A s s i gn me n t 10 D e s i gn R e p or t - G r ou p 1 ...johanv/Final10.490.pdf · A s s i gn me n t 10 D e s i gn R e p or t - G r ou p 1 Process Control Table 1:Detailed Description

10.490 ICE I - Continuous Process Design Fall 2019 Assignment 10 Design Report - Group 1

Plan Z Corporation Business Development TO: P.I. Barton and Y. Román FROM: Connor Chung, Sebastian Esquivel, Ziad Mansour, Johan Villanueva SUBJECT: Final Design for Ethylene Hydration Process Our team investigated the direct hydration of ethylene into ethanol using a solid phosphoric acid coated catalyst. We considered several factors when evaluating the feasibility and desirability of this process, including the overall US market for ethanol, safety concerns with this process, and the economic drivers. The market for ethanol is vast and diverse, as ethanol is a valuable component in many fuels, chemical products, and consumer products. Potential customers and business partners include research institutions and large corporations within the US who would be interested in purchasing ethanol for a variety of purposes. Annual production of ethanol in the US was approximately 15.5 billion gallons/year. Our projected annual revenue is $37.8 MM, and our projected annual gross profit is $2.6 MM. Our projected total capital costs are $55.3 MM. Though we are operating at a profit, the NPV of our process after 25 years is -$37.2 MM. This is largely due to the fact that we pay a 21% tax rate, and our future cash flows are discounted by 10%. There are several areas that could be improved, which could increase the NPV of the project. One large component of our costs is our total utility cost. Within utility costs, oil fuel, medium pressure steam, and electricity represent over 90% of our utility costs. Our team has worked to reduce these costs, but further optimization could yield a more optimistic NPV projection. Another large cost is our capital costs. For the cost projections in this project, we assumed that the plant would be built and operated within the US. We believe that moving this production plant overseas to a location with cheaper land and materials costs would present an opportunity for this process to be more profitable. Due the very negative net present value of this process, we highly recommend that Plan Z Corporation does not follow through with this process in its current state within the US. Best Wishes, PlanZ Corporation

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Process Flow Diagram and MB&EB Note: the numbering of the streams and naming of units blocks pertain to only the PFD.

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Piping and Instrument Diagram Note: the numbering of the streams and naming of units blocks pertain to only the PFD.

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Process Control Table 1:Detailed Description of the Process Controls:

# Controlled Variable Measured Variable Purpose

1F Flow of Organic (Ethylene) feed

Flow rate of organic feed stream

Ensure that the desired inlet flow rate is achieved and that the ethylene stock is not degraded

2T Inlet chilled water flow rate Outlet temperature of H1 exchanger

Ensure desired stream temperature is reached

3T Inlet cooled water flow rate Outlet temperature of H2 exchanger


4T Inlet LP steam flow rate Outlet temperature of H3 exchanger


5T Inlet oil fuel rate Outlet temperature of H4 furnace


6F Flow of Water (Aqueous) feed Flow rate of water feed stream

Ensure that the desired inlet flow rate is achieved and that the ethylene stock is not degraded



8P Flow rate of the bottoms stream leaving the Flash

Outlet pressure of the bottoms stream leaving the Flash

Ensure the stream is not dangerously over-pressurized

9T MP steam entering the reboiler of DISTILL1

Temperature of the bottoms stream from DISTILL1

Ensure that the DISTILL1 column is operating at desired specs

10T Cooled water entering the condenser of DISTILL2

Temperature of the distillate stream from DISTILL2


11T MP steam entering the reboiler of DISTILL2

Temperature of the bottoms stream from DISTILL2




13T Cooled water entering the condenser of DISTILL1

Temperature of the distillate stream from DISTILL1


14F Purge fraction in the purge splitter

Purge flow rate exiting the system

Ensure that the desired purge rate is being achieved

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Economic Estimates Overall Equipment/ISBL Costs Below is a table containing all of the unit operations of our process, as well as the correlation factors that were considered when estimating the equipment cost. We used Towler correlations to estimate the cost of each equipment piece except the reactors. The details of how we estimated the cost of the reactors is included in a separate section below for reactors. The total estimated equipment cost was $11.040 MM, and the total estimated ISBL cost was $28.636 MM.

Table 2: Breakdown of capital cost

Unit Op Size S Units S Lower S Upper a b n

# Units 2010 Cost

Installation Factor ISBL Cost

Reactor - - - - - - - 4 $1,054,321.02 2.5 $2,635,802.55

Distill 1

pressure vessel 11399.1 kg 120 250000 17400 79 0.85 1 $335,142.88 4 $956,814.34

packaging 50.515 m^3 0 7600 1 1 $383,914.00 4 $1,535,656.00

trays 1 3.10485 diameter, m 0.5 5 130 440 1.8 6 $21,069.77 4 $84,279.07

trays 2 2.49979 diameter, m 0.5 5 130 440 1.8 9 $21,772.58 4 $87,090.31

Distill 2

pressure vessel 17465 kg 120 250000 17400 79 0.85 1 $456,617.08 4 $1,344,659.30

packaging 53.882 m^3 0 7600 1 1 $409,503.20 4 $1,638,012.80

trays 1 2.67774 diameter, m 0.5 5 130 440 1.8 13 $35,370.55 4 $141,482.20

trays 2 1.91448 diameter, m 0.5 5 130 440 1.8 13 $20,101.43 4 $80,405.72

Flash 1769.84 shell mass, kg 160 250000 11600 34 0.85 1 $31,198.50 4 $124,794.00

Compressors

C1 (kW) 885.072

driver power, kW 93 16800 260000 2700 0.75 1 $698,124.68 2.5 $1,745,311.69

C2 (kW) 4845.91

driver power, kW 93 16800 260000 2700 0.75 1 $1,828,177.69 2.5 $4,570,444.24

C3 (kW) 1242.23

driver power, kW 93 16800 260000 2700 0.75 1 $824,956.84 2.5 $2,062,392.11

Pumps

P1 (kW) 243.296 power, kW 1 2500 -1100 2100 0.6 1 $55,641.43 4 $222,565.72

P2 (kW) 6.33366 power, kW 1 2500 -1100 2100 0.6 1 $5,256.41 4 $21,025.65

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Unit Op Size S Units S Lower S Upper a b n

# Units 2010 Cost

Installation Factor ISBL Cost

Heaters

H1 (double pipe exchanger) 7.5397 area, m^2 1 80 1900 2500 1 1 $20,749.25 3.5 $72,622.39



H4 (box furnace) 16.484 duty, MW 0.2 60 80000 109000 0.8 1 $1,105,813.94 2 $2,211,627.88

H5 (box furnace) 3.8249 duty, MW 0.2 60 80000 109000 0.8 1 $398,805.62 2 $797,611.24

H6/H8 (U-tube exchanger) 1083.56 area, m^2 10 1000 28000 54 1.2 1 $264,710.96 3.5 $926,488.36

H7 (box furnace) 36.962 duty, MW 0.2 60 80000 109000 0.8 1 $2,037,169.81 2 $4,074,339.62

Mixers (static)

COMBO (50 mixers in series) 27.793 liters/s 1 50 570 1170 0.4 50 $249,671.52 2.5 $624,178.80

H2OMIX (50 mixers in series) 21.0932 liters/s 1 50 570 1170 0.4 50 $226,567.33 2.5 $566,418.33

ALLMIX (50 mixers in series) 43.902 liters/s 1 50 570 1170 0.4 50 $294,051.20 2.5 $735,128.00

RXN-MIX (50 mixers in series) 44.699 liters/s 1 50 570 1170 0.4 50 $295,969.95 2.5 $739,924.86

REC-MIX (50 mixers in series) 26.345 liters/s 1 50 570 1170 0.4 50 $244,989.31 2.5 $612,473.29

Total Cost $11.040 MM $28.636 MM

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Utilities Below is a table outlining all of the unit blocks that used utility, which utility was used, and how much of the utility was used.

Table 3: Breakdown of the amount of utilities used per block

Block Description Utility Units Amount H1 heat exchanger chilled water kg/hr 1936.49 H2 heat exchanger cooled water kg/hr 25059.9 H3 heat exchanger low pressure steam kg/hr 2464.6 H4 furnace oil cal/sec 3.94E+06 H5 furnace oil cal/sec 914182 H7 furnace oil cal/sec 8.83E+06

DISTILL1 distillation column cooled water kg/hr 1.31E+06 DISTILL1 distillation column medium pressure steam kg/hr 46987.3 DISTILL2 distillation column cooled water kg/hr 646288 DISTILL2 distillation column medium pressure steam kg/hr 14333.3

P1 pump electricity kW 11.1523 P2 pump electricity kW 6.33366 C1 condenser electricity kW 885.072 C2 condenser electricity kW 4845.91 C3 condenser electricity kW 1242.23

Note: For the distillation columns, cooled water was used in the partial condensers whereas medium pressure steam was used in the reboilers. Below is a table that tabulates the total utility cost of our process. The largest components of our total utility costs are Electric, Fuel Fire, and MP (medium pressure) Steam, which together comprise approximately 94% of the total utility cost. These represent potential areas to focus on improving in our system, as decreasing our total utility cost would increase our yearly operating profit.

Table 4: Overall cost of each utility

Units Units/Unit product Units/hr

Price $/unit $MM/yr

Electric kWh 0.0 6,989.96 0.11 6.267 HP Steam Mg 0.0 0.00 14.50 0.000 MP Steam Mg 0.0 61.33 10.90 5.448 LP Steam Mg 0.0 2.46 8.42 0.169

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Boiler Feed Mg 0.0 0.00 0.00 0.000 Condensate Mg 0.0 0.00 0.00 0.000

Cooling Water Mg 0.0 1,978.73 0.05 0.806 Chilled Water Mg 0.0 1.94 6.05 0.095

Fuel Fired GJ 0.0 106.29 6.90 5.977 Total Utilities (UTS) 18.763

Reactors For the 4 reactor blocks, we essentially calculated the cost of the physical reactor shell, assuming that it is made out of stainless steel. The calculations can be seen below. We assumed that the thickness of the reactor was 3.6 inches, and that the price of stainless steel is $1.58/pound. The total cost of one reactor was estimated to be $263,580.

Table 5: Capital cost estimate for a reactor

Outer diameter (m) 2.5908 Inner diameter (m) 2.4079 Length (m) 13.716 Shell Volume (m^3) 9.8479 Density of Steel (kg/m^3) 7700 Mass of Steel (lbs) 166823 Price of Steel ($/lb) 1.58 Total Price ($) $263,580

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Heat Exchangers For heat exchangers (units H1, H2, H3 H6/H8), the unit for size was m^2, or area of the exchanger. In other words, surface area is the dimension required to use the cost estimate correlation found in Towler. However, Aspen did not provide any physical dimensions for the heater blocks, so an area had to be calculated by other means. Using the equation below, an area was calculated based off of the heat duty, Q, which was provided by Aspen, as well as the inlet/outlet temperatures of the process stream and the utility stream, encapsulated in the LMTD term. Q = U*A*LMTD Q = heat duty (W) U = overall heat transfer coefficient (W/(m^2*K)) A = area (m^2)

LMTD = logarithmic average of the temperature difference, defined as: The overall heat transfer coefficient was chosen based on values found online that describe heat exchange occurring for water and oil flowing in tubes. This value was 700 W/(m^2K). The LMTD was calculated using the inlet/outlet temperatures of the process stream and the utility stream, based on the equation shown above. Once we had Q, U, and LMTD, an area could be calculated for each heat exchanger, thus allowing us to make a cost estimate based on the correlations in Towler. These calculations are shown in the table below, which contains all our cost estimate calculations. For the furnaces (H4, H5, H8), the cost estimate correlations in Towler require the heat duty, which is provided by Aspen. The table below shows the conversion from calorie/second to megawatt (MW), which is the required unit for the correlations. H6 and H8 heater blocks were modeled together as a heat exchanger. Essentially, we decided to model the heat expelled from H8 as the heat inputted in H6. Prior to this change, we had an actual heat exchanger block in our flowsheet, but our process would not run without mass balance errors. This was a creative way to implement a heat exchanger without actually using a heat exchanger block in Aspen.

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Table 6:Parameters required to calculate Heat Exchangers Areas

H1 H2 H3 H4 H5 H6/H8 H7

Description cooling

exchanger cooling

exchanger heating

exchanger furnace furnace exchanger furnace Q (cal/sec) -8058.99 103982 3.58E+05 3.94E+06 9.14E+05 8.07E+06 8.83E+06

Q (W) -33718.8 435060 1.50E+06 1.65E+07 3.82E+06 3.38E+07 3.70E+07 U (W/m^2*K) 700 700 700 - - 700 - T_hot_in (C) 25 250 125 - - 315 - T_hot_out (C) 13 140 124 - - 103 - T_cold_in (C) 5 25 84.27 - - 100 -

T_cold_out (C) 19.99 39.99 100 - - 128 -

LMTD 6.3888 157.77 31.798 - - 44.525 - A (m^2) 7.5397 3.9394 6.74E+01 - - 1.08E+03 -

Heat Duty (MW) - - - 1.65E+01 3.82E+00 - 3.70E+01

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Pressure Vessels (Flash) The cost estimate correlation for pressure vessels requires a shell mass in kilograms. Aspen did not provide any dimensions of our flash unit, so we had to calculate this on our own, using information found in the textbook and online. First, it was critical to find the inner diameter of the flash vessel. Assuming that the vessel was a vertical cylinder, we calculated a maximum vapor velocity based on the densities of the liquid and vapor phases, as shown in the equation below. V_max = k*sqrt( (rho_L - rho_V) / rho_V) Based on literature, k was determined to be 0.107 m/s. We assumed that the vapor stream was predominantly ethylene, so rho_V was 1.18 kg/m^3, or the density of ethylene. We then calculated the density of the liquid stream based on the mass fractions of water and ethylene, and found that rho_L was approximately 608 kg/m^3. Using these values, we calculated a V_max of 2.426 m/s. The area was calculated by the equation: A [m^2] = vapor flow rate [m^3/s] / vapor velocity [m/s] A vapor flow rate of 0.7367 m^3/s was provided by Aspen. We decided that the velocity of the vapor stream was not going to be the maximum velocity, because we assumed that our system was not performing at that level. We decided to use a vapor velocity of 1 m/s, to simplify our math as well as make our flash more realistic. Based on this, our area was calculated to be 0.7367 m^2. D [m] = (4A / pi)^0.5 The diameter of the cylindrical flash drum was calculated as shown above, and it was found to be 0.9685 m. This measurement physically seems realistic and feasible. Upon checking on our work, we realized that a higher vapor velocity would have yielded a much smaller flash unit, to the point where we would not trust the feasibility of such a piece of equipment. Since the measurement of area is the cross-sectional area of the inside of the flash, the resulting diameter is the inner diameter. W = rho*pi*d*t*L Now that we have an inner diameter measurement, we are able to calculate a shell weight, using the equation shown above. We assumed that the thickness of the shell was 26mm, or 0.026m. This is based off literature search. We also assumed that the flash was made of steel, which has a density (rho) of 7700 kg/m^3. The length of the flash was calculated using a 3:1 length:diameter ratio. Thus, the weight of the shell was calculated to be 1769.84 kg. Similar calculations were performed to estimate the cost of the pressure vessels in our two distillation columns.

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Key Insights As we can see from the results in the flow sheet we are losing a lot of money. In fact, our NPV is of -$37.7 MM after 25 years. We have considerably lowered our utility costs, from ~$28MM/year to ~$18MM/year. Our payback period is infinite so our analysis suggest that we will never pay it back. This is largely due to the fact that we pay a 21% tax rate, and our future cash flows are discounted by 10%. However, we were able to find some areas that can make our system more profitable ( such as reducing utility costs …). From the economic analysis, it seems clear that the utility costs need to be significantly reduced in order to break even or become profitable. We have thought of a couple of ways in which we can improve the profitability of the system.

1) Reduce some of the utility costs by keeping the streams going into the flash at 70 bars instead of 50 bars.

2) Use some Heat exchangers instead of heater that rely on fuel fire. 3) Operate the reaction at a lower temperature. 4) Increase the profitability of the system by moving the plant to Lebanon. In fact, we have

found proxies for the selling price of ethanol, ethylene as well as the other utilities and the system is much more profitable if we operate it there (Refer to Cost Sheet if you want more details). Some of the assumptions (Tax, overhead salary…) were adapted to reflect the standard and laws of Lebanon. We found proxies for the utility costs as well as ethylene and ethanol from a document written in arabic from Tripoli Oil Installation. As we can see from the table, it is clear that the plant will be more profitable in a country such as Lebanon.

Table 7: Utilities, Ethanol and Ethylene costs in Lebanon

Units Price

$/unit Electric kWh 0.083 HP Steam Mg 12.23 MP Steam Mg 7.86 LP Steam Mg 5.48 Ethylene kg 0.320 Ethanol kg 0.9720 Cooling Water Mg 0.02 Chilled Water Mg 4.05 Fuel Fired GJ 6.74

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Safety and Environmental Summary Throughout our Preliminary Hazard Analysis (PHA), our team was able to highlight the most pertinent underlying potential dangers to our operation. In handling a large inventory of highly flammable material, measures should be taken to ensure that any possible leaks can be quickly detected and isolated. Leaks of ethanol or ethylene leads to contamination that can put the plant, operators, and surroundings at serious risk for explosions. Using phosphoric acid as a coating for the solid catalyst in the packed bed reactor may result in flammable, toxic fumes upon combustion. Our team recommends that greater investments not be overlooked when considering piping and transportation equipment to minimize potential leaks, contaminations, and explosions. Potential equipment ruptures also pose significant dangers to our operations. Overpressurization in pumps or compressors can result in the release of flammable liquids, combustible dusts, or even thermal radiation exposure. Poor equipment maintenance may result corrosion and equipment failure. Therefore, it is very important to invest in equipment that limits potential ruptures or failure. Operating teams should ensure that pressure levels are maintained at safe operating conditions. Proper ventilation systems should be incorporated in the system, and proper lockout/tagout procedures should be taken when operating any sources of hazardous substances. Another important consideration for the success of any chemical plant operation is the location of the plant itself. Potential hazards may result from establishing chemical plants in areas that are prone to natural disasters or any sort of environmental calamities, such as earthquakes, tsunamis, hurricanes, etc. Avoiding these potential incidents saves significant amounts from any extreme safety measures, potential capital losses, or even reconstruction. The chemical plant should be located in an area not prone to these potential natural disasters and with viable transportation options for material and equipment. Moreover, in order to reduce the transportation cost, we recommend that the plant is situated in an area that has ethylene produced. Our recommended plant location is Beirut, Lebanon. One of the most critical tasks in any chemical process is designing a process which poses minimal harm to the environment. Throughout our work over the past three months, our team has generated many working iterations of a successful, high yielding ethanol plant. After accomplishing our target numbers, we were determined to minimize our environmental impact through the wastes produced by our operation. By incorporating a recycle stream to the bottoms of our second distillation column, we were able to drastically lower operating costs for treating our waste water by looping it back into our system. Currently, the only waste that exists comes from the purge stream of the ethylene recycle stream. As for waste treatment, this purge constitutes only about 7% of that of the ethanol product stream in terms of mass flow rates. This makes treatment options very manageable.

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Aspen Plus Model The following section provides a detailed description of the Aspen flowsheet “PSET10V6.bkp” as well as a detailed description of the assumptions used. The text will often refer to streams and unit operations as they are named in the aspen flow sheet. Property Method The property method chosen for our process was NRTL. In order to choose the model, we predicted the T-xy diagram of the non ideal mixture of ethanol and water. First we used the ideal gas model but quickly realized that it is a pretty poor model because it assumes ethanol and water behave “ideally” or in other words that there are no intermolecular interaction between ethanol and water. Well as we know, there exist some intermolecular interactions (Hydrogen bonding, dipole-dipole) between ethanol and water. Hence, the ideal gas model does not capture the thermodynamics of the mixture appropriately. It is also clear that there is a big difference between the Tx-y diagram given by the ideal gas law and the ones given by UNIFAC, NRTL, and experimental data from NSIT. Then we looked at the prediction between UNIFAC and the NRTL models. These models were found to be more appropriate. In fact, these models give extremely similar results to each other and were similar to the experimental data from NSIT. In fact, these models take into account the intermolecular forces (H-bond) between water and ethanol and the positive azeotrope. The main difference between the models comes from the estimation of the activity coefficient. In order to choose the most suitable model, we checked litterature and saw that in the case of ethanol and water, NRTL has the smallest root mean squared error between predicted and experimental values. It also seems that NRTL predicts better the azeotropic portion of the T-xy diagram.

Description of the Process and Key Assumptions

We have an ethylene feed coming in at 25 ̊C and 1.5 bar. We start by cooling the feed using a cooler H1 to 13 ̊C. Then we compress the ethylene feed to 37 bar in order for it to match the pressure of the recycling stream. PM3 and the Tear stream will mix in the Combo and then this stream will go through a cooler, compressor and finally a furnace. These 3 units of operations will bring the stream to the desired reaction conditions of 300 ̊C and 70 bar. It is important to note that we use two coolers in this phase because the streams cannot be at a temperature higher than 250 ̊ in the streams from a safety standpoint. If you omit the coolers, then the outlets of the compressors will be at unfeasible conditions above 300 ̊C, hence it was necessary to add coolers.

On the other hand, a feed water comes in at 20 ̊C and 1.5 barg. This feed will be pressurized by a pump to 70 bar and then heated to 300 ̊C by a furnace. This feed of water will be mixed in H2OMIX with a recycling water stream coming from the bottom of the second distillation columns.

The water stream and ethylene stream will be mixed and then split in RXNSPLIT. The split fraction is 0.5. Then the streams will go into the reactor. We used 4 adiabatic packed bed reactors that have a length of 45 feet and a diameter of 8.5 ft. The packing is a solid catalyst with

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a phosphoric acid coating and has the following properties: a bed voidage of 0.4; particle density of 2500 kg/cum; the particle diameter of 1cm; shape factor of 1. The configuration and the number of reactors were calculated from the basis of achieving a 4.78% conversion. Achieving the target required the usage of a massive reactor with unrealistic dimension (reactor cannot fit in a container to be transported). The pressure drop through the reactor was calculated using an Ergun frictional correlation with a scaling factor of 1 and a roughness 4.572 e-05 meter. We mounted 2 sets of 2 reactors in series in parallel. This configuration ensured that the reactor size was feasible. Putting reactor in series reduces the length of a reactor whereas putting reactors in parallel reduces the diameter of the reactor. The stoichiometry of the reaction is 1 molecule of ethylene and 1 molecule of water react to form 1 molecule of ethanol. The reaction class is PowerLaw and the reaction occurs in the vapor phase. The driving force is the partial pressure measured in bar and the rate basis is Cat(wt). Moreover, the values of the different kinetic factors k=60, n=0, E=30 kcal/mol and the reaction rate units was set to kmol/kg-s.

MIXED1 (at the outlet of the reactor) goes through a valve that reduces the pressure of the stream from 70 bar to 50.5 bar. Then the stream will be cooled in a heat exchanger before heading to an adiabatic flash operating at 102.5 ̊C and 50 bar. An ethylene stream will be taken out whereas the rest of the components will go through a series of 2 Distillation columns. In these distillation columns, we had to optimize for 2 main parameters. The first parameter is the number of trays in the column which represents the capital cost and the second parameter is the reflux ratio which represents the operating cost.

We started using Group methods (GM) in order to estimate the number of trays needed as in an initial guess. Basically Group methods uses approximate calculations to relate the outlet stream properties to the inlet stream specifications and number of equilibrium trays. These approximation procedures are called group methods because they provide only an overall treatment of the stages in the cascade without considering detailed changes in the temperature and composition of individual stages. However, they are much easier to solve because they involve fewer variables and constraints. Although we know that Group methods is not ideal to estimate the number of trays in the case of the distillation, we still used it because of the simplicity of the calculation and the fact that it would give us a good initial guess from where we can do a sensitivity analysis in order to determine what is optimal for our model. In order to determine the optimal reflux ratio, we decided to use an adapted version of the cuckoo optimization algorithm. This method gave me some educated guess which we had to change in my model in order to obtain my desired separation. We also ran the same analysis on aspen using the DSTWU columns, and drew the graph of Reflux ratio vs number of trays and obtained similar numbers to the ones described above.

It is important to note that when we ran both of the optimization problems, we set the recovery rate of ethanol as 0.99 (~approaching 1) for the bottoms in the first column and for the distillate for the second column. After getting initial guesses, we adjusted the values in order to achieve the required specification (Mass fraction of ethanol is 92.4% and the mass fraction of organics in the ethanol stream is less than 0.0001%, Mass flow of commercial ethanol needs to be greater than 10,000 kg/hr). The first distillation column has 16 stages and a reflux ratio of 1.3 whereas the second distillation column has 26 stages and a reflux ratio of 3.94. Moreover, the feeds were

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placed on stages 2 and 13 for distillation columns 1 and 2 respectively. Finally, using mass balances we specified the distillate to feed ratio to be 0.328 (by mass) for the second column and 0.0972 for the second column. Both distillation have a kettle reboiler as well as a partial condenser.

The top of the first distillation column (Organics Stream) will be mixed with the Ethylene stream. Then this stream will be split into a purge stream and a recycle stream (TEAR stream). The purge fraction is of 0.0045. The split fraction was set in order to ensure that the amount of ethylene lost in the purge stream is less than 10% than the amount of ethylene fed through FEEDMIX (in our design the value of 8.22%). It is important to add an ethylene recycling stream since the conversion of ethylene is less than 5% and it would be a huge waste and a massive operating cost to not recycle it.

The bottoms stream of the second distillation columns (Bottoms) is 99.87% water and 0.13% Ethanol (in terms of mass fraction) and basically contains no other organic component. Since the bottoms is purely aqueous and the amount of water needed for the reaction is in the order of 10^5 kg/hr. We decided that from a cost effective standpoint it would be useful to recycle water. We decided to not include a purge stream for Bottoms because the stream was mainly aqueous. Recycling water not only reduces feed and waste costs but it also helps in reducing utility cost. In fact, water comes out of the second distillation column at 99 ̊C whereas purchased feed water is at 20 ̊C. The Bottoms stream will be treated before getting mixed with the FEED-H2O. In fact, the Bottoms stream liquid and need to be vaporized. Hence this stream, will go through a pump and then will go through a counter-current heat exchanger. On the Aspen file, the counter-current heat exchanger is represented by a cooler H8 that will cool the stream coming out of the reactor from 315 ̊C to 103 ̊C and a heater H6 that will vaporize 50% of the liquid of PM10 (Bottoms stream) and raise the temperature to 127 ̊C. We were not able to get the flow sheet to converge on Aspen with a heat exchanger so we decided to represent the heat exchanger by a set of 2 “Heater” Aspen blocks. From the description we can note that recycling the Bottoms stream reduces the utility cost associated with cooling MIXED2 as well as the cost needed to vaporize 50% of the aqueous stream. The aqueous stream will be vaporized and heated in a furnace and then pressurized in a compressor. The stream will be recycled afterwards.

An important note is the fact that the flow sheet contains many coolers that might not make total sense for the readers. The reason we used many coolers is to make sure the temperature of the streams is not above 250-275 ̊C (except the outlets and inlets of the reactors). This is a safety measure that is necessary to be taken.

Other Assumptions Used in the design

- A pressure drop of 0.5 bar is set across each heater. - The compressors are all isentropic and their efficiencies is set to 0.80. - The pumps efficiencies are set to 0.80. - The tear stream is PM15 and its composition, temperature and pressure are specified. - The convergence method used for tears is Wegstein and the tolerance is set to 0.0001.

List of Utilities used by each block

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- Chilled Water : H1 - Cooled Water : H2, DISTILL1 (Condenser), DISTILL2 (Condenser) - Electric: P1, P2, C1, C2, C3 - LP steam: H3 - MP Steam: DISTILL1 (Reboiler), DISTILL2 (Reboiler) - HP Steam: No blocks use this utility - Oil: H4, H5, H7

Main Outputs

- Purge 8.22% of Inlet Ethylene Feed. - Ethanol flow rate 10,539.4 kg/hr. - Commercial Ethanol Yield is 88.15%. - Pure Ethanol Yield is 81.5%. - 0.0788% of Ethanol is lost in the Purge. - Reaction conversion is equal to 4.78%.

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Aspen Main Flowsheet

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A s s i gn me n t 10 D e s i gn R e p or t - G r ou p 1 ...johanv/Final10.490.pdf · A s s i gn me n t 10 D e s i gn R e p or t - G r ou p 1 Process Control Table 1:Detailed Description