Design and control of dual condensers in distillation columns

Design and control of dual condensers in distillation columns

Accepted Manuscript Title: Design and Control Of Dual Condensers in Distillation Columns Author: William L. Luyben PII: DOI: Reference: S0255-2701(13...

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Accepted Manuscript Title: Design and Control Of Dual Condensers in Distillation Columns Author: William L. Luyben PII: DOI: Reference:

S0255-2701(13)00191-8 http://dx.doi.org/doi:10.1016/j.cep.2013.08.007 CEP 6329

To appear in:

Chemical Engineering and Processing

Received date: Revised date: Accepted date:

20-5-2013 11-8-2013 23-8-2013

Please cite this article as: W.L. Luyben, Design and Control Of Dual Condensers in Distillation Columns, Chemical Engineering and Processing (2013), http://dx.doi.org/10.1016/j.cep.2013.08.007 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.

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Paper submitted to Chemical Engineering and Processing; Process Intensification

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Design and Control Of Dual Condensers in Distillation Columns

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William L. Luyben

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Department of Chemical Engineering Lehigh University Bethlehem, PA 18015 USA

May 18, 2013 Revised August 11, 2013

[email protected]; 610-758-4256; FAX 610-758-5057

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Abstract Most distillation columns use water-cooled condensers cooling water is inexpensive. With cooling water supplied at 305 K, a reflux-drum temperature of 322 K

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is normally used for column design. This temperature and the distillate composition set the require column pressure.

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This paper demonstrates that the use of refrigeration in a second condenser in series with the primary water-cooled condenser has economical advantages in some

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separations. The dual-condenser process can be less expensive than the single-condenser process in cases in which a lighter-than-light-key component is present in the feed and in

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which the separation between the key components is difficult (low relative volatility systems). Using a small refrigerated condenser permits the column to operate at a lower

refrigeration required.

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pressure, which reduces reboiler duty enough to compensate for the small amount of

The dynamic control of the dual-condenser process is also studied, and an

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Key Words:

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effective control structure is developed.

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refrigerated condenser, distillation control, dual condensers, de-isobutanizer, c3-splitter

1. Introduction

The design and operating pressures in the majority of distillation columns are set

by the desire to use inexpensive cooling water in the condenser. A minimum reflux-drum temperature of about 322 K is usually used to permit a reasonable temperature differential between the cooling water and the process. Typical inlet and outlet cooling water temperatures are 305 K and 316 K. This temperature and the specified composition of the distillate product set the column pressure.

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If the column feed is strictly a binary mixture of light-key and heavy-key components, the typical column design uses a conventional water-cooled condenser. However, in almost all cases, there are small amounts of lighter-than-light-key components in the feed to the column. These all must leave in the distillate, and their

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presence raises the required pressure at the fixed 322 K reflux-drum temperature. This paper explores this situation and illustrates that the use of a second condenser with

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refrigerant in series with the primary water-cooled condenser can reduce overall energy

costs in some systems. Only qualitative discussions of this dual-condenser configuration

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have been found in the literature (Kister1).

For example, in the de-isobutanizer column discussed later in this paper with 2

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mol% ethane in the feed that is mostly isobutane and normal butane, the distillate composition is 3.92 mol% ethane, 91.28 mol% isobutane and 4.80 mol% n-butane (the design specification). Assuming a total condenser and a reflux-drum at 322 K, a

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bubblepoint calculation gives a pressure of 7.735 atm. The reboiler duty of this column with a water-cooled condenser is 10.41 MW. If the column pressure could be reduced to

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6.8 atm, the reboiler duty could be reduced 9.929 MW. This can be achieved by using a dual refrigerated condenser with only a small refrigeration load (0.4189 MW).

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Several chemical separations and parameter values are considered in the

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following sections to assess the effects of (1) relative volatility between the key components, (2) the volatility and amount of the lighter-than-light-key component in the feed and (3) the product purity levels.

2. De-Isobutanizer Column (DIB) The separation of isobutane and n-butane is one of the most important industrial

distillation systems. These components have normal boiling points that are fairly close (261.43 K for iC4 and 272.65 K for nC4) so the separation requires a fairly large number of trays and a fairly high reflux ratio. Typical columns contain 50 trays and have reflux ratios around 6. Operation at elevated pressure (6 atm) is required if cooling water is used as the heat sink. The mixture of C4 components to be separated often contains small amounts of lighter components such as propane or ethane.

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As an example, consider a DIB column with 51 stages and a feed with 2 mol% ethane (C2) and equimolar concentrations of iC4 and nC4. The feed flowrate is 500 kmol/h. Note that ethane, not propane, is used as the lighter-than-light key-component. Later in this section we will demonstrate that the lighter-than-light-key impurity must be

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sufficiently more volatile than the light key component for the dual-condenser setup to be economical.

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The column is designed to produce a bottoms product with an impurity of 5 mol% iC4 and a distillate with an impurity of 4.804 mol% nC4. Since all the C2 in the feed goes

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overhead in the distillate, its composition is 3.92 mol% C2. The composition of the distillate on a C2-free basis is 95 mol% iC4. Aspen simulations with Choa-Seader

2.1 Single Water-Cooled Condenser

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physical properties are used in the designs.

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Figure 1 gives the flowsheet for this conventional system. The pressure in the condenser is 7.735 atm for the specified distillate composition at 322 K. The reboiler duty required to meet the two product specifications is 10.41 MW, and the reflux ratio is

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temperature is 345 K.

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7.074. Low-pressure steam at 433 K ($7.78 per GJ) is used in the reboiler since the base

These results with 2 mol% light C2 impurity in the feed should be compared to

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what the design would be if the feed contained no ethane. The column pressure would be lower (6.179 atm versus 7.735 atm) for the same 322 K reflux-drum temperature, and the reboiler duty would be lower (9.946 MW versus 10.41 MW).

2.2 Dual Condensers

Figure 2 gives the flowsheet for process with two condensers in series. The first

uses cooling water and achieves a 322 K temperature in the reflux drum. Now, however, the pressure in the condenser is 6.8 atm. The lower pressure in the column makes the separation easier and reduces the reboiler duty to 9.929 MW and reflux ratio to 6.79. However, not all the overhead vapor from the top of the column is condensed at this lower pressure. A small vapor stream is removed from the top of the reflux drum (78.61 kmol/h). The stream is more rich in ethane (9.16 mol%) than the final distillate (3.92 mol%) and requires a temperature of 305 K to condense it in the refrigerated

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condenser. Chilled water can be used (supplied at 278 K and returned at 288 K) at a cost of $4.43 per GJ (Turton et al2). The refrigeration load is very small 0.4189 MW compared to the heat duty in the first condenser (9.378 MW) since all of the required reflux (1731 kmol/h) and a large

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portion (176.4 kmol/h) of the total distillate (255 kmol/h) are condensed in the water-

cooled condenser. Only the 78.61 kmol/h of vapor not condensed in the first condenser

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must be cooled and condensed in the second condenser.

Table 1 provides an economic comparison of the conventional and dual-condenser

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cases. Designs over a range of pressures for the dual-condenser flowsheet are shown. The optimum pressure in terms of total annual cost (TAC) is seen to be the case at a pressure

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of 6.8 atm.

Lower pressures produce larger vapor flowrates (DV) from the reflux drum, which increase the refrigeration load (Qrefrig). But reboiler duty (QR) also decreases

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because the separation is easier. Total energy cost (reboiler steam plus refrigeration) goes through a minimum at a pressure of 6.8 atm. Total capital investment decreases as

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pressure is decreased. The optimum pressure in the dual-condenser system is 6.8 atm.

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Comparing the conventional and the dual-condenser cases shows that the dualcondenser flowsheet has lower capital investment, lower total energy costs and a lower TAC.

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2.3 Propane as Light Impurity

The effect of having a less volatile lighter-than-light-key impurity in the feed is

explored by assuming there was 2 mol% propane instead of ethane. Assuming the same column and specifications used above, the conventional column operates at 6.523 atm with a reboiler duty of 9.856 MW. A pressure of 6.45 atm in the dual-condenser system gives only a slightly lower

reboiler duty (9.796 MW) than the conventional design. Of course there is the additional cost of refrigeration. The refrigeration load is 0.2808 MW to condense the 57.27 kmol/h of vapor from the reflux drum. The economics show that the small decrease in reboiler energy is too small to compensate for the refrigeration cost. Total energy cost increases from $2,418,000 per year to $2,443,000 per year. Capital investment also increases from $1,785,000 to $1,800,000. Clearly the conventional design is better in this situation. 5

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This result indicates that the dual-condenser design is only economical when the lighter-than-light-key component is significantly lighter than the light key. However, the effect of pressure on the separation and the difficulty of the separation are other important

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parameters. In the next section we look at a more difficult separation.

2.4 Effect of Product Purity

The impurity levels assumed in the studies described in previous sections were 5

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mol%. If these specifications are lowered to produce higher purity products, the

separation is more difficult and the effect of reducing pressure is more pronounced. To

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illustrate the effect of this important variable, the two designs are compared for the case in which the bottoms specification is 2 mol% iC4 and the distillate specification is 2

the dual-condenser design is kept at 6.8 atm.

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mol% nC4 (on a C2-free basis). The number of stages is kept at 51, and the pressure in

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Results are presented in Table 2 and should be compared with those given in Table 1 for the 5 mol% impurity case. The higher purities require more reboiler energy, which increases both energy and capital costs (larger diameter column and large heat-

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exchanger area). But there are larger differences in the economics between the

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conventional and dual-condenser design because of the increase in the difficulty of separation. The reduction in pressure in the dual-condenser design is more pronounced.

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The reboiler energy difference between the two alternative flowsheets is 0.48 MW with 5 mol% impurities but is 1.20 MW with 2 mol% impurities. The improvement in the economics (difference between the conventional TAC and the dual-condenser TAC) has increased to 5.8% in the 2 mol% case from only about 2% in the 5 mol% case. These results demonstrate that product purities are another important parameter

that must be considered along with relative volatility differences, LLK feed concentrations.

3. C3 Splitter The separation between propylene and propane is another very important distillation process in chemical plants and petroleum refineries around the world. The relative volatility of these components is only 1.136 at 322 K. The relative volatility of

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the iC4/nC4 system is 1.321 at 322 K. Therefore the distillation columns (called C3 splitters) for the propylene/propane system have more trays, higher reflux ratios and higher pressures than the DIB columns. As a numerical example, we look at a column with 90 trays and producing

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products that contain 5 mol% impurities. The feed composition is 2 mol% ethane, 49

mol% propylene and 49 mol% propane. The distillate composition is 3.92 mol% C2 with

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the 2 mol% C2 in the feed.

The C3 splitter with one conventional water-cooled condenser must operate at

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20.365 atm with a 322 K reflux-drum temperature. A reflux ratio of 18.18 is required to make the specified separation, giving a reboiler duty of 16.19 MW.

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In the dual-condenser design, the optimum pressure is 20 atm, which only slightly reduces the required reflux ratio to 17.86 and the reboiler duty to 15.99 MW. The refrigeration load is 0.2557 MW. The dual-condenser design has capital and energy costs

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that are only slightly smaller than the conventional single-condenser design. The TAC of the conventional process is $5,108,400 per year. The TAC of the dual-condenser process

2 mol% ethane.

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is $5,083,000 per year. So there is little economic advantage when the feed contains only

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However, if the C2 concentration in the feed is increased to 4 mol%, the dual-

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condenser process becomes much more attractive. With the higher amount of lighterthan-light-key component in the feed, the design of the conventional column has a pressure of 21.71 atm, a reflux ratio of 18.5 and a reboiler duty of 16.49 MW. The optimum design of the dual-condenser process has a pressure of 20.24 atm, a reflux ratio of 17.22 and a reboiler duty of 15.68 MW. Refrigeration load is 0.8582 MW. Now the dual-condenser process has a lower capital cost ($3,369,000 versus

$3,524,000) and a lower total energy cost ($3,967,000 per year versus $4,046,000 per year). The TAC is therefore significantly smaller ($5,096,000 per year versus $5,220,000 per year). These results illustrate that the amount of lighter-than-light-key component in the feed is an important factor in comparison of the alternative flowsheets.

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4. Debutanizer The ease of separation (relative volatilities) are significantly affected by changes in pressure in the two system studied above (isobutane/n-butane and propylene/propane).

fairly substantial reductions in the required reboiler duty.

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Reducing the pressure in the column by using a second refrigerated condenser causes

In this section we take a look at the n-butane/n-pentane system that operates at

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lower pressure and shows less sensitivity of relative volatility to pressure reductions. The

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separation is fairly easy, so fewer trays and lower reflux ratios are required. As a numerical example, we consider a 21-stage column with a feed that is 2 mol% C2, 49 mol% nC4 and 49 mol% nC5. Design specifications are 2 mol% nC4 in the bottoms and

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1.923 mol% nC5 in the distillate (giving a distillate that is 2 mol% nC5 on a C2-free basis since the distillate composition is 3.92 mol% C2).

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The single-condenser design operates at 6.11 atm with a reboiler duty of 4.219 MW. The dual-condenser design operates at 5.8 atm with a reboiler duty that is only slightly smaller (4.146 MW). Capital investment is higher, and total energy cost is only

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slightly smaller. The result is a TAC for the dual-condenser process that is $1,291,700

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per year compared to the conventional process TAC of $1,294,500 per year. These results demonstrate that the separation between the key components must

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be significantly affected by pressure reductions in order for the dual-condenser process to be economical. Note that in this system the lighter-than-light-key component (ethane) is much more volatile than the light-key component (n-butane). Despite this fact, the dualcondenser configuration is economically unattractive. To summarize the result of this study, we have found that the important factors

are:

1. Difference in volatilities between the lighter-than-light-key component and the light-key component. 2. The concentration of the lighter-than-light-key component in the feed. 3. The pressure sensitivity of ease of separation between the key components. 4. The purities of the product streams.

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5. Control of High-Purity De-Isobutanizer Column The control of a conventional single-condenser de-isobutanizer has been discussed in the distillation literature for several decades. The high-reflux ratio and

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difficult separation with little temperature change in the column require a control

structure with several unique features. The traditional control structure has the following

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loops:

1. Feed flowrate is flow controlled (or set by an upstream process).

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2. Pressure is controlled by manipulating the flowrate of the condenser cooling water.

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3. Since the reflux ratio is high, reflux-drum level is controlled by the flowrate of the reflux.

4. Base level is controlled by the flowrate of the bottoms product.

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5. Dual composition control is used since the separation is difficult and the flat temperature profile makes temperature control ineffective in many cases. There

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are two composition controllers. The impurity of iC4 in the bottoms is controlled

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by manipulating reboiler heat input. The impurity of nC4 in the distillate is

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controlled by manipulating the liquid distillate flowrate from the reflux drum.

5.1 Control Structure

The control of the dual-condenser process has apparently not been discussed in

the literature. Since the 2 mol% impurity case provides more economic incentive than the 5 mol% impurity case, we use the high-purity DIB column to explore dynamic control. The control of pressure is the essential issue. We do not want to manipulate the

flowrate of the cooling water for pressure control because its flowrate should be set at the maximum so that the use of expensive refrigeration is minimized. However, heat removed by the very large flowrate of cooling water (41,000 kmol/h) is 13.19 MW and condenses 96.6% of the vapor coming out the top of the column. Only 3.4% of this vapor flows to the refrigerated condenser whose duty is only 0.4777 MW. The flowrate of the chilled water is fairly small (4000 kmol/h) and only contributes a small fraction of the total condensation in the overhead system.

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Therefore trying to control pressure by manipulating the refrigerant flowrate is very ineffective. Such a control structure would be able to handle only very small disturbances. An appropriate analogy would be trying to stop a freight train by dragging your feet. Dynamic simulations quickly revealed this fundamental problem.

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This process offers a nice example of the inherent conflict between design and

control that occurs in many chemical processes. A process design that looks good from a

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steady-state perspective is not always easily controllable.

One approach to solve this problem would be to attempt to develop a

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sophisticated complex control system. However, as is often the case in process control, it is much more practical and effective to modify the process to improve its basic

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controllability.

The process modification made in the dual-condenser process involves installation of a control valve in the vapor line between the reflux drum and the

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refrigerated condenser. From a basic control perspective, this process change provides an additional control degree of freedom since the position of the vapor valve can be

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manipulated to control the pressure in the reflux drum, which establishes the pressure in the column.

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The control structure implemented uses the two manipulated variables (vapor

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valve position and refrigerant flowrate) to control two pressures (reflux-drum pressure and refrigerated-drum pressure). The effectiveness of the structure is demonstrated below.

Of course some pressure drop is required to be taken over the new vapor valve.

The pressure in the reflux drum is 6.8 atm. We select a pressure in the second drum of 5 atm. With a 0.1 atm pressure drop over the refrigerated condenser, the design pressure drop over the vapor valve is 1.7 atm. The vapor flowrate from the reflux drum is 90.4 kmol/h. The temperature of the vapor leaving the reflux drum and going to the refrigerated condenser is 322 K The lower pressure in the refrigerated condenser produces a lower process exit temperature leaving the refrigerated condenser of 294 K versus 306 K as shown in Table 2. The area of the refrigerated condenser must be increased to account for the smaller temperature driving force between the process and the refrigerant (inlet at 288 K and exit

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at 292 K with a flowrate of 4000 kmol/h). The new design area is 45.1 m2 compared to 31.15 m2 shown in Table 2 for the 6.8 atm case). In order to be able to handle disturbances, the refrigeration area was further increased to 75 m2. This represents only a small increase in capital investment. As shown

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in Table 2, the area of the water-cooled condenser is 1210 m2 and the area of the reboiler is 250.6 m2. So a relatively small 44 m2 increase in total heat exchanger area has a minor

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effect on capital investment. The larger area changes the refrigerant load only slightly

(0.4777 MW in the original design compared to 0.5198 MW in the modified design). The

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flowrate of the refrigerant is reduced to 1010 kmol/h at design conditions. The refrigerant valve is designed to be only 7% open at design conditions so large increases in the

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refrigerant flowrate can be achieved if needed to handle disturbances.

Figure 3 shows the control structure with the two pressure control loops. The control valve in the cooling water line remains at a fixed position, so the flowrate of

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cooling water is constant. The pressure in the column is controlled by manipulating the valve in the vapor line from the reflux drum to the refrigerated condenser. The pressure in

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the drum after the refrigerated condenser is controlled by manipulating the flowrate of the chilled water to the refrigerated condenser. The level in the refrigerated drum is

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controlled by manipulating the flowrate of liquid from this drum. The rest of the loops are

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the same as outlined above for the conventional column. Notice that the total distillate stream is the sum of the liquid streams from the two

drums, not including the reflux flowrate. The composition of this total distillate is controlled by manipulating the flowrate of the liquid withdrawn from the reflux drum (not including the reflux, which is used to control the reflux-drum level because of the high reflux ratio). Thus the total distillate is a blend of the two liquid streams. The liquid from the refrigerated drum is 87.94 kmol/h, and the liquid distillate from the reflux drum is 167.0 kmol/h. The flowrate of reflux is 2431 kmol/h.. Implementing the two-condenser system in Aspen was achieved by using a Radfrac model for the column that is a stripper (no condenser). The two condensers are both implemented using HeatX models so that the flowrates of the cooling water and the chilled water can be specified. The cooling water flowrate is held constant in the proposed control system, not the condenser heat removal. Both drums were implemented

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using Flash2 models. The control valve in the vapor line from the refrigerated drum is completely closed during the simulations.

5.2 Dynamic Simulation

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The steady-state file in Aspen Plus is exported to Aspen Dynamics after holdups

in the column base and the two drums have been specified to provide 5 minutes of holdup

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when half full, based on the total liquid flowrate in or out of the vessel.

The three level controllers are proportional with KC = 2. The pressure controllers

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have the Aspen default settings of KC = 20 and I = 12 minutes. The two composition controllers are tuned sequentially. Since the CCxD loop is blending two streams, its

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dynamics are fast with the composition deadtime the only dynamic element in the loop. So the distillate composition controller CCxD is tuned with the bottoms loop on manual. Then the bottoms composition controller CCxB is tuned with the distillate loop on

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automatic. Deadtimes of 3 minutes are inserted in the loops to account for composition measurement delays. Relay-feedback testing and Tyreus-Luyben settings give controller

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tuning constants KC = 0.40 and I = 37 minutes for the CCxB controller and KC = 22 and

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I = 13 minutes for the CCxD controller.

The effectiveness of the control structure is demonstrated by imposing step

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disturbances in feed flowrate or in feed composition. Figure 4 gives the responses of the system to step increases in the setpoint of the feed flow controller at time equal 0.5 hours. The solid lines are for a 15% increase, and the dashed lines are a 10% increase. The +10% increase is easily handled. Both pressures (P and Prefrig) are brought back to their setpoints, as are the compositions of the two product streams (xD and xB). Notice that the flowrate of vapor from the reflux drum DV and the flowrate of the refrigerant Frefrig are more than double their design values. However, the +15% disturbance causes the valve in the vapor line from the reflux drum to go wide open, so the column pressure P is not held at its setpoint. Despite the change in column pressure, the two composition controllers are able to return the product compositions to their setpoints. The pressure in the refrigerated drum also cannot be controlled with the refrigerant valve wide open and the refrigerant flowrate about 15

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times the design value. Notice that the heat removal in the water-cooled condenser (QC) changes despite the fact that the flowrate of the cooling water is held constant. Figure 5 gives results when the feed is reduced by 10 or 15%. Column pressure P is held at its specified value. However, the pressure in the refrigerated drum is not held at

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5 atm. It increases to about 6.5 atm despite the fact that the vapor flowing to the

refrigerated condenser is reduced from about 90 to10 kmol/h. The composition of this

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stream is richer in the light C2 component (increases from 8.6 mol% C2 at design

conditions to 17.7 mol% C2), so the pressure is higher despite having a low temperature

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(288 K) because of the refrigerant flow rising to its maximum.

Figure 6 gives the response of the system to step changes in the feed composition

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at time equal 0.5 hours. The solid lines are when the composition of C2 in the feed is increased from 2 to 3 mol%, with an appropriate decrease in iC4. The dashed lines are when the composition of C2 in the feed is decreased from 2 to 1 mol%, with an

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appropriate increase in iC4. Both of these disturbances are well handled by the control structure in terms of product composition control. The pressure in the refrigerated

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condenser cannot be held at 5 atm when more C2 enters in the feed. The flowrate of vapor to the refrigerated condenser (DV) increases, and the refrigerant valve goes wide

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open (maximum flowrate of the chilled water Frefrig).

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Figure 7 gives results when the temperature of the cooling water entering the first condenser changes. This type of disturbance can occur when rapid changes in weather conditions affect cooling-tower operation. The flowrate of cooling water is constant. The solid lines are for an increase in temperature from the design value of 305 K up to 308 K. The dashed lines are for a decrease in temperature from the design value of 305 K down to 303 K. These are quite large disturbances since the differential temperature driving forces in the heat exchanger are only 326 – 317 = 9 K at the hot end and 322 – 305 = 17 K at the cold end.

The increase in temperature causes a decrease in the heat transferred in the watercooled condensers (QC), which increases the vapor DV to the refrigerated condenser. But column pressure is held at 6.8 atm. The refrigerated drum pressure stay close to the design 5 atm, but the refrigerant valve eventually goes wide open. The distillate from the

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first drum DL goes to zero, which causes the purity of the total distillate xD to end up at a slightly higher impurity. The decrease in cooling water inlet temperature causes an interesting counterintuitive effect. We would expect the colder water would make it easier to maintain

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column pressure. However, as the upper left graph in Figure 7 shows, the column

pressure rises to a steady state of 7 atm. This occurs because the flowrate of vapor DV to

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the refrigerated condenser becomes very small and very rich in the light C2 component. Thus even with the refrigerant valve wide open and a resulting low temperature, the

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pressure in the refrigerated drum rises to 7 atm, which raises the column pressure to 7 atm. The temperature in the first drum falls from 322 to 317 K, and the liquid distillate

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DL from this drum becomes the total distillate product DTOT. Despite the offset in column pressure away from the setpoint, product compositions are maintained at the specifications.

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The proposed control structure handles all these disturbances and provide

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6. Conclusion

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effective regulatory-level stable closedloop control.

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The use of dual condensers is economically attractive in systems that feature difficult separations, have lighter-than-light-key impurities in the feed that are significantly more volatile that the light-key component and have significant amounts of this component in the feed.

Effective control requires a slight modification of the basic process to provide an

additional control degree of freedom.

References 1. Kister, H. Z. Distillation Design 1992, McGraw-Hill, p. 97. 2. Turton, R., Bailie, R. C., Whiting, W. B., Shaelwitz, J. A. Analysis, Synthesis and Design of Chemical Processes 2nd Edition, 2003, Prentice Hall.

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Table 1 –Economic Results; De-Isobutanizer with 2 mol% C2 in Feed and 5 mol% Impurities of Key Components

Pressure

(atm)

7.735

6.5

6.8

QR TR AR QC TC AC DV ID Qrefrig Trefrig Arefrig

(MW) (K) (m2) (MW) (K) (m2) (kmol/h) (m) (MW) (K) (m2) (106 $)

10.41 345 208.5 10.17 322 978.5 0 2.891 0 NA 0

9.736 338 180.6 8.698 322 836.9 183.5 2.781 0.9504 311 47.88

9.929 340 190.0 9.379 322 902.4 78.62 2.811 0.4189 305 29.45

Energy Cost

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0.9524 0.8950

2.389

2.436

2.463

0

0.1328

0.0585

0.0367

1.857 2.554

1.824 2.522

1.847 2.495

1.856 2.500

3.173

3.130

3.111

3.119

2.554

(106 $) (106 $/y) (106 $/y)

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0.9524 0.8950

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System Totals Capital Reboiler + Refrig. TAC

10.04 341 192.1 9.620 322 925.6 49.6 2.829 0.2624 301 23.51

0.9412 0.8832

(106 $/y)

Reboiler Energy Refrig. Energy

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0.9809 0.8758

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Shell HX

7.0

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Capital Investment

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One Dual Dual Dual Condenser Condensers Condensers Condensers

15

Page 15 of 29

Table 2 –Economic Results; De-Isobutanizer with 2 mol% C2 in Feed and 2 mol% Impurities of Key Components in Products

Energy Cost

Total Capital Reboiler + Refrig. TAC

13.195 340.3 250.6 12.58 322 1210 90.41 3.327 0.4777 306 31.15

us

cr

ip t

14.40 345.8 290.7 14.15 322 1361 0 3.512 0 NA 0

1.139 1.069 0.2406 0.0308

3.533 0

3.237 0.0667

(106 $) (106 $/y)

2.5521 3.533

2.4797 3.304

(106 $/y)

4.3837

4.1307

(106 $/y)

Ac ce p

Reboiler Energy Refrig. Energy

6.8

1.207 1.086 0.2591 0

te

Shell HX Drum1 Drum2

d

Capital Investment

7.785

an

QR TR AR QC TC AC DV ID Qrefrig Trefrig Arefrig

(atm) (MW) (K) (m2) (MW) (K) (m2) (kmol/h) (m) (MW) (K) (m2) (106 $)

Dual Condensers

M

Pressure

One Condenser

16

Page 16 of 29

Figure Captions Figure 1 – Conventional water-cooled condenser; De-Isobutanizer

ip t

Figure 2 – Dual condenser flowsheet; De-isobutanizer Figure 3 – Dual condenser control structure Figure 4 – Feed flowrate increases

cr

Figure 5 – Feed flowrate decreases

us

Figure 6 – Feed composition disturbances

Ac ce p

te

d

M

an

Figure 7 – CW inlet temperature disturbances

17

Page 17 of 29

Ac ce p

te

d

M

an

us

cr

ip t

Highlights  Using one water-cooled and one refrigerated condenser shown to be economical in some systems  The reduction in pressure results in lower energy consumption in columns with difficult separations.  An effective dynamic control structure is developed and tested.

18

Page 18 of 29

i

Figure(s)

us

cr

Figure 1 – Conventional Water-Cooled Condenser; De-Isobutanizer

M an

10.17 MW (CW)

R = 1805 kmol/h

Feed 500 kmol/h

ed

322 K 0.02 C2 0.49 iC4 0.49 nC4

ce pt

25

Ac

RR=7.074 ID=2.89 m

8.136 atm 345 K

7.735 atm 322 K

Distillate 255 kmol/h 0.0392 C2 0.9127 iC4 0.0481 nC4

50

10.41 MW (LP Steam)

Bottoms 245 kmol/h 0.05 iC4 0.095 nC4

Page 19 of 29

i

us

cr

Figure 2 – Dual Condenser Flowsheet; De-Isobutanizer

0.4189 MW (Chilled Water)

M an

9.378 MW (CW)

V = 78.61 kmol/h

322 K 0.02 C2 0.49 iC4 0.49 nC4

50

Ac

RR=6.79 ID=2.811 m

ce pt

25

7.20 atm 340 K

6.8 atm 322 K

0.0916 C2 0.8701 iC4 0.0383 nC4

ed

R = 1731 kmol/h

Feed 500 kmol/h

6.8 atm 305 K

L = 176.4 kmol/h 0.0159 C2 0.9317 iC4 0.0524 nC4

9.929 MW (LP Steam)

Distillate 255 kmol/h 0.0392 C2 0.9127 iC4 0.0481 nC4

Bottoms 245 kmol/h 0.05 iC4 0.095 nC4

Page 20 of 29

25

i Chilled Water

PC

PC LC

ce pt

ed

LC

M an

CW

FC

cr

us

Figure 3 – Dual Condenser Control Structure

CC

Ac

50

LC

CC

Page 21 of 29

cr

i 5.5

2

4

6

DV (kmol/h)

+10%

600

2

4

8

+15%

550 500 0

6

+10% 2

4

Time (h)

6

8

+15%

200 0

R (kmol/h)

0

400

8

+15%

5 4.5

ed

0

ce pt

Prefrig (atm)

+10%

Ac

P (atm)

6.8 6.7

F (kmol/h)

+15%

Frefrig (Mmol/h)

10/15% Feed Flowrate Increase 6.9

M an

us

Figure 4 – Feed Flowrate Increase

0

+10% 2

20

4

6

8

6

8

6

8

+15%

10 0

+10% 0

2

4

2800

+15% +10%

2600 2400 0

2

4

Time (h) Page 22 of 29

i cr

+10% 4

6

+15%

+10% 0

1.5 1 0.5

2

4

0

2

6

+15% +10%

4

Time (h)

6

8

+10%

3200

2

4

6

8

6

8

6

8

200 +10% +15%

100 0

8

QC (MW)

2

+15%

3400

3000 0

8

ed

2

ce pt

xD (%nC4)

0

2.1

1.9

Qrefrig (MW)

QR (MW)

2.5

3600

DL (kmol/h)

+15%

2

M an

3

Ac

xB (%iC4)

us

Figure 4 – continued

0

2

14

4

+15%

13.5 +10%

13 12.5 0

2

4

Time (h) Page 23 of 29

cr

i 4

6

ce pt

8 -15%

6

-10%

0

500

2

4

450 400 0

6

8

R (kmol/h)

4

-10%

-15% 2

4

Time (h)

6

8

50 0

8

ed

2

Ac

F (kmol/h)

Prefrig (atm)

6.75 0

DV (kmol/h)

P (atm)

6.8

100

Frefrig (Mmol/h)

10/15% Feed Flowrate Decrease 6.85

M an

us

Figure 5 – Feed Flowrate Decrease

0

-10% 2

4

6

8

2

4

6

8

6

8

20 10 0

0

2500 -10% -15% 2000 0

2

4

Time (h) Page 24 of 29

i cr QR (MW)

4

0

1

2

ce pt

2

8

4

6

8

0

2

4

Time (h)

6

8

-10%

3000

2500 0

-15% 2

4

6

8

2

4

6

8

6

8

250 200 150 0

QC (MW)

0.5 0

6

Ac

xD (%nC4)

2

2.5

1.5

Qrefrig (MW)

0

3500

DL (kmol/h)

2 1

M an

3

ed

xB (%iC4)

us

Figure 5 – continued

13 -10% 12 -15% 11

0

2

4

Time (h)

Page 25 of 29

i cr

6

5.5

250

2

4

6

8

3% C2

1% C2 200 0

2

4

Time (h)

50 0

R (kmol/h)

300

ce pt

4.5 0

1% C2

6

8

3% C2

100

8

3% C2 5

150

Frefrig (Mmol/h)

4

ed

2

Ac

DTOT (kmol/h)

Prefrig (atm)

6.75 0

M an

P (atm)

6.8

DV (kmol/h)

C2 Composition Disturbances 6.85

us

Figure 6 – Feed C2 Composition Disturbances

1% C2 2

20

4

6

8

4

6

8

4

6

8

3% C2

10 1% C2

0

0

2

2600

1% C2

2400 3% C2 2200 0

2

Time (h) Page 26 of 29

i cr QR (MW)

2

0

6

2

4

6

QC (MW)

1

3% C2

0.5

1% C2 0

0

2

4

Time (h)

6

8

3100

3% C2 2

4

6

8

4

6

8

4

6

8

250 1% C2

200 150 100 0

8

1% C2

3200

3000 0

8

ed

4

ce pt

xD (%nC4)

2

2.5

1.5

Qrefrig (MW)

0

3300

DL (kmol/h)

2 1.5

M an

2.5

Ac

xB (%iC4)

us

Figure 6 – continued

3% C2 2

13.5 13

1% C2

12.5

3% C2 12 0

2

Time (h) Page 27 of 29

i cr

us

Figure 7 – CW Inlet Temperature Disturbances

7

+3 4

600 400

2

4

6

8

R (kmol/h)

0

200 0

0

2

0

8

+3

4

Time (h)

6

8

+3

200

Frefrig (Mmol/h)

5

ce pt

3

6

6

ed

2

7

4

DTOT (kmol/h)

0

Ac

Prefrig (atm)

6.5

400

M an

P (atm)

3

DV (kmol/h)

CW Temp Disturbances 7.5

3

0

2

4

6

8

2

4

6

8

4

6

8

20 10 +3 0

0

2800 2600

3

2400 +3 2200 0

2

Time (h) Page 28 of 29

i cr us

6

+3

2 3 1

4

0

2 1

2

4

0 -1 0

8

6

2800 0

+3 2

4

6

8

4

6

8

4

6

8

400 3 200

0

8

+3 0

2

16 +3

3 2

3000

QC (MW)

3

2

ed

0

ce pt

xD (%nC4)

QR (MW)

+3

3

3200

DL (kmol/h)

2

1

Qrefrig (MW)

3400

3

M an

3

Ac

xB (%iC4)

Figure 7 – continued

4

Time (h)

3 14 12 +3

6

8

10 0

2

Time (h) Page 29 of 29