Flash Drum

PROBLEM The flash drum process shown in Figure 1.1 preheats a feed stream to a distillation column. The flash drum vaporizes 40 mole percent of the feed, and the heat exchanger further heats the flash drum liquid. Both the flash drum and heat exchanger have 1 psi (6.9 kPa) pressure drops. A total of 1.0 MM Btu/hr (1.05 MM kJ/hr) is available to the process.

Find the liquid fraction and temperature of the PROD stream and the two duties.

Table 1.1 shows the feed composition and condition. Use the NRTL thermodynamic system, but replace its default vapor enthalpy and vapor density methods with SRKM.

Figure 1.1: Preheat Process

Table 1.1: Feed StreamComponent Mole %

Toluene 80

Methyl-Ethyl Ketone (MEK) 19

Water 1

Flowrate 100 lb-mol/hr 45 kg-mol/hr

Pressure 20 psia 138kPa

Temperature 190 °F 88 °C

Hint: Use one stream Specification and one Define.

1. Define button is in HX spec.

2. Cold side is process Stream

SOLUTION To vaporize 40% of the flash drum feed, use the following stream speci -fication:

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Steam VAPOR Flowrate of All Components on a Wet basis in LBMOL/ HR / Total Unit Feed to FD Flowrate of All Components on a wet basis in LB-MOL / HR = 0.40000 within the default tolerance

Hint: In prod. Spec., click = , then replace it with a / (divide sign).

The duty available to the heat exchanger is 1 MM Btu/hr (1.05 MM kJ/ hr) minus however much duty the flash drum requires to vaporize 40% of its feed. Since the flash duty is unknown before the run, you cannot assign a value to the heat exchanger's duty. Instead, you should have PRO/II calculate the FLASH duty, and use the define feature to pass the calculated flash duty to the heat exchanger. The Define statement for the HEAT EXCHANGER will has the form:

Hear Exchanger HX Duty in x 10^6 BTU/HR = 1.000 – Flash FD Duty in 10^6 BTU/HR

Table 1.2 summarizes the results.

Table 1.2: Results of Flash Drum

English Units Metric Units

Flash DutyHeat Exchanger Duty PROD Temperature PROD Liquid Fraction

0.7353 MM Btu/hr 0.2647 MM Btu/hr 233.96 °F0.6844

0.7678 MM kJ/hr 0.2822 MM kJ/hr 112.20°C 0.6787

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Hydrocarbon/Water Separation

PROBLEM This exercise illustrates the differences between a rigorous VLLE flash and a VLE flash with the water decant option. Simulate a three phase separator as an isothermal flash at 300°F and 300 psia for the crude stream given in the tables below. Use VLLE SRK Kabadi-Danner (SRKKD) thermodynamics and VLE SRK thermodynamics with the water decant option, along with the EOS water solubility correlation (water in the HC phase). Show the K-values for the VLE flash and for the VLLE flash. You can draw 2 separate flash units with the feed to the second unit referring to the feed stream to the 1st unit. This exercise allows you to answer the following questions:■ Is the assumption of a pure water phase valid for this simulation?■ How do the results vary with respect to the water solubility correlations?

Table 2.1: Feed Stream Data

Pure Components Rate Petroleum Rate

(lb moles/hr) Components (lb moles/hr)

Water 3000 Cut 11 165

Carbon Dioxide 35 Cut 12 303

Nitrogen 30 Cut 13 560

Methane 890 Cut 14 930

Ethane 300 Cut 15 300

Propane 520

i-Butane 105

n-Butane 283

i-Pentane 100

n-Pentane 133

Temperature 150°F

Pressure 1000 psia

Table 2.2: Petroleum Fraction Properties

Fraction Molecular Wt. API Gravity NBP (°F)

Cut 11 91 64 180

Cut 12 100 61 210

Cut 13 120 55 280

Cut 14 150 48 370

Cut 15 200 40 495

SOLUTION Defining four thermodynamic sets allows PRO/II to flash the feed iso-thermally four times within one simulation run. Defining the LLE key components can greatly decrease the computation time.

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Note: To enter LLE key components through PROVISION, you must modify the defined VLLE thermodynamic system's LLE K-value data. Within the LLE K-value window, click on the LLE Key Components button, select the user-specified radio buttons, and select your light and heavy liquid phase key components. (Dominant Species).

Figure 33 shows the Flash Drum Summary from the PRO/II output report. Note that this summary only reports the water decant flashes. PRO/II reports the total liquid mole fraction and breaks down the liquid phases into the liquid hydrocarbon phase and the free water phase. USE ONLY DCNT-EOS. Note that PRO/II provides this type of report only when you invoke the water decant option for a hydrocarbon-water system.

The K-value printout for the first FLASH unit (DCNT-EOS) follows the standard Flash Drum Summary. This report gives the vapor and liquid mole fractions and the K-value for each component. This gives you information about the relative volatilities of the components of the products.

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Figure 34 shows the thee-phase Flash Drum Summary from the output report. Note that this summary only reports the rigorous VLLE FLASH units. If you invoke the VLLE option for a FLASH, but only one liquid phase exists, then PRO/II provides the results in the standard Flash, but only one liquid phase exists, then PRO/II provides the results in the standard Flash Drum Summary. The results for the rigorous three-phase FLASH are very similar to those in Figure 34. The Vapor and liquid mole fractions for each component in the three product phases are at the end of the standard three-phase Flash Drum Summary. Pro/II uses these results to calculate the two sets of K-values. You can obtain a set of LLE K-values by dividing the second set of VLE K-values by the first.

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Naphtha Assay - Part 1

PROBLEM Enter the assay data for the Naphtha Feed described below to a flash drum and generate a set of petroleum components. Also create a plot of the calculated TBP curve and component cuts. The light ends make up 5% of the total liquid volume. Use SRK thermodynamics with API liquid density to model the system. To represent the boiling curve more closely, increase the number of pseudocomponents by using twenty-one 18°F (10°C) cuts from 80 cF to 460°F (30°C to 240°C). The process data is given in Table 3.1.

Table 3.1: Naphtha Feed Data

ASTM D86 Data Light Ends

LV% T(°F) T(°C) Component Mole %

3 90 32 isobutane 0.70

5 125 52 butane 2.15

10 195 90 isopentane 0.86

30 250 121 pentane 3.58

50 280 138

70 310 154 Total 5% Liq. volume

90 390 199

95 418 214

98 430 221

Flowrate 1000 Ib/hr 450 kg/hr

Temperature 100°F 38"C

Pressure 1 atm

Average API gravity 54.2 Average Specific Gravity 0.762

SOLUTION Build the flow sheet and enter the UOM, components, and thermodynamic data, as usual.

Step 1 Modify the Assay Characterization Data

To represent the boiling curve more closely, you must increase the number of pseudocomponents by using twenty-one 18°F (10GC) cuts from SOT to 460°F (30°C to 240°C).

Click the Assay Characterization button on the toolbar

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Click Modify Primary Set...

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In the Primary TBP Cutpoints Definition dialog box, change theMinimum Temperature for First Interval to 80°F (30°C).

In first row, enter 460°F (240° C) as the Maximum Temperature for the Interval and 21 as the Number of Pseudocomponents in Interval. Leave the default data for the second and third intervals as they are.

Click OK to return to the Assay Cutpoints and CharacterizationOptions dialog box.

Click Characterization Options... and change the Initial Point from its default to 3%.

Step 2 Define the Assay Stream

Double-click on the stream and select Petroleum Assay as the Stream Type.

Click Flowrate and Assay… and enter the rate.

To define a new assay:

In the Stream Data - Flowrate and Assay dialog box, click Define/Edit Assay...

Choose ASTM D86 and enter the distillation data in the grid. You can use the <Tab> key to move from cell to cell in the grid.

Select API Gravity and enter the average value.

Figure 35 shows these entries.

To enter the lightends composition:

In the Assay Definition dialog box, click Lightends...

Choose Percent of Assay and enter 5 as the LV percent figure.

Check the Assay lightends information for stream NAPTHA_FEED box and enter the molar composition of each component in the grid.

Because the lightends compositions do not add up to 1 or 100, youmust check the Normalize box.

Figure 36 shows these entries.

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Back in the Stream Data dialog box, enter Temperature and Pressure values. Click OK to return to the flowsheet.

You will be asked whether you want to generate the pseudocomponents now or at a later time. Answer "No.

Step 3 Generate the Pseudocomponents

You can generate the pseudocomponents at any time by selecting Gener-

ate Assay Components from the Input menu. Do this now. Then:

To view the list of pseudocomponents, open the Component Selection dialog box.

To look at the petroleum property data generated for the pseudocom-ponents, generate an output report and view it.

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To view Processing Curve, click View Curve... in the Assay Definition dialog

box. (Hint: Stream/Flowrate Assay/Assay Definition/View Curve, Figure 38)

Figure 38: Assay

Processing Plot

Naphtha Assay - Part 2

PROBLEM In this exercise you will use the assay stream you created in NAPHTH1 and manipulate it using two flash drums. The first flash drum predicts the temperature at which 80% of the feed is recovered in the vapor. The second flash drum predicts the quantity of liquid that drops out when a temperature drop is imposed on that vapor.

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Table 3.3: Equipment Data and Operating Conditions __ Unit Description Data

D-1 Flash drum Pressure = 0.8 atm80/20 vapor/liquid split

D-2 Flash drum Temperature Drop = 18°F (10°C)Pressure Drop = 0.1 atm

SOLUTION Open the simulation named NAPHTH1.PRZ and save it under a new name, NAPHTH2.PRZ.

Step 1 Build the Process Flow Diagram

Add two flash drums to the flowsheet and connect, as shown below.

Connect the stream NAPHTHA_FEED to the first flash drum.

Figure 3.3: PRO/IIFlowsheet

Step 2 Enter Unit Operating Conditions

Enter unit data from Table 3.3 into both flash drums.

Express the recovery specification in Flash F1 as the ratio of one stream's flowrate to that of another. You can define a recovery on the basis of one component or a range of components or, as here, on the total stream. Set the specification so that it matches the figure shown below.

Steam F1VAPOR Flowrate of All Components on a Wet basis in LBMOL/ HR / Stream NAPHTHA FEED Flowrate of All Components on a wet basis in LB-MOL / HR = 0.80000 within the default tolerance

Because the feed temperature of flash F2 is unknown until the first flash has executed, you cannot enter an absolute value at this time. You can, however, enter a relative value, as shown below (Hint: Define T).

Flash Drum F2 Temperature in F = Stream F1VAPOR Temperature in F – 18.000

Step 3 Run the Simulation and View the Results

Add a STREAM PROPERTY TABLE to display a Short Property List for all

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the streams in the flowsheet.

Run the simulation and observe the temperatures, pressures and flowrates on the STREAM PROPERTY TABLE. Verify that they accurately reflect the specifications you entered.

Figure 40: Short Property Table

Stream Name

NAPHTHA_FEED

F1 VAPOR F2VAPOR F2LIQUID F1 LIQUID

Temperature F 100.000 281.195 263.195 263.195 281.195Pressure ATM 1.000 0.B00 0.700 0.700 0.800

Flowrate LB-MOL/HR 8.574 6.859 6.346 0.514 1.715

Phase Liquid Vapor Vapor Liquid Liquid

Stream Name

NAPHTHA_FEED

F1VAPOR F2VAPOR F2LIOUID F1 LIQUID

Temperature C 33.000 138.442 128.442 128.442 138.442Pressure ATM 1.000 0.800 0.700 0.700 0.800Flowrate KG-MOUHR 3.858 3.087 2.856 0.231 0.772Phase Liquid Vapor Vapor Liquid Liquid

(OUTPUT/ Stream Property Table/Double Click Stream Property Table to add desired streams)

or (Right Mouse Stream Icon/ View Results)

Naphtha Assay - Part 3

PROBLEM Using the assay stream and flashes from NAPHTH2, you will now examine the effects of varying the vapor/liquid split in the first flash drum on the duty of the second flash drum. Run cases at vapor fractions from 80% (the base case) to 40% in 10% increments.

SOLUTION Open the simulation named NAPHTH2.PRZ and save it under a new name, MAPHTH3.PRZ.

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Step 1 Enter Case Study Data

Click the Case Study button on the toolbar to open the Case Study Parameters and Results dialog box.

Check the Define Case Study box.

In the first row of the Parameter field, click Parameter .

Select Flash from the Unit/Stream list, Fl from the Unit Name list.

Click Parameter and select Product Specification.

Specify 5 cycles and a step size of -0.1.

In the first row of the Result field, click Parameter and select Flash F2 Duty.

Figure 41 shows the completed dialog box.

Figure 41: Completed Case Study

Step 2 Run the Simulation

Click the Run button to run the simulation.

Step 3 View Output in a Tabular Format

Select Output/Case Study from the menu bar and then Tables from the submenu to open the Case Study Tables dialog box.

Enter a table name in the first row and click Data...

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Move the variables you want to appear in the table from the Available Variables list to the Selected Variables list.

Click on Columns to change the default to Display cycles in Rows .

Click View Tales... to view the results.

Step 4 View Output in a Graphical Format

Define a plot that shows the change in F2 duty with the change in F1 split.

Select Output/Case Study from the menu bar and then Plots from the submenu to open the Case Study Plots dialog box.

Enter a plot name in the first row and click Data....

Highlight F1SPLIT from the Available Variables list and X-Axis from the other list.

Enter a label in the Label/Legend field, if desired.

Click Update , followed by Add ->

Highlight F2DUTY from the Available Variables list and Y-Axis#1 in the other list.

Enter a label in the Label/Legend field, if desired.

Click Update , followed by Add ->

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Cyclohexane Plant Part 1: Base Case

Solution

PROBLEM Solve the heat and material balances for the Cyclohexane plant shown in Figure 4.1. Table 4.1 gives the stream conditions. The process conditions and reaction data follow. Do not break the thermal recycle loop or use any special techniques to accelerate convergence.

Figure 4.1 Flowsheet for Cyclohexane Plant

Table 4.1: Stream Data

Process

Conditions

(Hint: Input/Recycle Convergence)

Increase the number of loop trials to 80.

Use the Minimum Tear Streams sequence (the default).

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Benzene(SI) Makeup H2 {S2}

Hydrogen (mol %)

0.0 97.5

Methane {mol %) 0.0 2.5

Benzene (mol %) 100 0.0

Temperature (DF) 100 100

Pressure (psig) 500 500

Flowrate 100 bbl/hr 600,000 {std vap ft3

/ hr)

Use SRK thermodynamics with API liquid density (default).

Heat exchangers have a 5 psi pressure drop on both sides (Cold Product).

Heat exchanger E1 is a shell and tube heat exchanger with one pass shell

side and 2 passes tube side. U=100 Btu/°F-ft2-hr and Area=650 ft2.

Set up the splitter specification so that 20% of the splitter feed exits as the

purge gas.

Reaction Data (Hint: Click Reaction Data button on the tool bar; Conversion Reactor)

Vapor phase reaction: C6H6 + 3H2 —> C6H12

Reactor is jacketed and controlled to 435°F

99.9% conversion of benzene to Cyclohexane (Extent of Reaction)

Heat of reaction (at 77°F) = -87x103 Btu/lb-mol benzene

■

Enter the reaction through the Reaction Data and Reaction Definition dialog boxes, accessed through the Reaction Data button. Figure 4.2 shows the Reaction Definition dialog box after the data has been entered. Don't forget to enter the heat of reaction data, by clicking H...

Figure 4.2: Reaction Definitions Dialog Box:

Run A: Base Case Simulation

From the Output menu, choose Report Format and then Miscellaneous Data. Mark the Include Overall Flowsheet Mass Balance check box. Generate the output report.

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1. How many recycle loop iterations are necessary for the simulation to converge? How much time did this take?

2. What is the percent deviation in the hydrogen and methane material balance?

3. What calculation sequence and tear streams did PRO/II select for this flowsheet?

4. How many loops did PRO/II use in solving the flowsheet?

(Thought: Why is recycle necessary? Why is purge necessary?}

Run B: Tighten TolerancesTighten the recycle stream tolerances to make the relative component tolerance 0.001 and to exclude components with mole fractions below 0.001. Rerun the problem.

1. How many additional iterations are necessary?

2. Do the hydrogen and methane material balances improve?

SOLUTION Run A: Base Case Simulation

1. The run required 29 iterations (10.5 seconds) Naturally the time depends on the computer you use. Don't worry if your time doesn't match this value; we are only looking for trends.

2. The percent deviation is 4.32% for Hydrogen and 4.00% for methane.

3. The sequence is: Ml, El, Fl, SP1, C1, E2, and RX1. Heat exchanger E3 does not appear, because we chose to merge it with the FLASH. The tear streams are S6 and S13. The SIMSCI sequence option always selects the calculation sequence that yields the minimum number of tear streams. When more than one sequence satisfies this minimum tear streams criterion, the sequence selected by PRO/II can depend on the order you entered the unit operations. For this flowsheet, the sequence: Ml, E2, R1, E1, F1, SP1, C1 also yields two tear streams, S13 andS4.

4. One loop is used in solving the flowsheet

Run B: Tighten Tolerances

1. 21 additional iterations (14.5 seconds) were required for a total of 50=29+21 (25 seconds).

2. The hydrogen and methane material balances so improve. The material balance errors decreased to 0.37% for hydrogen and 0.34% for methane.

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LOCATING TERNARY AZEOTROPES ACETONE-METHANOL-CHOLOROFORM

PROBLEM In this exercise, you will first use a multivariable controller to locate the ternary azeotrope for an acetone-chloroform-methanol system. In Part B, you will simulate the same process but use an optimizer to locate the azeotrope. Process data are provided below.

Figure 5.1: PRO/II Flowsheet

Table 5.1: Stream Data

Acetone Chloroform Methanol

Acetone (mol %) 100 0 0

Chloroform (mol %) 0 100 0

Methanol (mol %) 0 0 100

Temperature (°F) 150 120 150

Flowrate (lb-mol/hr) 0.367 0.213 0.420

Use NRTL thermodynamics and include two-phase LLE calculations.

Operate the flash drum at the bubble point and 0 psig.

Using a Multivariable Controller

Add a multivariable controller and a calculator to the flowsheet shown in Figure 5.1 and use it to locate the ternary azeotrope for an acetone-methanol-chloroform system.

.

SOLUTION

In the Calculator dialog box, change the stream sequence to VAPOR, LIQUID1, and LIQUID2

and enter the procedure shown in Figure 5.2.

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Set up the MVC as shown in Figure 5.3:

Figure 5.3:

Multivariable

Controller

Box

Hint: Click the number under variables and then click insert to add a specification

Definition of Azeotrope: mole fraction of combined liq / mole fraction of Vap = 1

Figure 5.4 shows the output results for the product streams, VAPOR and LIQUID1.

Figure 5.4: ProductStreams Summary

This example converges rapidly, provided your initial guesses are not too far from the solution. At 0 psig,

PRO/II Work Book 41

Stream Name Stream

DescriptionVAPOR LIQUID1

Phase Vapor Liquid

Temperature PressureF

PSIG134.534

0.000134.534

0.000

Flowrate LB-MOLJHR 0 .000 1.000

Composition

ACETONE

METHANIOL

CLFR

0.3660.4210.213

0.3670.4200.213

the ternary azeotrope occurs at 134.53 °F, Xacetone = 0.367, Xmethanol = 0.420, Xchloroform = 0.213. (Hint:

Bubble Point No Vapor Flow) (Right Mouse on a stream, select view Results or Data Review

window)

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Propane-Propylene Splitter Part 1: Base case simulation

PROBLEM Solve the heat and material balances for the propane-Propylene splitter shown in Figure 6.1. Process data and specifications are provided below.

Figure 6.1: FlowsheetFor Propane-

Propylene Splitter

Table 6.1: Feed Stream Data

Component Mole %

Propane 14.6112Propylene 85.3888Flowrate 25550 lb/hrTemperature 120 FPressure 310 psia

Column Specifications:

Condenser is at Bubble point and at pressure of 272 psia.

Pressure Profile

SP

22

92

92

13

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Other Column Specifications:

1. Reboiler type = Kettle

2. Overhead Estimate = 500 lbmole/hr.

3. Feed Location Stage = 105

Use SRK thermodynamics with Costald liquid density.

Use the I/O algorithm with the Conventional initial estimate generator.

Specify the mole fraction of propylene in the overhead stream to be 0.9968.

Specify the recovery of propylene in the bottom stream to be 5 % of the propylene in the feed.

Vary the reboiler and condenser duties to meet these specifications.

Hint: Bottom Propylene Flowrate/Feed propylene Flow Rate = 0.05

1. What does the flat portion of separation factor plot tell you about the current feed tray location?

2. What is the reboiler duty?

SOLUTION

1. Output/Generate Plot Your separation factor plot should look similar to Figure 6.2. The rel -atively flat region between approximately trays 100 & 130 indicates that the feed tray locations is not optimal. The trays in this region are not effectively contributing to the separation of the two components.

2. The reboiler duty is approximately 47.82 x 10^6 BTU/hr

Figure 6.2

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Propane-Propylene Splitter Part2: optimize the Feed Tray Location

PROBLEM In this and succeeding parts of this problem, you will use PRO-II’s OPTIMIZER utility to determine the feed tray location that minimizes the reboiler duty for the propane-propylene splitter flowsheet.

Add an OPTIMIZER to the flowsheet you built for Part 1 of this problem.

Define the OPTIMIZER objective function to minimize the reboiler duty.

Set up the OPTIMIZER to search for the best feed tray location between trays 50 and 150.

Change the following convergence tolerances on the COLUMN:

(Double Click Column Icon and Click on column convergence data).

• Relative tolerance for both column specifications to 1 x 10-5 (Performance Spec.)• Bubble point tolerance to 0.000001• Enthalpy balance tolerance to 0.000001• Equilibrium K-value tolerance to 0.000001

Answer the following questions:

1. What is the new feed tray location? Is this what you expected basedon the separation factor plot?

2. What is the new reboiler duty?

3. Plot the separation factor. How has the new feed tray location affected the separation factor profile?

SOLUTION Remember to select Input/Restore Input Data from the menu bar between runs.

The New feed tray location for Feed is stage 130

The reboiler duty is now 44.7 x 10^6 BTU/hr

Figure 6.3 shows the

new plot. Notice that

the flat spot is gone,

indicating an

improvement in the

feed tray location.

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Propane-Propylene Splitter Part3: Changing the Optimizer Parameters

PROBLEM In Part 2 of this problem, you changed the COLUMN convergence tolerances which had a significant impact on the calculated optimum feed tray location.

In this part, change the following OPTIMIZER parameters:

Set the absolute maximum step size to 5. (Step size)

Set the defined absolute step size to 1.

Set the minimum relative change for the objective function to 0.0001. (Option)

Rerun the problem with the new parameters and answer the following questions:

1. What are the new feed tray locations?

2.How big an effect does changing the OPTIMIZER parameters have on the feed tray location when compared to the effect of changing the COLUMN convergence tolerances?

SOLUTION 1. The feed tray location is now 130.6

2. The OPTIMIZER parameters had a smaller affect than the COLUMN

convergence tolerances. In flowsheets containing CONTROLLER and

OPTIMIZER loops, accuracy is often more strongly impacted by the tolerances of

the unit operations that make up the loop, rather than by the CONTROLLER or if

OPTIMIZER. A good rule thumb is to keep the unit operation tolerances an order

of magnitude lower than the CONTROLLER and OPTIMIZER tolerances.

Propane-Propylene Splitter Part 4: Maximize Profit

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PROBLEM Use PRO/II's CALCULATOR and OPTIMIZER units to evaluate and maximize the profit for the propane-propylene splitter. Base the profit on the economic data found as follows in the Tables.

Economic Data

Utility Cost ($/MM Btu)

LP SteamCooling Water

3.60.40

Product Product ($/LB)

PropylenePropane

0.1740.120

In this part of the exercise, vary the following parameters:

Feed tray location between trays 50 and 150

Recovery of Propylene in the bottoms product between 0.01 and 0.1. This corresponds to COL1SPEC1 or COL1SPEC2 in the OPTIMIZER.

Hint: Remember to Restore input data; Define Parameter in Calculator 1; Click the number then insert the 2nd variable; Ht Duty of Cooling Water is negative; Tolerances (Column: 1.E-6; Optimizer: 1.E-5)

Questions:

1. What is the new feed tray location?

2. What is die new value for the recovery specification?

3. Are either of the variables at the bounds you provided?

4. Explain why this recovery specification corresponds to the optimum

solution in this exercise?

SOLUTION

1. The New Feed tray location is stage 118.9.

2. The New recovery specification is 0.01.

3. The recovery specification is at its lower bound. It is quite common to find the optimum solution

at one or more of the variable bounds.

4. The OPTIMIZER guided the COLUMN to the solution that yields the least propylene in the

bottom product. Although this sharp split requires high reboiler and condenser duties, it frees

up more propylene for the high-value overhead product. In this example, the additional

energy costs are more than offset by the increased flowrate of the propylene stream.

Multiple-Feed Distillation Column Part 1: Base case simulation

PROBLEM Solve the heat and material balances for the propane-butane splitter shown in

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Figure 7.1. Process data and specifications are provided below.

Figure 7.1: Flowsheet

For Propane-Butane Splitter

Table 4: Feed Stream Data

Lb-moles per hr

Component FEED 1 to stage 15 from the bottom FEED 2 to stage 6 from the bottom

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Ethane 1.5 0.5Propane 24.0 10.0n-Butane 16.5 22.0n-Pentane 7.5 14.5n-Hexane 0.5 3.0

Both feeds are at Bubble point and at pressure of 250 psia.

Specifications:

4. Top tray Pressure = 250 psia

5. Condenser type = Partial

6. Condenser Pressure = 250 psia

7. Reboiler type = Kettle

8. Distillate rate = 36.0 lbmole/hr.

9. Reflux rate = 150 lb mole/hr.

10. Initial Estimates Method = Conventional

Reflux estimate (Bulk liquid rate) = 186 lbmole/hr

11. Column Performance Specification:

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Use SRK thermodynamics with Costald liquid density.

Use the I/O algorithm with the Conventional initial estimate generator.

Vary the reboiler and condenser duties to meet these specifications.

3. What does the flat portion of separation factor plot tell you about the current feed tray location?

4. What is the reboiler duty?

SOLUTION

3. Your separation factor plot should look similar to Figure 23. The relatively flat region between approximately trays 11 & 18 and 24 & 27 indicates that the feed tray locations are not optimal. The trays in this region are not effectively contributing to the separation of the two components.

4. The reboiler duty is approximately 1.1283 x 10^6 BTU/hr

5.

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Multiple-Feed Column Part2: optimize the Feed Tray Location

PROBLEM In this and succeeding parts of this problem, you will use PRO-II’s OPTIMIZER utility to determine the feed tray location that minimizes the reboiler duty for the propane-propylene splitter flowsheet.

Add an OPTIMIZER to the flowsheet you built for Part 1 of this problem.

Define the OPTIMIZER objective function to minimize the reboiler duty.

Set up the OPTIMIZER to search for the best feed tray location between trays 7 and 30.

Change the following convergence tolerances on the COLUMN:

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(Double Click Column Icon and Click on column convergence data).

• Relative tolerance for both column specifications to 1 x 10-5

• Bubble point tolerance to 0.001

• Enthalpy balance tolerance to 0.001

• Equilibrium K-value tolerance to 0.001

Answer the following questions:

4. What is the new feed tray location? Is this what you expected basedon the separation factor plot?

5. What is the new reboiler duty?

6. Plot the separation factor. How has the new feed tray location affected the

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separation factor profile?

SOLUTION Remember to select Input/Restore Input Data from the menu bar between runs.

The new feed tray location for Feed 1 is between 16 and 17

The New feed tray location for Feed 2 is between 21 and 22

The reboiler duty is now 1.1231 x 10^6 BTU/hr

Figure 24 shows the new plot. Notice that the flat spot is gone, indicating an improvement in the feed tray location.

Figure 24:

Multiple-Feed Column Part3: Changing the Optimizer Parameters

PROBLEM In Part 2 of this problem, you changed the COLUMN convergence tolerances which had a significant impact on the calculated optimum feed tray location.

In this part, change the following OPTIMIZER parameters:

Set the absolute maximum step size to 1.

Set the defined absolute step size to 1.

Set the minimum relative change for the objective function to 0.0001.

Rerun the problem with the new parameters and answer the following questions:

3. What are the new feed tray locations?

4. How big an effect does changing the OPTIMIZER parameters have on the feed tray location when compared to the effect of changing the COLUMN convergence tolerances?

SOLUTION 1. The feed tray locations are now between 16 and 17 for Feed 1

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and 23 and 24 for Feed 2.

2. The OPTIMIZER parameters had a smaller affect than the COLUMN

convergence tolerances. In flowsheets containing CONTROLLER and

OPTIMIZER loops, accuracy is often more strongly impacted by the tolerances of

the unit operations that make up the loop, rather than by the CONTROLLER or if

OPTIMIZER. A good rule thumb is to keep the unit operation tolerances an order

of magnitude lower than the CONTROLLER and OPTIMIZER tolerances.

Multiple-Feed Column Part 4: Maximize Profit

TASK Use PRO/II's CALCULATOR and OPTIMIZER units to evaluate and maximize the profit for the propane-propylene splitter. Base the profit on the economic data found as follows in the Tables.

Economic Data

Utility Cost ($/MM Btu)

LP SteamCooling Water

3.64.0

Product Product Selling Rate ($/LB)

EthanePropanen-Butanen-Pentanen-Hexane

0.001240.23

0.1960.250.30

In this part of the exercise, vary the following parameters:

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Feed tra

y locations between trays 16 & 17 for Feed 1 and 23 & 24 for Feed 2

Recovery of Propane in the bottoms product between 0.01 and 0.1. This corresponds to COL1SPEC1 or COL1SPEC2 in the OPTIMIZER.

Result Summary of Calculator:

Calculator 'CA1'

User Input Calculated

Parameter 1 N/A 1.1232E+00

2 N/A -8.2622E-01

3 N/A 1.5614E+03

4 N/A 4.1248E+03

5 N/A 5.5555E-02

6 N/A 9.4053E-01

7 N/A 3.9194E-03

8 N/A 2.3831E-08

9 N/A 1.8520E-16

10 N/A 3.1722E-07

11 N/A 2.2085E-03

12 N/A 5.9936E-01

13 N/A 3.4375E-01

14 N/A 5.4687E-02

15-50 N/A

Result 1 N/A 4.3742E+00

2 N/A 3.3907E+02

3 N/A 9.0879E+02

4 N/A 1.2479E+03

5 N/A 2.4520E+02

6-200 N/A

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