1, 1, 1, 1]; %atm z_D3Bea = [0, 0, 0, 0]; z_D3Ba = [0, 0, 0, 0]; z_D3Bw = [1, 1, 1, 1]; z_D3Tea = [0.083, 0.0859, 0.1264, 0.1591]; z_D3Ta = [0.807, 0.8041, 0.7636, 0.7309]; z_D3Tw = [0.110, 0.110, 0.110, 0.110]; D3Bflowkmols = [9.1453*10^-4, 1.4369*10^-3, 2.3329*10^-3, 2.4065*10^-3]; D3Bflow = D3Bflowkmols*1000*31556926; %kmol/s*(1000mol/kmol)*(31556926s/yr) = mol/yr D3Tflowkmols = [0.0244, 0.0196, 0.0115, 0.0108]; D3Tflow = D3Tflowkmols*1000*31556926; %mol/yr D3stages = [31.5, 31.3, 27.8, 26.2]; D3cond = [-1.0294*10^7, -8.253*10^6, -4.7675*10^6, -4.4408*10^6]; %J/s D3reboil = [1.0549*10^7, 8.458*10^6, 4.9243*10^6, 4.6167*10^6]; %J/s Tcin3 = [349, 349, 349, 348]; %K Tcout3 = [348, 348, 348, 347]; %K Treboil3 = [373, 373, 373, 373]; %K %Compressor 1 gamma1 = 0.23; pg.154 %reactP is in atm %atm*(2116.22 lbf/ft^2)/atm = lbf/ft^2 Pin_comp1 = reactP*2116.22; %lbf/ft^2 %dist1P is in atm Pout_comp1 = dist1P*2116.22; %lbf/ft^2 Fc_comp1 = 1.00; %centrifugal, motor %Compressor 2 gamma2 = 0.23; pg.154 %dist1P is in atm %atm*(2116.22 lbf/ft^2)/atm = lbf/ft^2 Pin_comp2 = dist1P*2116.22; %lbf/ft^2 %dist2P is in atm Pout_comp2 = dist2P*2116.22; %lbf/ft^2 Fc_comp2 = 1.00; %Compressor 3 gamma3 = 0.23; pg.154 %dist2P is in atm %atm*(2116.22 lbf/ft^2)/atm = lbf/ft^2 Pin_comp3 = dist2P*2116.22; %lbf/ft^2 %dist3P is in atm Pout_comp3 = dist3P*2116.22; %lbf/ft^2 Fc_comp3 = 1.00;

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%Cooling fluid == water Tfin = 4.4+273; %Cooling fluid temperature in K Tfout = 10+273; %Cooling fluid temperature out in K %fraction of start up capital = SU/FC = alpha_su alpha_su = 0.1; %fraction of fixed capital spent in start up years a_2 = 0; a_1 = 0.3; a0 = 0.7; %fraction of p_bt obtained in start up years = 3 startupyrs = 3; b1 = 0.5; b2 = 0.8; b3 = 1.0; %construction rate = CR, finance rate = FR, tax rate = TR, complementary tax rate = TRC CR = 0.1; FR = 0.08; TR = 0.5; TRC = 1-TR; %sigma_b = used in NPV0 calculation; note need to change if start up yr > 3 sigma_b = b1*(1+FR)^-1+b2*(1+FR)^-2+b3*(1+FR)^-3+(1+FR)^-3*((1-(1+FR)^-7)/FR); %sigma = used in NPV0 calculation; not for n = 10 years = lifetime of plant sigma = (1-(1+FR)^-10)/FR; % needed values PS = 2; % project start time PL = 12; % project lifetime for m = 1:1:PlotNumber for h = 1:1:length(xa) %calculating recycle flowrate = mol/yr if h == 1 || h == 3 Ra(h) = 2*Pea/s*((1-xa(h))/xa(h)); end if h ==2 || h == 4 Ra(h) = Fa*(1-xa(h)); end %calculating total ethanol flow to reactor = mol/yr Fain(h) = Fa + Ra(h); %outlet temperature of reactor T_out2(h)=T_in-(dH0_a*Fa+dH0_b*Pea+dH0_c*Pde)/(Fain(h)*Cp_a); %total molar flow rate leaving reactor = mol/yr TFLR(h) = Pea + Ps + Pde + Pw + Ra(h); %mole fractions leaving reactor za(h) = Ra(h)/TFLR(h); zea(h) = Pea/TFLR(h); zs(h) = Ps/TFLR(h); zde(h) = Pde/TFLR(h); zw(h) = Pw/TFLR(h); %total molar flow rate to separation system = mol/yr %after removing all diether with split block TFTS(h) = Pea + Pw + Ra(h) + Ps;

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%mole fractions zas(h) = Ra(h)/TFTS(h); zeas(h) = Pea/TFTS(h); zws(h) = Pw/TFTS(h); zss(h) = Ps/TFTS(h); %solving for reactor volume %K's Ka = exp(5890/(R1*reactT)-6.40); %atm^-1 Ks = exp(6850/(R1*reactT)-7.18); %atm^-1 k = exp(-16130/(R1*reactT)+16.25); %mol hr^-1 g-cat^-1 Pain(h) = reactP; %(mol/L)*(K)*(L atm K^-1 mol^-1) = atm Paout(h) = reactP*za(h); %atm step(h) = (Pain(h) - Paout(h))/100; Pa = Paout(h):step(h):Pain(h); %mol/L Psin(h) = 0; Psout(h) = reactP*zs(h); step2(h) = (Psin(h) - Psout(h))/100; ps = Psout(h):step2(h):Psin(h); for g = 1:length(Pa) %concentration of acetaldehyde is zero funct(h,g) = 20*10*(1+Ka*Pa(g)+Ks*ps(g))^2/(k*Ka*Pa(g)); end AreaUnderCurve(h) = trapz(Pa,funct(h,:)); %hr g-cat/L tau(h) = AreaUnderCurve(h); %catalyst weight catW(h) = Fa*tau(h)*(1/8765.81277); %mol/yr*(hr g-cat/mol)*conversion = g-cat %reactor volume [=] g-cat*(cm^3/g-cat)*(m^3/1000000 cm^3) = m^3 Vcat(h) = catW(h)*(1/catdensity); %cm^3 Vreact(h) = 2*Vcat(h)*(10^-6); %m^3 %finding dimensions of reactor %assume tube diameter is 2 cm, and length of reactor is 10 m d = 2/100; %m r = d/2; %m hcyl(h) = Vreact(h)/(pi*r^2); %m %number of tubes in heat exchanger Ntubes(h) = hcyl(h)/10; %surface area of tubes A(h) = 2*pi*r*10*Ntubes(h); %m^2 Aft(h) = A(h)*10.76; %ft^2 %Reynolds number Calculations NEED TO CHECK %mol/yr*g/mol*cm^3/g*1/m^2*m^3/cm^3*yr/s = m/s velocity(h) = Fain(h)*MWa*(1/dagcm)*(1/(pi*r^2))*(1/100)^3*(1/31556926); %PROBLEM HERE???

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visa = 0.001095; %kg m/s^2 s/m^2 = kg s^-1 m^-1 %g/cm^3*m/s*m*m*s/kg*1kg/1000g*(100cm/m)^3 = unitless Re(h) = dagcm*velocity(h)*d*(1/visa)*(1/1000)*100^3; %Flash Drum Calculations %We assume there is a split block between the reactor and the reactants that gets rid of diethyl %We need to bring the feed's temperature as low as 320K to flash the hydrogen from the rest of the species %componenets entering the flash drum %%% Hydrogen BP: 20.3K (-253C) %Antoine Constants in K and Bar A1=3.54314; B1=99.395; C1=7.726; % Ethyl Accetate BP: 350K (77.1C) %Antoine Constants in K and Bar A2=4.22809; B2=1245.702; C2=-55.189; % Ethanol BP: 351K (78C) %Antoine Constants in K and Bar A3=5.37229; B3=1670.409; C3=-40.191; % Water BP: 373K (100C) %Antoine Constants in K and Bar A4=4.65430; B4=1435.264; C4=-64.848; z_E(h) = zas(h); z_EA(h) = zeas(h); z_H(h) = zss(h); z_W(h) = zws(h); F_flash(h) = TFTS(h); %mol/yr %Operating Pressure P_drum=1; %bar %Operating Temperature Note:we need to adjust this to get maximum separation T_drum=255; %K %Antoine equation %psat all in bar psat_H=10^(A1-B1/(T_drum+C1)); %bar psat_EA=10^(A2-B2/(T_drum+C2)); %bar psat_E=10^(A3-B3/(T_drum+C3)); %bar psat_W=10^(A4-B4/(T_drum+C4)); %bar K1=psat_H/P_drum; K2=psat_EA/P_drum; K3=psat_E/P_drum; K4=psat_W/P_drum; k1=1/(K1-1); k2=1/(K2-1); k3=1/(K3-1); k4=1/(K4-1); % Solve Rachford-Rice equation numerically to find a=V/F: a(h)=fzero(@(a) z_H(h)/(k1+a) + z_EA(h)/(k2+a) + z_E(h)/(k3+a) + z_W(h)/(k4+a) , .5); %Hydrogen molar composition in the bottoms xflash_H(h)=z_H(h)/(1+a(h)*(K1-1)); %Ethyl acetate's composition in the bottoms xflash_EA(h)=z_EA(h)/(1+a(h)*(K2-1)); %Ethanol molar composition in the bottoms xflash_E(h)=z_E(h)/(1+a(h)*(K3-1));

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%water molar composition in the bottoms xflash_W(h)=z_W(h)/(1+a(h)*(K4-1)); %Hydrogen molar composition in the vapor stream yflash_H(h)=K1*xflash_H(h); %Ethyl acetate's composition in the vapor stream yflash_EA(h)=K2*xflash_EA(h); %Ethanol molar composition in the vapor stream yflash_E(h)=K3*xflash_E(h); %water molar composition in the vapor stream yflash_W(h)=K4*xflash_W(h); %Bottoms flow rate L_flash(h)=(1-a(h))*F_flash(h); %Vapor stream flow rate V_flash(h)=a(h)*F_flash(h); %Gamma need to be varying??? %Compressor 1 %molar density of fluid to compressor [=] mol/L mdcompfluid1(h) = xflash_H(h)*mds + xflash_EA(h)*mdea + xflash_E(h)*mda + xflash_W(h)*mdw; %volumetric flow rate = flow rate to compressor/molar density [=] ft^3/min %mol/yr*L/mol*0.0353146667ft^3/L*(1yr/525948.766min) = ft^3/min qin1(h) = L_flash(h)*(1/mdcompfluid1(h))*0.0353146667*(1/525948.766); hp1(h) = ((3.03*10^-5)/gamma1)*Pin_comp1*qin1(h)*((Pout_comp1(h)/Pin_comp1)^gamma1-1); bhp1(h) = hp1(h)/0.8; %Distillation Columns Height_1(h) = 3*0.61 + 0.61*D1stages(h); Height_2(h) = 3*0.61 + 0.61*D2stages(h); Height_3(h) = 3*0.61 + 0.61*D3stages(h); Diameter_1(h) = 1/6*Height_1(h); Diameter_2(h) = 1/6*Height_2(h); Diameter_3(h) = 1/6*Height_3(h); PeaT(h) = D1Bflow(h)*z_D1Bea(h) + D2Bflow(h)*z_D2Bea(h) + D3Bflow(h)*z_D3Bea(h); PsT(h) = yflash_H(h)*V_flash(h); PwT(h) = D1Bflow(h)*z_D1Bw(h) + D2Bflow(h)*z_D2Bw(h) + D3Bflow(h)*z_D3Bw(h); PaBot(h) = D1Bflow(h)*z_D1Ba(h) + D2Bflow(h)*z_D2Ba(h) + D3Bflow(h)*z_D3Ba(h); PeaT_kg(h) = PeaT(h)*MWea*(1/1000); %kg/yr PsT_kg(h) = PsT(h)*MWs*(1/1000); %kg/yr PwT_kg(h) = PwT(h)*MWw*(1/1000); %kg/yr PaBot_kg(h) = PaBot(h)*MWa*(1/1000); %need in kg/yr IN(h) = Fa_kg; OUT(h) = Pde_kg + PeaT_kg(h) + PsT_kg(h) + PwT_kg(h) + PaBot_kg(h); %Costing % Revenue % R = EA val + H2 val - A cost (Water Sellable???) % R($/yr) = ($/kg)*(mol/yr)*(g/mol [MW])*(kg/1000g) [=] $/yr

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Sell(h) = Cost_pea*PeaT(h)*MWea*(1/1000); Pay(h) = Cost_a*Fa*MWa*(1/1000); %$/lb*mol/yr*(g/mol [MW])*(0.0022lbs/g) Extra(h)= Cost_s*PsT(h)*MWs*0.0022; R(h) = Sell(h)-Pay(h)+Extra(h); %Operating Cost Calculations % Operating Costs are the utility cost to run that piece of equipment % C = React + Sep System %cm^3*$/kg*g/cm^3*1kg/1000g*2/yr CatCost(h) = Vcat(h)*Cost_cat*catdensity*(1/1000)*2; %use Dowtherm over steam %using a process furnace to circulate the Dowtherm %operating cost is the cost fuel %$/kg*MJ/yr*kg/kJ*10^6J/MJ*1kJ/1000J %Cost_Heat(h) = reactsteamcost*abs(QrJ)*(1/Hsteamreact)*(10^6/1000); %MJ/yr*(1.62 $/MM Btu)*(10^6 J/MJ)*(1 Btu/1055 J)*(1 MM Btu/10^6 Btu) = $/yr Cost_Heat(h) = abs(QrJ)*1.62/1055; React(h) = CatCost(h) + Cost_Heat(h); %$/yr %compressor operating cost utilityreq1(h) = bhp1(h)/0.9; %hp hptokW = 0.75; %1 hp = 0.75 kW [=] kW/hp opcostcomp1(h) = utilityreq1(h)*hptokW*operating_hours*electricity_cost; %distillation operating cost %Condenser Costs = coolant %coolant = refrigerated water coolant_cost = 5.7; %$/GJ %$/GJ*J/s*(GJ/10^9 J)*(31556926s/yr) C_1(h) = coolant_cost*abs(D1cond(h))*(31556926)*10^-9; C_2(h) = coolant_cost*abs(D2cond(h))*(31556926)*10^-9; C_3(h) = coolant_cost*abs(D3cond(h))*(31556926)*10^-9; %Reboiler = heater %$/kg steam_cost1 = 6.74/1000; steam_cost2 = 2.38/1000; %kJ/kg H_steam1 = 1755; H_steam2 = 2213; %$/kg*kg/kJ*J/s*31556926s/yr*1kJ/1000J R_1(h) = steam_cost1*abs(D1reboil(h))*(1/H_steam1)*31556926*(1/1000); R_2(h) = steam_cost1*abs(D2reboil(h))*(1/H_steam1)*31556926*(1/1000); R_3(h) = steam_cost2*abs(D3reboil(h))*(1/H_steam2)*31556926*(1/1000); dist_op_cost(h) = C_1(h)+C_2(h)+C_3(h)+R_1(h)+R_2(h)+R_3(h); Sep(h) = opcostcomp1(h) + dist_op_cost(h); C(h) = React(h)+Sep(h); %$/yr % Profit Before Tax = Revenue - Operating Costs P_bt(h) = (R(h) - C(h)); %$/yr %Equipment Cost Calculations %installed costs %Reactor Installed Cost

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%reactD = reactor diameter in ft, reactH = reactor height in ft %reactH = 6*reactD --> Vreact = 6*pi*reactD^3/4 reactD(h) = (4*Vreact(h)/(6*pi))^(4/3); reactH(h) = 6*reactD(h); %installed_react_cost(h) = (MS/280)*(101.9*reactD(h)^1.066*reactH(h)^0.802*(2.18+Fc_react)); installed_react_cost(h) = (MS/280)*(101.3*Aft(h)^0.65)*(Fh_heat+2.29); %Compressor Installed Cost installed_comp_cost1(h) = (MS/280)*517.5*bhp1(h)^0.82*(2.11+Fc_comp1); %Distillation Installed Cost %tray_h is the tray stack height in ft tray_h1(h) = Height_1(h)*3.2808399; tray_h2(h) = Height_2(h)*3.2808399; tray_h3(h) = Height_3(h)*3.2808399; %d is diameter in ft d1(h) = Diameter_1(h)*3.2808399; d2(h) = Diameter_2(h)*3.2808399; d3(h) = Diameter_3(h)*3.2808399; %Fs depends on tray spacing 24 in = 1, 18 in = 1.4, 12 in = 2.2 %tray spacing 0.6 meters Fs_dist1 = 1; Fs_dist2 = 1; Fs_dist3 = 1; %Sieve tray type Ft_dist = 0; %Tray material: CS = 0, SS = 1.7, Monel = 8.9 Fm_dist = 0; Fc_dist1 = Fs_dist1+Ft_dist+Fm_dist; Fc_dist2 = Fs_dist2+Ft_dist+Fm_dist; Fc_dist3 = Fs_dist3+Ft_dist+Fm_dist; installed_dist_cost_1(h) = (MS/280)*4.7*d1(h)^1.55*tray_h1(h)*Fc_dist1; installed_dist_cost_2(h) = (MS/280)*4.7*d2(h)^1.55*tray_h2(h)*Fc_dist2; installed_dist_cost_3(h) = (MS/280)*4.7*d3(h)^1.55*tray_h3(h)*Fc_dist3; %distillation heat exchangers %distillation column heat exchangers Fd_dr = 1.35; Fp_dr = 0; Fm_dr = 1; Fd_cr = 1; Fp_cr = 0; Fm_cr = 1; Fh_dr = (Fd_dr+Fp_dr)*Fm_dr; Fh_dc = (Fd_cr+Fp_cr)*Fm_cr; TlmC1(h) = ((Tfin-Tcout1(h))-(Tfout-Tcin1(h)))/log((Tfin-Tcout1(h))/(Tfout-Tcin1(h))); TlmC2(h) = ((Tfin-Tcout2(h))-(Tfout-Tcin2(h)))/log((Tfin-Tcout2(h))/(Tfout-Tcin2(h))); TlmC3(h) = ((Tfin-Tcout3(h))-(Tfout-Tcin3(h)))/log((Tfin-Tcout3(h))/(Tfout-Tcin3(h))); Ts1 = 242+273; Ts2 = 121+273; TavR1(h) = Ts1 - Treboil1(h); TavR2(h) = Ts1 - Treboil2(h); TavR3(h) = Ts2 - Treboil3(h); %steam and organic liquid Ur = 820;

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%W/(m^2*K) %steam and water Ur3 = 1430; %cooling water and organic Uc = 800; %J/s*(m^2 K s/J)*K^-1 [=] m^2 %m^2*(10.76 ft^2/m^2) [=] ft^2 Ar1(h) = D1reboil(h)/(Ur*TavR1(h))*10.76; Ar2(h) = D2reboil(h)/(Ur*TavR2(h))*10.76; Ar3(h) = D3reboil(h)/(Ur3*TavR3(h))*10.76; Ac1(h) = D1cond(h)/(Uc*TlmC1(h))*10.76; Ac2(h) = D2cond(h)/(Uc*TlmC2(h))*10.76; Ac3(h) = D3cond(h)/(Uc*TlmC3(h))*10.76; %area needs to be in feet squared installed_heat_cost_dr1(h)=(MS/280)*(101.3*Ar1(h)^(0.65))*(Fh_dr+2.29); installed_heat_cost_dr2(h)=(MS/280)*(101.3*Ar2(h)^(0.65))*(Fh_dr+2.29); installed_heat_cost_dr3(h)=(MS/280)*(101.3*Ar3(h)^(0.65))*(Fh_dr+2.29); installed_heat_cost_dc1(h)=(MS/280)*(101.3*Ac1(h)^(0.65))*(Fh_dc+2.29); installed_heat_cost_dc2(h)=(MS/280)*(101.3*Ac2(h)^(0.65))*(Fh_dc+2.29); installed_heat_cost_dc3(h)=(MS/280)*(101.3*Ac3(h)^(0.65))*(Fh_dc+2.29); installed_heat_cost_d(h) = installed_heat_cost_dr1(h)+installed_heat_cost_dr2(h)+installed_heat_cost_dr3(h)+installed_heat_cost_dc1(h)+installed_heat_cost_dc2(h)+installed_heat_cost_dc3(h); %MJ/yr*(10^6 J/MJ)*(Btu/1055 J)*(yr/8000 hr)*(10^-6) [=] 10^6 Btu/hr %QrBtu needs to be in 10^6 Btu/hr {1Btu = 1055 J} QrBtu = QrJ/(1055*8000); Fd_furn = 1.00; %process heater Fm_furn = 0; %carbon steel Fp_furn = 0; %Up to 500 psi Fc_furn = Fd_furn + Fm_furn + Fp_furn; installed_furn_cost(h) = (MS/280)*(5.52*10^3)*QrBtu^0.85*(1.27+Fc_furn); %Fixed Capital = Direct Costs + Indirect Costs %the cost of mixer is negligible sepfix(h) = installed_comp_cost1(h) + installed_dist_cost_1(h) + installed_dist_cost_2(h) + installed_dist_cost_3(h) +installed_heat_cost_d(h); OnsiteCosts(h) = real(installed_react_cost(h) + sepfix(h) + installed_furn_cost(h)); FC(h) = (OnsiteCosts(h))*1.25*1.45; %Working Capital Calculation Fa_wc = 2*Pea/1; %mol/yr %mol/yr*$/kg*g/mol [MW]*1kg/1000g wc_Fa_cost = Fa_wc*Cost_a*MWa*(1/1000); WC = (wc_Fa_cost)/4;

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%fraction of working capital = WC/FC = alpha_wc alpha_wc(h) = WC/FC(h); %sigma_a = used in TCI calculation sigma_a(h) = a_2*(1+CR)^2+a_1*(1+CR)+a0+alpha_wc(h)+alpha_su; aprime = TRC*sigma_b; bprime(h) = (0.1*TR*(1+alpha_su)-TRC*FR*sigma_a(h))*sigma+(TRC*(alpha_wc(h)+alpha_su)-sigma_a(h))*(1+FR)^-10; aa = aprime*(1+FR)^-PS; b(h) = bprime(h)*(1+FR)^-PS; c(h) = (aa*(1+alpha_wc(h)+alpha_su))/(sigma_a(h)*PL); d(h) = b(h)/(sigma_a(h)*PL); %TCI needed quantity according to packet TCI(h) = (FC(h)*sigma_a(h)); TI(h) = FC(h)*(1+alpha_wc(h)+alpha_su); % return on investment before taxes = ROI_bt ROI_bt(h) = (P_bt(h)/TI(h))*100; % net present value at year 0 = NPV0 (0 = start of operations period) NPV0(h) = (TRC*P_bt(h)*sigma_b-TRC*FR*TCI(h)*sigma+0.1*TR*(FC(h)*(1+alpha_su))*sigma+(TRC*FC(h)*(alpha_su+alpha_wc(h))-TCI(h))*(1+FR)^-10); NPVproj(h) = (NPV0(h)*(1+FR)^-PS); NPVpercent(h) = ((c(h)*ROI_bt(h)+d(h))); end end xa1 = [0.8 0.85 0.9 0.95]; hold on; fig1 = figure(1); set(fig1,'Color','white') plot(xa1, ROI_bt, plotcolor) xlabel('Reactor Conversion (0.05 = no recycle stream)') ylabel('Return on Investment (%/yr)') legend('T = 225 P = 1','T = 255 P = 1', 'T = 285 P = 1', 'T = 225 P = 5','T = 255 P = 5', 'T = 285 P = 5','T = 225 P = 10','T = 255 P = 10', 'T = 285 P = 10','Location','SouthEast') hold on; fig2 = figure(2); set(fig2,'Color','white') plot(xa1, TCI/10^6, plotcolor) xlabel('Reactor Conversion (0.05 = no recycle stream)') ylabel('Total Capitalize Investment (MM $)') legend('T = 225 P = 1','T = 255 P = 1', 'T = 285 P = 1', 'T = 225 P = 5','T = 255 P = 5', 'T = 285 P = 5','T = 225 P = 10','T = 255 P = 10', 'T = 285 P = 10','Location','SouthEast') hold on; fig3 = figure(3); set(fig3,'Color','white') plot(xa1, NPVproj/10^6, plotcolor) xlabel('Reactor Conversion (0.05 = no recycle stream)')

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ylabel('Net Present Value of Project (MM $)') legend('T = 225 P = 1','T = 255 P = 1', 'T = 285 P = 1', 'T = 225 P = 5','T = 255 P = 5', 'T = 285 P = 5','T = 225 P = 10','T = 255 P = 10', 'T = 285 P = 10','Location','SouthEast') hold on; fig4 = figure(4); set(fig4,'Color','white') plot(xa1, NPVpercent, plotcolor) xlabel('Reactor Conversion (0.05 = no recycle stream)') ylabel('Net Present Value Percent (%/yr)') legend('T = 225 P = 1','T = 255 P = 1', 'T = 285 P = 1', 'T = 225 P = 5','T = 255 P = 5', 'T = 285 P = 5','T = 225 P = 10','T = 255 P = 10', 'T = 285 P = 10','Location','SouthEast') hold on; fig5 = figure(5); set(fig5,'Color','white') plot(xa1, PeaT_kg/10^6, plotcolor) xlabel('Reactor Conversion (0.05 = no recycle stream)') ylabel('Amount of Product to Spec (MM kg/yr)') legend('T = 225 P = 1','T = 255 P = 1', 'T = 285 P = 1', 'T = 225 P = 5','T = 255 P = 5', 'T = 285 P = 5','T = 225 P = 10','T = 255 P = 10', 'T = 285 P = 10','Location','SouthEast') hold on; fig6 = figure(6); set(fig6,'Color','white') plot(xa1,Vreact, plotcolor) xlabel('Reactor Conversion (0.05 = no recycle stream)') ylabel('Reactor Volume (m^3)') legend('T = 225 P = 1','T = 255 P = 1', 'T = 285 P = 1', 'T = 225 P = 5','T = 255 P = 5', 'T = 285 P = 5','T = 225 P = 10','T = 255 P = 10', 'T = 285 P = 10','Location','SouthEast')

Obtaining Necessary Values %index desired is 3 which corresponds to a conversion of 90% r = 3; fprintf('Feed flow of ethanol: %i mol/hr\n', Fa/operating_hours) fprintf('Product flow of hydrogen: %i mol/hr\n', PsT(r)/operating_hours) fprintf('Product flow of ethyl acetate: %i mol/hr\n', PeaT(r)/operating_hours) fprintf('Product flow of diethyl ether: %i mol/hr\n\n', Pde/operating_hours) fprintf('Reactor:\n') fprintf('Energy Removed by Reaction: %i MW\n', Qr) fprintf('Reactor Diameter: %i m\n', reactD(r)) fprintf('Reactor Height: %i m\n', reactH(r)) fprintf('Reactor Volume: %i m^3\n\n', Vreact(r)) fprintf('Distillation 1:\n') fprintf('Diameter: %i m\n', Diameter_1(r)) fprintf('Height: %i m\n', Height_1(r))

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fprintf('Stages: %i \n', D1stages(r)) fprintf('Condenser Energy: %i J/s\n', D1cond(r)) fprintf('Condenser Surface Area: %i units\n', Ac1(r)) fprintf('Reboiler Energy: %i J/s\n', D1reboil(r)) fprintf('Reboiler Surface Area: %i units\n\n', Ar1(r)) fprintf('Distillation 2:\n') fprintf('Diameter: %i m\n', Diameter_2(r)) fprintf('Height: %i m\n', Height_2(r)) fprintf('Stages: %i \n', D2stages(r)) fprintf('Condenser Energy: %i J/s\n', D2cond(r)) fprintf('Condenser Surface Area: %i units\n', Ac2(r)) fprintf('Reboiler Energy: %i J/s\n', D2reboil(r)) fprintf('Reboiler Surface Area: %i units\n\n', Ar2(r)) fprintf('Distillation 3:\n') fprintf('Diameter: %i m\n', Diameter_3(r)) fprintf('Height: %i m\n', Height_3(r)) fprintf('Stages: %i \n', D3stages(r)) fprintf('Condenser Energy: %i J/s\n', D3cond(r)) fprintf('Condenser Surface Area: %i units\n', Ac3(r)) fprintf('Reboiler Energy: %i J/s\n', D3reboil(r)) fprintf('Reboiler Surface Area: %i units\n\n', Ar3(r)) fprintf('Reactor Operating Cost: %i MM $/yr\n', React(r)/10^6) fprintf('Distillation 1 Operating Cost: %i MM $/yr\n', (C_1(r)+R_1(r))/10^6) fprintf('Distillation 2 Operating Cost: %i MM $/yr\n', (C_2(r)+R_2(r))/10^6) fprintf('Distillation 3 Operating Cost: %i MM $/yr\n\n', (C_3(r)+R_3(r))/10^6) fprintf('Reactor Fixed Cost: %i MM $/yr\n', installed_react_cost(r)/10^6) fprintf('Distillation 1 Fixed Cost: %i MM $/yr\n', (installed_dist_cost_1(r)+installed_heat_cost_dr1(r)+installed_heat_cost_dc1(r))/10^6) fprintf('Distillation 2 Fixed Cost: %i MM $/yr\n', (installed_dist_cost_2(r)+installed_heat_cost_dr2(r)+installed_heat_cost_dc2(r))/10^6) fprintf('Distillation 3 Fixed Cost: %i MM $/yr\n\n', (installed_dist_cost_3(r)+installed_heat_cost_dr3(r)+installed_heat_cost_dc3(r))/10^6) fprintf('Econ Stuff:\n') fprintf('Profit Before Taxes: %f MM $/yr\n', P_bt(r)/10^6) fprintf('Fixed Capital: %f MM $\n', FC(r)/10^6) fprintf('alpha_wc: %f\n', alpha_wc(r)) fprintf('TCI: %f MM $\n', TCI(r)/10^6) fprintf('Return on Investment (TI): %f percent/yr\n', ROI_bt(r)) fprintf('NPV(0): %f MM $\n', NPV0(r)/10^6) fprintf('NPV percent: %f percent/yr\n', NPVpercent(r)) fprintf('NPV project: %f MM $\n', NPVproj(r)/10^6) fprintf('Operating Cost: %f MM $/yr\n', C(r)/10^6) fprintf('Cost of ethanol: %i MM $/yr\n', Pay(r)/10^6) fprintf('Value of ethyl acetate: %i MM $/yr\n', Sell(r)/10^6) fprintf('Value of hydrogen: %i MM $/yr\n\n', Extra(r)/10^6)

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Design2EconHYSYS clear;clc; %molecular weight in g/mol MWa = 46.07; MWea = 88.105; MWw = 18.02; MWs = 2.016; MWde = 74.12; %need to know values PeaT = 1.1392*10^9; PsT = 2.3983*10^9; %selectivity s = 10/11; %flow rates Pea_kg = 105.7*10^6; %kg/yr Pea = Pea_kg*1000*(1/88.105); %mol/yr [=] MM kg/yr*1000g/kg*mol/g Fa =2*Pea/s; Ps = 2*Pea; %mol/yr Pde = Pea*(1/s - 1); %mol/yr Pw = Pea*(1/s - 1); %mol/yr Fa_kg = Fa*MWa*(1/1000); %kg/yr Pde_kg = Pde*MWde*(1/1000); %kg/yr %costs Cost_pea = 1.30; %ethyl acetate $/kg Cost_a = 0.70; %ethanol $/kg Cost_s = 0.31; %hydrogen $/lb Cost_cat = 4; %catalyst $/kg electricity_cost = 0.06; %$/kWh %%%%%stuff from HYSYS %Process Furnances Fd_furn = 1.00; %process heater Fm_furn = 0; %carbon steel Fp_furn = 0; %Up to 500 psi Fc_furn = Fd_furn + Fm_furn + Fp_furn; %process furnace (heater 2) before distillation 1 QBR = 3.76269*10^6*1000*8000; %J/yr QBR_Btu = QBR/(1055*8000)*10^-6; %10^6 Btu/hr %cooler (heater 3) in between d2 and d3

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QBD = 915546*1000*8000; %J/yr %if don't have Acbet use the following with QBD Tfin = 4.4+273; Tout = 75.65+273; Tfout = 10+273; Tin = 163.5+273; TlmCafter = ((Tfin-Tout)-(Tfout-Tin))/log((Tfin-Tout)/(Tfout-Tin)); Uc = 800; %J/yr*m^2 K s/J*1/K*1 yr/31556926 s = m^2 %m^2*10.76 ft^2/m^2 = ft^2 Acbet = abs(QBD/(Uc*(TlmCafter))*10.76/31556926); %Reactor process furnace QrJ = 7.562*10^6 *1000*8000; %J/yr QrBtu = QrJ/(1055*8000)*10^-6; %10^6 Btu/hr %Reactor Vol = 41.5; %m^3 Vcat = Vol/2; r = 1/100; %m Aft = 2*Vol/r*10.76; %ft %cooler in between reactor and flash drum (heater 1) QAR = -2.38701*10^7*1000*8000; %energy removed = J/yr %coolant = refrigerated water coolant_cost = 5.7; %$/GJ %if don't have Acafter use the following with QAR Tfin = 4.4+273; T_drum = -18.15+273; Tfout = 10+273; reactT = 285+273; TlmCafter = ((Tfin-T_drum)-(Tfout-reactT))/log((Tfin-T_drum)/(Tfout-reactT)); Uc = 800; %J/yr*m^2 K s/J*1/K*1 yr/31556926 s = m^2 %m^2*10.76 ft^2/m^2 = ft^2 Acafter = abs(QAR)/(Uc*(TlmCafter))*10.76/31556926; %distillation columns %distillation column heat exchangers Fd_dr = 1.35; Fp_dr = 0; Fm_dr = 1; Fd_cr = 1; Fp_cr = 0; Fm_cr = 1; Fh_dr = (Fd_dr+Fp_dr)*Fm_dr; Fh_dc = (Fd_cr+Fp_cr)*Fm_cr; H_steam1 = 1755; H_steam2 = 2213; steam_cost1 = 6.74/1000; steam_cost2 = 2.38/1000; %distillation column

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%Fs depends on tray spacing 24 in = 1, 18 in = 1.4, 12 in = 2.2 %tray spacing 0.6 meters Fs_dist = 1; %Sieve tray type Ft_dist = 0; %Tray material: CS = 0, SS = 1.7, Monel = 8.9 Fm_dist = 0; Fc_dist1 = Fs_dist+Ft_dist+Fm_dist; Fc_dist2 = Fs_dist+Ft_dist+Fm_dist; Fc_dist3 = Fs_dist+Ft_dist+Fm_dist; D1stages = 35.1; D2stages = 30.8; D3stages = 25.8; %tray_h is the height in ft tray_h1 = 3*2 + 2*D1stages; tray_h2 = 3*2 + 2*D2stages; tray_h3 = 3*2 + 2*D3stages; %d is diameter in ft d1 = 1/6*tray_h1; d2 = 1/6*tray_h2; d3 = 1/6*tray_h3; D1reboil = 4.439*10^6; %J/s D2reboil = 3.4165*10^6; D3reboil = 5.2389*10^6; D1cond = -4.9115*10^6; D2cond = -4.8935*10^6; D3cond = -5.0817*10^6; %use if don't have surface area Tfin = 4.4+273; Tfout = 10+273; Tcout1 = 441; Tcin1 = 335; Tcout2 = 441; Tcin2 = 441; Tcout3 = 348; Tcin3 = 349; Treboil1 = 464; Treboil2 = 464; Treboil3 = 373; TlmC1 = ((Tfin-Tcout1)-(Tfout-Tcin1))/log((Tfin-Tcout1)/(Tfout-Tcin1)); TlmC2 = ((Tfin-Tcout2)-(Tfout-Tcin2))/log((Tfin-Tcout2)/(Tfout-Tcin2)); TlmC3 = ((Tfin-Tcout3)-(Tfout-Tcin3))/log((Tfin-Tcout3)/(Tfout-Tcin3)); Ts1 = 242+273; Ts2 = 121+273; TavR1 = Ts1 - Treboil1; TavR2 = Ts1 - Treboil2; TavR3 = Ts2 - Treboil3; %steam and organic liquid Ur = 820; %W/(m^2*K) %steam and water Ur3 = 1430; %cooling water and organic Uc = 800; %J/s*(m^2 K s/J)*K^-1 [=] m^2 %m^2*(10.76 ft^2/m^2) [=] ft^2 Ar1 = D1reboil/(Ur*TavR1)*10.76; Ar2 = D2reboil/(Ur*TavR2)*10.76; Ar3 = D3reboil/(Ur3*TavR3)*10.76; Ac1 = D1cond/(Uc*TlmC1)*10.76;

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Ac2 = D2cond/(Uc*TlmC2)*10.76; Ac3 = D3cond/(Uc*TlmC3)*10.76; CuOdensity = 6.31; CoOdensity = 6.44; Cr2O3density = 5.22; catdensity = 0.94*CuOdensity+0.05*CoOdensity+0.01*Cr2O3density; %marshal and swift factor for 210Q1 MS=1461.3; %Reactor %shell material use carbon steel Fm_react = 1.00; %P = 1 atm Fp = 1; P = 5 atm Fp = 1.05; P = 10 atm Fp = 1.15 %1 atm = 14.7 psi Fp_react = 1.00; Fc_react = Fm_react + Fp_react; Fm_heat=1.00; %material factor for CS/Ti Fp_heat=0; %pressure factor for up to 150psi Fd_heat=1.0; %type factor Floating head Fh_heat=(Fd_heat+Fp_heat)*Fm_heat; %F factor from gutheric Apendix E.2 of Douglas %fraction of start up capital = SU/FC = alpha_su alpha_su = 0.1; %fraction of fixed capital spent in start up years a_2 = 0; a_1 = 0.3; a0 = 0.7; %fraction of p_bt obtained in start up years = 3 startupyrs = 3; b1 = 0.5; b2 = 0.8; b3 = 1.0; %construction rate = CR, finance rate = FR, tax rate = TR, complementary tax rate = TRC CR = 0.1; FR = 0.08; TR = 0.5; TRC = 1-TR; %sigma_b = used in NPV0 calculation; note need to change if start up yr > 3 sigma_b = b1*(1+FR)^-1+b2*(1+FR)^-2+b3*(1+FR)^-3+(1+FR)^-3*((1-(1+FR)^-7)/FR); %sigma = used in NPV0 calculation; not for n = 10 years = lifetime of plant sigma = (1-(1+FR)^-10)/FR; % needed values PS = 2; % project start time PL = 12; % project lifetime %%%%%Economic Parameters %Costing % Revenue

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% R = EA val + H2 val - A cost (Water Sellable???) % R($/yr) = ($/kg)*(mol/yr)*(g/mol [MW])*(kg/1000g) [=] $/yr Sell = Cost_pea*PeaT*MWea*(1/1000); Pay = Cost_a*Fa*MWa*(1/1000); %$/lb*mol/yr*(g/mol [MW])*(0.0022lbs/g) Extra= Cost_s*PsT*MWs*0.0022; R = Sell-Pay+Extra; %Operating Cost Calculations % Operating Costs are the utility cost to run that piece of equipment % C = React + Sep System %cm^3*$/kg*g/cm^3*1kg/1000g*2/yr CatCost = Vcat*Cost_cat*catdensity*(1/1000)*2; %use Dowtherm over steam %using a process furnace to circulate the Dowtherm %operating cost is the cost fuel %J/yr*(1.62 $/MM Btu)*(1 Btu/1055 J)*(1 MM Btu/10^6 Btu) = $/yr Cost_Heat = abs(QrJ)*1.62/1055*10^-6; %Heater before reactor operating cost %J/yr*$/MM Btu*Btu/J*1 MM Btu/10^6 Btu Cost_Heat_2 = abs(QBR)*1.62/(1055 *10^6); %Cooler after reactor operating cost %J/yr*$/GJ*GJ/10^9J Cost_Cool = abs(QAR)*coolant_cost*10^-9; Cost_Cool_2 = QBD*coolant_cost*10^-9; React = CatCost + Cost_Heat + Cost_Heat_2 + Cost_Cool+Cost_Cool_2; %$/yr %distillation operating cost %Condenser Costs = coolant %$/GJ*J/s*(GJ/10^9 J)*(31556926s/yr) C_1 = coolant_cost*abs(D1cond)*(31556926)*10^-9; C_2 = coolant_cost*abs(D2cond)*(31556926)*10^-9; C_3 = coolant_cost*abs(D3cond)*(31556926)*10^-9; %Reboiler = heater %$/kg*kg/kJ*J/s*31556926s/yr*1kJ/1000J R_1 = steam_cost1*abs(D1reboil)*(1/H_steam1)*31556926*(1/1000); R_2 = steam_cost1*abs(D2reboil)*(1/H_steam1)*31556926*(1/1000); R_3 = steam_cost2*abs(D3reboil)*(1/H_steam2)*31556926*(1/1000); dist_op_cost = C_1+C_2+C_3+R_1+R_2+R_3; Sep = dist_op_cost; C = React+Sep; %$/yr % Profit Before Tax = Revenue - Operating Costs P_bt = (R - C); %$/yr %Equipment Cost Calculations %installed costs %Reactor Installed Cost installed_react_cost = (MS/280)*(101.3*Aft^0.65)*(Fh_heat+2.29);

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%Distillation Installed Cost %distillation 1 installed_dist_cost_1 = (MS/280)*4.7*d1^1.55*tray_h1*Fc_dist1; %distillation 2 installed_dist_cost_2 = (MS/280)*4.7*d2^1.55*tray_h2*Fc_dist2; %distillation 3 installed_dist_cost_3 = (MS/280)*4.7*d3^1.55*tray_h3*Fc_dist3; %area needs to be in feet squared %reboiler distillation 1 installed_heat_cost_dr1=(MS/280)*(101.3*Ar1^(0.65))*(Fh_dr+2.29); %reboiler distillation 2 installed_heat_cost_dr2=(MS/280)*(101.3*Ar2^(0.65))*(Fh_dr+2.29); %reboiler distillation 3 installed_heat_cost_dr3=(MS/280)*(101.3*Ar3^(0.65))*(Fh_dr+2.29); %condenser distillation 1 installed_heat_cost_dc1=(MS/280)*(101.3*Ac1^(0.65))*(Fh_dc+2.29); %condenser distillation 2 installed_heat_cost_dc2=(MS/280)*(101.3*Ac2^(0.65))*(Fh_dc+2.29); %condenser distillation 3 installed_heat_cost_dc3=(MS/280)*(101.3*Ac3^(0.65))*(Fh_dc+2.29); %total distillation heaters installed_heat_cost_d = installed_heat_cost_dr1 + installed_heat_cost_dr2 + installed_heat_cost_dr3 + installed_heat_cost_dc1 + installed_heat_cost_dc2 + installed_heat_cost_dc3; %reactor process furnance installed_furn1_cost = (MS/280)*(5.52*10^3)*QrBtu^0.85*(1.27+Fc_furn); %before reactor heater installed_furn2_cost = (MS/280)*(5.52*10^3)*QBR_Btu^0.85*(1.27+Fc_furn); %cooler in between reactor and flash drum installed_heatafter_cost = (MS/280)*(101.3*Acafter^0.65)*(Fh_dc+2.29); %total installed_heatbetween_cost = (MS/280)*(101.3*Acbet^0.65)*(Fh_dc+2.29); %total installed_furn_cost = installed_furn1_cost + installed_furn2_cost + installed_heatafter_cost + installed_heatbetween_cost; %Fixed Capital = Direct Costs + Indirect Costs %the cost of mixers and pumps are negligible sepfix = installed_dist_cost_1 + installed_dist_cost_2 + installed_dist_cost_3 + installed_heat_cost_d; OnsiteCosts = real(installed_react_cost + sepfix + installed_furn_cost); FC = (OnsiteCosts)*1.25*1.45; %Working Capital Calculation Fa_wc = 2*Pea/1; %mol/yr %mol/yr*$/kg*g/mol [MW]*1kg/1000g wc_Fa_cost = Fa_wc*Cost_a*MWa*(1/1000); WC = (wc_Fa_cost)/4;

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%fraction of working capital = WC/FC = alpha_wc alpha_wc = WC/FC; %sigma_a = used in TCI calculation sigma_a = a_2*(1+CR)^2+a_1*(1+CR)+a0+alpha_wc+alpha_su; aprime = TRC*sigma_b; bprime = (0.1*TR*(1+alpha_su)-TRC*FR*sigma_a)*sigma+(TRC*(alpha_wc+alpha_su)-sigma_a)*(1+FR)^-10; aa = aprime*(1+FR)^-PS; b = bprime*(1+FR)^-PS; c = (aa*(1+alpha_wc+alpha_su))/(sigma_a*PL); d = b/(sigma_a*PL); %TCI needed quantity according to packet TCI = (FC*sigma_a); TI = FC*(1+alpha_wc+alpha_su); % return on investment before taxes = ROI_bt ROI_bt = (P_bt/TI)*100; % net present value at year 0 = NPV0 (0 = start of operations period) NPV0 = (TRC*P_bt*sigma_b-TRC*FR*TCI*sigma+0.1*TR*(FC*(1+alpha_su))*sigma+(TRC*FC*(alpha_su+alpha_wc)-TCI)*(1+FR)^-10); NPVproj = (NPV0*(1+FR)^-PS); NPVpercent = ((c*ROI_bt+d)); fprintf('Econ Stuff:\n') fprintf('Profit Before Taxes: %f MM $/yr\n', P_bt/10^6) fprintf('Fixed Capital: %f MM $\n', FC/10^6) fprintf('alpha_wc: %f\n', alpha_wc) fprintf('TCI: %f MM $\n', TCI/10^6) fprintf('Return on Investment (TI): %f percent/yr\n', ROI_bt) fprintf('NPV(0): %f MM $\n', NPV0/10^6) fprintf('NPV percent: %f percent/yr\n', NPVpercent) fprintf('NPV project: %f MM $\n', NPVproj/10^6) fprintf('Operating Cost: %f MM $/yr\n', C/10^6) fprintf('Cost of ethanol: %i MM $/yr\n', Pay/10^6) fprintf('Value of ethyl acetate: %i MM $/yr\n', Sell/10^6) fprintf('Value of hydrogen: %i MM $/yr\n\n', Extra/10^6) fprintf('Reactor Operating Cost: %i MM $/yr\n', (CatCost + Cost_Heat)/10^6) fprintf('Heat Exchanger 1: %i MM $/yr\n', Cost_Cool/10^6) fprintf('Heat Exchanger 2: %i MM $/yr\n', Cost_Heat_2/10^6) fprintf('Heat Exchanger 3: %i MM $/yr\n', Cost_Cool_2/10^6) fprintf('Distillation 1 Operating Cost: %i MM $/yr\n', (C_1+R_1)/10^6) fprintf('Distillation 2 Operating Cost: %i MM $/yr\n', (C_2+R_2)/10^6) fprintf('Distillation 3 Operating Cost: %i MM $/yr\n\n', (C_3+R_3)/10^6) fprintf('Reactor Fixed Cost: %i MM $\n', (installed_react_cost+installed_furn1_cost)/10^6) fprintf('Heat Exchanger 1: %i MM $\n', installed_heatafter_cost/10^6) fprintf('Heat Exchanger 2: %i MM $\n', installed_furn2_cost/10^6)

60

fprintf('Heat Exchanger 3: %i MM $\n', installed_heatbetween_cost/10^6) fprintf('Distillation 1 Fixed Cost: %i MM $\n', (installed_dist_cost_1+installed_heat_cost_dr1+installed_heat_cost_dc1)/10^6) fprintf('Distillation 2 Fixed Cost: %i MM $\n', (installed_dist_cost_2+installed_heat_cost_dr2+installed_heat_cost_dc2)/10^6) fprintf('Distillation 3 Fixed Cost: %i MM $\n\n', (installed_dist_cost_3+installed_heat_cost_dr3+installed_heat_cost_dc3)/10^6)

Bibliography

Bernard Silberstein, H. B. (1969). Kinetics of Homogeneously Catalyzed Gas-Liquid Reactions.

Chlorination of Benzene with Stannic Chloride Catalyst. Industrial & Engineering

Chemistry Fundamentals, 8 (3), 365-374 .

Doherty, M. (2011). Production of Ethyl Acetate Solvent via Green Technology. Santa Barbara.

Douglas, J. M. (2011). Conceptual design of chemical processes. New York: Mcgraw Hill.

M&S. (2010, January ). Economic Indicators. Chemical Engineering, p. 64.

Mellichamp, D. A. (2011). Evaluating Plant Profitability in a Risk-Return Context. Santa Barbara:

University of California.

Ethyl Acetate Design Project - ideangineideangine.com/Ethyl_Acetate_Design_Project.pdf · Ethyl Acetate Design Project ... distillation column that removed water in the process, ...

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