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Design and control of distillation columns with inert venting
Journal Pre-proof
Design and Control of Distillation Columns with Inert Venting William L. Luyben PII: DOI: Reference:
S0098-1354(19)31341-9 https://doi.org/10.1016/j.compchemeng.2020.106725 CACE 106725
To appear in:
Computers and Chemical Engineering
Received date: Revised date: Accepted date:
15 December 2019 7 January 2020 8 January 2020
Please cite this article as: William L. Luyben , tion Columns with Inert Venting, Computers and https://doi.org/10.1016/j.compchemeng.2020.106725
Design and Control of DistillaChemical Engineering (2020), doi:
This is a PDF file of an article that has undergone enhancements after acceptance, such as the addition of a cover page and metadata, and formatting for readability, but it is not yet the definitive version of record. This version will undergo additional copyediting, typesetting and review before it is published in its final form, but we are providing this version to give early visibility of the article. 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. © 2020 Published by Elsevier Ltd.
Highlights " Inert components in the feed to a distillation require venting a vapor stream from the top of the reflux drum. " Valuable products are inevitably lost in the vent stream. " Vent losses can be reduced by increasing pressure, but this increases energy cost. " An economic balance finds the optimum pressure.. " An effective control structure uses a reflux-drum temperature controller to adjust the vent flowrate.
1
Paper submitted to Computers and Chemical Engineering
Design and Control of Distillation Columns with Inert Venting William L. Luyben Department of Chemical Engineering Lehigh University Bethlehem, PA 18015
December 15, 2019 Revised January 7, 2020
[email protected]; 610-758-4256; FAX 610-758-5057 2
Abstract The feed streams in many distillation columns contain small amounts of very light components that must be removed in the overhead system. To avoid having to operate at the high pressures or low temperatures required to totally condense the distillate, a small vapor vent stream is removed from the top of the reflux drum. This paper considers the economic and controllability issues involved in the design and control of these inert venting systems. An important engineering trade-off exists between product losses in the vent and energy consumption leads to an optimum operating pressure, which varies with the concentration of inert in the feed.
Key Words Inert venting; partial condensers; vapor distillate; distillation control
1. Introduction An exceptionally rich literature has developed over the last century in the area of distillation design and distillation control. Thousands of papers and scores of textbooks have appeared dealing with a subject that is quite vast in scope because of the great variety of types of columns and types of vapor-liquid equilibrium relationships. This extensive coverage is not unexpected in light of the importance of distillation in the chemical processing industries. Despite many predictions of imminent demise by farsighted “experts”, distillation remains the pre-eminent separation method in petroleum refineries, chemical plants and energy facilities around the world. Distillation columns are used to separate mixtures of components that differ in volatility. More-volatile, low-boiling light components are removed in the overhead distillate product stream. Less-volatile, higher-boiling heavy components are produced as the bottoms product. The distillate is typically removed from the reflux drum as a liquid product, which can be pumped to whatever pressure is required. This requires that the column operate at the bubble-point temperature and pressure of the given distillate composition. In order to use inexpensive cooling water in the condenser, reflux-drum temperatures are often set at 50 oC. The pressure of the column is then set at the bubble-
3
point pressure of the distillate. Low-boiling components require high column pressures. For example, a C3-Splitter column producing a propylene distillate operates at 20 atm. On the other hand, a methanol/water column can operate at 0.55 atm. If a vapor distillate is desired, the reflux drum operates at the dew-point temperature of the distillate mixture. If the distillate has a high purity, the difference between the bubble-point and dew-point temperatures is small. However, if the distillate is a mixture of components with differing volatilities, this difference can become large and required a much higher pressure when the distillate is liquid (total condenser) than when the distillate is vapor (partial condenser). In many important applications the stream fed to the column contains a small amount of a very low-boiling component. The production of methanol from synthesis gas is a good example. The gaseous mixture of hydrogen, carbon monoxide and carbon dioxide reacts in a vapor phase reactor to form methanol and water. The reactor operates at high temperature (250 oC) and high pressure (10 atm). The reactor effluent is cooled and partially condenses to form some liquid. The mixed-phase stream is flashed to lower pressure and the liquid from the flash drum is fed into a distillation column to separate the methanol from the water. However, the gas stream leaving the flash drum is mostly unreacted light components that are compressed and recycled back to the reactor. Naturally there are small amounts of these very light (almost inert) components that are dissolved in the liquid stream. In the methanol process example the composition of the liquid fed to the column is about 1 mol% CO2, 81 mol% methanol and 18 mol% water. In this paper we begin by exploring the design of these systems to determine the important design optimization variables for developing an economically optimum process. A range of values of the inert composition in the feed is investigated. Then we develop an effective control structure that handles large disturbances in production rate and feed compositions.
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2. Process Studied The numerical example used in this paper has a liquid feed stream that is a mixture of methanol and water with small amounts of carbon dioxide. Inert feed compositions from zero to 5 mol% CO2 are considered. Figure 1 gives the flowsheet for the case with 2 mol% CO2 in the feed. The column has 30 stages with feed introduced on Stage 22. A partial condenser is used with most of the methanol removed in the liquid phase (DL) but there is also a small vapor purge stream (DV) removed from the top of the reflux drum. The reflux-drum temperature is 50 oC and the feed flowrate is 100 kmol/h in all cases. As we will demonstrate in the next section of this paper, the optimum pressure for the 2 mol% CO2 feed composition shown in Figure 1 is 1.03 bar. The vent stream flowrate is 3.945 kmol/h with a composition of 46.53 mol% CO2 and 53.37 mol% methanol. The flowrate of the distillate liquid product is 37.69 kmol/h with a methanol purity of 99 mol%. The loss of methanol in the vent stream is 2.106 kmol/h, which corresponds to a yield loss of 5.3 % of the 40 kmol/h of methanol in the feed (recovery is 94.7 %).. Note that if there were no inert in the feed, there would be no need for a vent purge. The temperature in the reflux drum is 50 oC but the temperature of the vapor leaving the top stage in the column is 65 oC. Simulations use Aspen NRTL physical properties. The column uses a Radfrac model with the heat-transfer option in the condenser selected as “LMTD” (so the flowrate of cooling water is manipulated in the dynamic simulations reported later). The heattransfer option in the reboiler is “Condensing” so the flowrate of the steam to the reboiler is manipulated in the dynamic simulations.
3. Cases Explored Several cases are considered over a range of inert compositions of the feed stream. Product purities are fixed at 99 mol%. Reflux-drum temperature is fixed at 50 oC. Pressure is varied for each case to find the “sweet spot” at which energy cost and the value of the methanol lost in the vent are balanced.
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3.1 Base Case with No Inert: We begin by designing a column with no inert in the feed and therefore no vent purge stream. With a 99 mol% methanol distillate, a total condenser and a 50 oC reflux drum, the operating pressure is 0.55 bar. The reflux ratio needed to also produce a 99 mol% water bottoms is 0.8035, which gives a reboiler duty of 795 kW. The recovery of methanol is 99.5 %. The temperature of the vapor leaving the top of the column is 50.6 oC, which is very close to the reflux-drum temperature. Contrast this case with that shown in Figure 1 when there is 2 mol% CO2 in the feed and the overhead vapor temperature (65 oC) is significantly higher than the reflux-drum temperature. Note that the reboiler duty in no-inert feed is lower (795 kW) than that required in the 2 mol% feed case shown in Figure 1 (852 kW). At the same time the methanol recovery in the no-inert feed is higher (99.5 %) than the 2 mol% case (94.7 %). These results illustrate the adverse effects of having inerts in the feed.
3.2 Feed with 0.5 mol% CO2: With inert components in the feed, we have to find the optimum operating pressure. Logic would lead us to expect that higher pressure would make the methanol/water separation more difficult and energy requirements would increase. Logic would also lead us to expect that higher pressure (at the same 50 oC reflux-drum temperature) would decrease the concentration of the methanol in the vapor vent and increase the concentration of inert. Both of these effects lead to a smaller loss of methanol in the vent. So there is a trade-off between energy cost and the value of the lost methanol with pressure being the design optimization variable. For the economic calculation, the cost of the low-pressure reboiler steam1 is $7.87 per GJ, and the value of the lost methanol2 is $350 per metric ton. Capital costs are derived from Turton et al1. With pressure specified in the simulations, two Aspen design spec vary functions are used to keep the composition of the bottoms at 99 mol% water by varying the bottoms flowrate and the composition of the liquid distillate product at 99 mol% methanol by varying the reflux ratio. Figure 2 and Table 1 gives results for a range of pressures. As pressure is increased, the flowrate of the vent stream DV and its methanol concentration yD both decrease, which rapidly reduces the loss of methanol in the vent.
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The reboiler duty and the energy cost, as well as the annual cost of capital, increase slowly as pressure increases. Similar trends occur for the 2 mol% case as shown in Figure 3. However, note that the flowrate of the vapor distillate DV is much larger. Since almost all of the CO2 fed into the column must get out in vent stream, this result is no surprise. The question now to be answered is what is the optimum pressure?
3.3 Optimum Pressure: Capital is a fixed-cost one-time investment. Energy and product loss are operating expenditures that continue as long as the plant is running. Therefore it seems reasonable to select the optimum operating pressure as the point where energy costs are just balanced by the savings in product loss. We might intuitively expect that the pressure where the vent loss versus curve and energy versus curve intersect is the optimum operating pressure for the given inert feed composition case. It turns out that this is a good approximation. What we really need to look at to find the rigorous optimum pressure is the profit produced by the column. Profit is the value of the on-spec liquid distillate DL minus the sum of the value of the methanol lost in the vent DV and the total annual cost of energy and capital. Figure 4 gives economic results for several feed compositions. Four curves are shown in each plot: cost of vent loss, cost of energy, annual cost of capital and profit. All curves have units of K$ per year. The profit curve is divided by ten so it can be displayed on the same graph as the less costly expenses. Capital cost is annualized by using a payback period of 3 years. The optimum pressure is where the profit reaches a maximum. In all cases the profit curve becomes quite flat soon after the point of intersection of the vent loss curve and the energy curves. So we assume this intersection point gives a quick and easy approximation of the rigorous economic pressure. The results shown in Figure 4 cover inert feed compositions from 0.5 mol% CO2 to 5 mol% CO2 The intersection point between energy and vent loss shifts to higher pressures as the composition of inert CO2 in the feed increases. Note that energy and capital costs do not change much with feed composition, but the vent losses are strong functions of the amount of inert in the feed since all of the CO2 in the feed must go out in the vent.
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Figure 5 shows how the important process variable change for each feed composition. As the inert feed composition increases, the optimum pressure, the reboiler duty and the flowrate of the vapor vent increase while the methanol concentration of the vent stream decreases. The results presented above demonstrate a logical design procedure for distillation columns with inert venting. In the following section we develop a control structure that provides effective regulatory-level control of these columns.
4. Control Structure The controllability of a distillation column is just as important as its design. Stable operation in the face of inevitable disturbances is necessary for efficient operation in terms of maintaining product quality, minimizing energy cost and avoiding regions of unsafe operation. Simulations are developed using Aspen Dynamics and results for the 2 mol% case are presented. Some discussion of the specific models used may be useful since Aspen provide several alternatives.
4.1 Aspen Models: The column is simulated using a Radfrac model with the type of dynamic heat-transfer options for the condenser and reboiler selected to give realistic manipulated variables. In the condenser the Aspen “LMTD” model is chosen. The temperature of the cooling water is set at 30 oC. The “Temperature approach” parameter is set at 10 oC, which means that the exit cooling water temperature is 10 oC lower than the temperature of the vapor leaving the top of the column. From the known condenser heat duty and the known temperatures of the reflux drum and the overhead vapor, Aspen calculates the required flowrate of the cooling water and the required “UA” for use in the dynamic simulation. For example, the cooling water flowrate is 25,970 kg/h for the 2 mol% case. Then in Aspen Dynamics the flowrate of the cooling can be manipulated. In the reboiler the Aspen “Condensing” model is chosen. From the known reboiler duty, Aspen calculates the required flowrate of steam for use in the dynamic
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simulation. For example, the steam flowrate is 1359 kg/h for the 2 mol% case. Then in Aspen Dynamics the flowrate of the steam can be manipulated.
4.2 Control Structure: Figure 6 gives the Aspen Dynamics flowsheet with a basic control structure consisting of the following loops; 1. Feed is flow controlled. 2. The temperature on Stage 28 is controlled by manipulating the flowrate of the steam. This stage is selected since temperature changes from tray to tray are large at this location. 3. Column pressure is controlled by manipulating the flowrate of the cooling water. 4. The temperature in the reflux drum (Stage 1) is controlled by manipulating power of the compressor. 5. The mass flowrate of the reflux is ratioed to the mass flowrate of the feed. The alternative structure of fixing the reflux ratio gave similar results. 6. Base level is controlled by manipulating the control valve in the bottoms line. 7. Reflux-drum level is controlled by manipulating the control valve in the distillate line. Item 4 in the list above is the key feature in an inert-venting system. For a given pressure, as more inert is accumulated in the reflux drum, the temperature decreases. Venting more material removes more inerts. Therefore the temperature controller must have reverse action (when temperature goes up, too much is being vented, so compressor power should be decreased). Standard flow, level and pressure control parameters are used. The two temperature controllers have 1-minute deadtimes and are tuned by running relayfeedback test and applying Tyreus-Luyben tuning rules. The mass ratio of reflux to feed is 0.563. In some chemical plants manipulation of cooling water flowrate is avoided because of issues with heat-exchanger fouling and corrosion due to high cooling water return temperatures. In this case a control degree of freedom is lost. Pressure is then controlled by manipulating the flowrate of the vapor vent stream. This structure only
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works when there is a significant vent flowrate, which would not be the case for very small inert compositions in the feed.
4.3 Dynamic Results: The control structure is tested by making step disturbances in either throughput and feed composition. Dynamic simulation results for 20% changes in the setpoint of the feed flow controller at 0.5 hours are given in Figure 7. Solid lines are increases in feed; dashed lines are decreases. As expected, an increase in feed flowrate causes increases in the flowrates of steam and cooling water. Both temperature control loops bring the temperatures back to the desired values without about 1 hour. Bottoms water purity (xB) and liquid distillate methanol purity (xD) come to new steadystate values that are close to specification. In Figure 8 changes in feed composition are imposed on the column. Solid lines are when the CO2 concentration of the feed is increased from 2 to 3.5 mol% with the composition water reduced appropriately. Dashed lines are when the CO2 concentration of the feed is decreased from 2 to 0.5 mol% with the composition water increased appropriately. Stable regulatory control is achieved. The flowrate of the vapor vent (DV) changes drastically for these large changes in the inert that must be removed. These very large changes in inert feed composition put a severe rangeability load on the compressor that is pulling the vent stream out of the reflux drum. As the lower left graph in Figure 8 shows, the flowrate of the vent DV changes by a factor of about 7 for changes in inert feed composition from 0.5 to 3.5 mol% CO2. If you look closely at the 20% increase in feed given in Figure 7, you can see a large transient dip in Stage 28 temperature due to the large step change in the cold feed stream. The effect is also felt in the bottom purity, which drops to about 97 mol% for a brief time. Figure 9 illustrates the effectiveness of using a steam-to-feed ratio with the Stage 28 temperature controller resetting the ratio. This combination of feedforward and feedback control works well for attenuating transient deviations. The mass steam-to-feed ratio is 0.563 in the 2 mol% case used in the control study.
5. Conclusion
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The presence of inert components in the feed to a distillation column necessitates the use of inert venting from the top of the reflux drum. There is an inevitable loss of valuable components in this vent stream. For a fixed minimum reflux-drum temperature, these losses can be reduced by operating at higher pressure. However higher pressure normally reduces volatility, which results in higher energy consumption. We have illustrated this important engineering trade-off to find the optimum column pressure. A control structure has been developed that effectively handles large disturbance. The heart of this structure is a reflux-drum temperature controller that adjusts the vent flowrate to prevent the accumulation of inert in the overhead system.
References (1) Turton, R., Bailie, R. C., Whiting, W. B., Shaelwitz, J. A. Analysis, Synthesis and Design of Chemical Processes 2nd Edition, 2003, Prentice Hall. (2) ICIS; https://www.icis.com/explore/commodities/chemicals/methanol/
Figure Captions Figure 1 – Flowsheet; design with 2 mol% CO2 Figure 2 – Effect of pressure with 0.5 mol% CO2 feed composition Figure 3 – Effect of pressure; 2 mol% CO2 feed Figure 4A – Economics; 0.5 mol% CO2 Figure 4B – Economics; 1 mol% CO2 Figure 4C – Economics; 2 mol% CO2 Figure 4D – Economics; 3 mol% CO2 Figure 4E – Economics; 5 mol% CO2 Figure 5 – Effect of feed composition Figure 6 – Aspen Dynamics control structure Figure 7 – 20% feed rate disturbances Figure 8 – Feed CO2 composition disturbances Figure 9 - +20% increase in feed with and without steam-to-feed ratio
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Table 1 – Effect of Design Pressure; 0.5 mol% CO2 Pressure QR (bar)
DV
(kW) (kmol/h)
yD
Vent
Energy
Capital/3
(mf
Loss
(106
(106 $/y)
MeOH) (106 $/y) $/y)
0.6
819
6.015
0.919
0.4939
0.1834
0.2843
0.65
824
3.081
0.846
0.2334
0.186
0.2872
0.68
810
2.411
0.808
0.174
0.181
0.287
0.7
829
2.113
0.785
0.1487
0.188
0.2883
0.8
839
1.324
0.687
0.0816
0.188
0.289
1
859
0.7544
0.550
0.0372
0.192
0.291
1.2
881
0.491
0.458
0.0202
0.197
0.294
12
13
14
15
16
17
18
19
20
21
22
23
24
25
26
Conflict of Interest The author has no conflict of interest. Author Statement No conflict
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Design and Control of Distillation Columns with Inert Venting William L. Luyben PII: DOI: Reference:
S0098-1354(19)31341-9 https://doi.org/10.1016/j.compchemeng.2020.106725 CACE 106725
To appear in:
Computers and Chemical Engineering
Received date: Revised date: Accepted date:
15 December 2019 7 January 2020 8 January 2020
Please cite this article as: William L. Luyben , tion Columns with Inert Venting, Computers and https://doi.org/10.1016/j.compchemeng.2020.106725
Design and Control of DistillaChemical Engineering (2020), doi:
This is a PDF file of an article that has undergone enhancements after acceptance, such as the addition of a cover page and metadata, and formatting for readability, but it is not yet the definitive version of record. This version will undergo additional copyediting, typesetting and review before it is published in its final form, but we are providing this version to give early visibility of the article. 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. © 2020 Published by Elsevier Ltd.
Highlights " Inert components in the feed to a distillation require venting a vapor stream from the top of the reflux drum. " Valuable products are inevitably lost in the vent stream. " Vent losses can be reduced by increasing pressure, but this increases energy cost. " An economic balance finds the optimum pressure.. " An effective control structure uses a reflux-drum temperature controller to adjust the vent flowrate.
1
Paper submitted to Computers and Chemical Engineering
Design and Control of Distillation Columns with Inert Venting William L. Luyben Department of Chemical Engineering Lehigh University Bethlehem, PA 18015
December 15, 2019 Revised January 7, 2020
[email protected]; 610-758-4256; FAX 610-758-5057 2
Abstract The feed streams in many distillation columns contain small amounts of very light components that must be removed in the overhead system. To avoid having to operate at the high pressures or low temperatures required to totally condense the distillate, a small vapor vent stream is removed from the top of the reflux drum. This paper considers the economic and controllability issues involved in the design and control of these inert venting systems. An important engineering trade-off exists between product losses in the vent and energy consumption leads to an optimum operating pressure, which varies with the concentration of inert in the feed.
Key Words Inert venting; partial condensers; vapor distillate; distillation control
1. Introduction An exceptionally rich literature has developed over the last century in the area of distillation design and distillation control. Thousands of papers and scores of textbooks have appeared dealing with a subject that is quite vast in scope because of the great variety of types of columns and types of vapor-liquid equilibrium relationships. This extensive coverage is not unexpected in light of the importance of distillation in the chemical processing industries. Despite many predictions of imminent demise by farsighted “experts”, distillation remains the pre-eminent separation method in petroleum refineries, chemical plants and energy facilities around the world. Distillation columns are used to separate mixtures of components that differ in volatility. More-volatile, low-boiling light components are removed in the overhead distillate product stream. Less-volatile, higher-boiling heavy components are produced as the bottoms product. The distillate is typically removed from the reflux drum as a liquid product, which can be pumped to whatever pressure is required. This requires that the column operate at the bubble-point temperature and pressure of the given distillate composition. In order to use inexpensive cooling water in the condenser, reflux-drum temperatures are often set at 50 oC. The pressure of the column is then set at the bubble-
3
point pressure of the distillate. Low-boiling components require high column pressures. For example, a C3-Splitter column producing a propylene distillate operates at 20 atm. On the other hand, a methanol/water column can operate at 0.55 atm. If a vapor distillate is desired, the reflux drum operates at the dew-point temperature of the distillate mixture. If the distillate has a high purity, the difference between the bubble-point and dew-point temperatures is small. However, if the distillate is a mixture of components with differing volatilities, this difference can become large and required a much higher pressure when the distillate is liquid (total condenser) than when the distillate is vapor (partial condenser). In many important applications the stream fed to the column contains a small amount of a very low-boiling component. The production of methanol from synthesis gas is a good example. The gaseous mixture of hydrogen, carbon monoxide and carbon dioxide reacts in a vapor phase reactor to form methanol and water. The reactor operates at high temperature (250 oC) and high pressure (10 atm). The reactor effluent is cooled and partially condenses to form some liquid. The mixed-phase stream is flashed to lower pressure and the liquid from the flash drum is fed into a distillation column to separate the methanol from the water. However, the gas stream leaving the flash drum is mostly unreacted light components that are compressed and recycled back to the reactor. Naturally there are small amounts of these very light (almost inert) components that are dissolved in the liquid stream. In the methanol process example the composition of the liquid fed to the column is about 1 mol% CO2, 81 mol% methanol and 18 mol% water. In this paper we begin by exploring the design of these systems to determine the important design optimization variables for developing an economically optimum process. A range of values of the inert composition in the feed is investigated. Then we develop an effective control structure that handles large disturbances in production rate and feed compositions.
4
2. Process Studied The numerical example used in this paper has a liquid feed stream that is a mixture of methanol and water with small amounts of carbon dioxide. Inert feed compositions from zero to 5 mol% CO2 are considered. Figure 1 gives the flowsheet for the case with 2 mol% CO2 in the feed. The column has 30 stages with feed introduced on Stage 22. A partial condenser is used with most of the methanol removed in the liquid phase (DL) but there is also a small vapor purge stream (DV) removed from the top of the reflux drum. The reflux-drum temperature is 50 oC and the feed flowrate is 100 kmol/h in all cases. As we will demonstrate in the next section of this paper, the optimum pressure for the 2 mol% CO2 feed composition shown in Figure 1 is 1.03 bar. The vent stream flowrate is 3.945 kmol/h with a composition of 46.53 mol% CO2 and 53.37 mol% methanol. The flowrate of the distillate liquid product is 37.69 kmol/h with a methanol purity of 99 mol%. The loss of methanol in the vent stream is 2.106 kmol/h, which corresponds to a yield loss of 5.3 % of the 40 kmol/h of methanol in the feed (recovery is 94.7 %).. Note that if there were no inert in the feed, there would be no need for a vent purge. The temperature in the reflux drum is 50 oC but the temperature of the vapor leaving the top stage in the column is 65 oC. Simulations use Aspen NRTL physical properties. The column uses a Radfrac model with the heat-transfer option in the condenser selected as “LMTD” (so the flowrate of cooling water is manipulated in the dynamic simulations reported later). The heattransfer option in the reboiler is “Condensing” so the flowrate of the steam to the reboiler is manipulated in the dynamic simulations.
3. Cases Explored Several cases are considered over a range of inert compositions of the feed stream. Product purities are fixed at 99 mol%. Reflux-drum temperature is fixed at 50 oC. Pressure is varied for each case to find the “sweet spot” at which energy cost and the value of the methanol lost in the vent are balanced.
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3.1 Base Case with No Inert: We begin by designing a column with no inert in the feed and therefore no vent purge stream. With a 99 mol% methanol distillate, a total condenser and a 50 oC reflux drum, the operating pressure is 0.55 bar. The reflux ratio needed to also produce a 99 mol% water bottoms is 0.8035, which gives a reboiler duty of 795 kW. The recovery of methanol is 99.5 %. The temperature of the vapor leaving the top of the column is 50.6 oC, which is very close to the reflux-drum temperature. Contrast this case with that shown in Figure 1 when there is 2 mol% CO2 in the feed and the overhead vapor temperature (65 oC) is significantly higher than the reflux-drum temperature. Note that the reboiler duty in no-inert feed is lower (795 kW) than that required in the 2 mol% feed case shown in Figure 1 (852 kW). At the same time the methanol recovery in the no-inert feed is higher (99.5 %) than the 2 mol% case (94.7 %). These results illustrate the adverse effects of having inerts in the feed.
3.2 Feed with 0.5 mol% CO2: With inert components in the feed, we have to find the optimum operating pressure. Logic would lead us to expect that higher pressure would make the methanol/water separation more difficult and energy requirements would increase. Logic would also lead us to expect that higher pressure (at the same 50 oC reflux-drum temperature) would decrease the concentration of the methanol in the vapor vent and increase the concentration of inert. Both of these effects lead to a smaller loss of methanol in the vent. So there is a trade-off between energy cost and the value of the lost methanol with pressure being the design optimization variable. For the economic calculation, the cost of the low-pressure reboiler steam1 is $7.87 per GJ, and the value of the lost methanol2 is $350 per metric ton. Capital costs are derived from Turton et al1. With pressure specified in the simulations, two Aspen design spec vary functions are used to keep the composition of the bottoms at 99 mol% water by varying the bottoms flowrate and the composition of the liquid distillate product at 99 mol% methanol by varying the reflux ratio. Figure 2 and Table 1 gives results for a range of pressures. As pressure is increased, the flowrate of the vent stream DV and its methanol concentration yD both decrease, which rapidly reduces the loss of methanol in the vent.
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The reboiler duty and the energy cost, as well as the annual cost of capital, increase slowly as pressure increases. Similar trends occur for the 2 mol% case as shown in Figure 3. However, note that the flowrate of the vapor distillate DV is much larger. Since almost all of the CO2 fed into the column must get out in vent stream, this result is no surprise. The question now to be answered is what is the optimum pressure?
3.3 Optimum Pressure: Capital is a fixed-cost one-time investment. Energy and product loss are operating expenditures that continue as long as the plant is running. Therefore it seems reasonable to select the optimum operating pressure as the point where energy costs are just balanced by the savings in product loss. We might intuitively expect that the pressure where the vent loss versus curve and energy versus curve intersect is the optimum operating pressure for the given inert feed composition case. It turns out that this is a good approximation. What we really need to look at to find the rigorous optimum pressure is the profit produced by the column. Profit is the value of the on-spec liquid distillate DL minus the sum of the value of the methanol lost in the vent DV and the total annual cost of energy and capital. Figure 4 gives economic results for several feed compositions. Four curves are shown in each plot: cost of vent loss, cost of energy, annual cost of capital and profit. All curves have units of K$ per year. The profit curve is divided by ten so it can be displayed on the same graph as the less costly expenses. Capital cost is annualized by using a payback period of 3 years. The optimum pressure is where the profit reaches a maximum. In all cases the profit curve becomes quite flat soon after the point of intersection of the vent loss curve and the energy curves. So we assume this intersection point gives a quick and easy approximation of the rigorous economic pressure. The results shown in Figure 4 cover inert feed compositions from 0.5 mol% CO2 to 5 mol% CO2 The intersection point between energy and vent loss shifts to higher pressures as the composition of inert CO2 in the feed increases. Note that energy and capital costs do not change much with feed composition, but the vent losses are strong functions of the amount of inert in the feed since all of the CO2 in the feed must go out in the vent.
7
Figure 5 shows how the important process variable change for each feed composition. As the inert feed composition increases, the optimum pressure, the reboiler duty and the flowrate of the vapor vent increase while the methanol concentration of the vent stream decreases. The results presented above demonstrate a logical design procedure for distillation columns with inert venting. In the following section we develop a control structure that provides effective regulatory-level control of these columns.
4. Control Structure The controllability of a distillation column is just as important as its design. Stable operation in the face of inevitable disturbances is necessary for efficient operation in terms of maintaining product quality, minimizing energy cost and avoiding regions of unsafe operation. Simulations are developed using Aspen Dynamics and results for the 2 mol% case are presented. Some discussion of the specific models used may be useful since Aspen provide several alternatives.
4.1 Aspen Models: The column is simulated using a Radfrac model with the type of dynamic heat-transfer options for the condenser and reboiler selected to give realistic manipulated variables. In the condenser the Aspen “LMTD” model is chosen. The temperature of the cooling water is set at 30 oC. The “Temperature approach” parameter is set at 10 oC, which means that the exit cooling water temperature is 10 oC lower than the temperature of the vapor leaving the top of the column. From the known condenser heat duty and the known temperatures of the reflux drum and the overhead vapor, Aspen calculates the required flowrate of the cooling water and the required “UA” for use in the dynamic simulation. For example, the cooling water flowrate is 25,970 kg/h for the 2 mol% case. Then in Aspen Dynamics the flowrate of the cooling can be manipulated. In the reboiler the Aspen “Condensing” model is chosen. From the known reboiler duty, Aspen calculates the required flowrate of steam for use in the dynamic
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simulation. For example, the steam flowrate is 1359 kg/h for the 2 mol% case. Then in Aspen Dynamics the flowrate of the steam can be manipulated.
4.2 Control Structure: Figure 6 gives the Aspen Dynamics flowsheet with a basic control structure consisting of the following loops; 1. Feed is flow controlled. 2. The temperature on Stage 28 is controlled by manipulating the flowrate of the steam. This stage is selected since temperature changes from tray to tray are large at this location. 3. Column pressure is controlled by manipulating the flowrate of the cooling water. 4. The temperature in the reflux drum (Stage 1) is controlled by manipulating power of the compressor. 5. The mass flowrate of the reflux is ratioed to the mass flowrate of the feed. The alternative structure of fixing the reflux ratio gave similar results. 6. Base level is controlled by manipulating the control valve in the bottoms line. 7. Reflux-drum level is controlled by manipulating the control valve in the distillate line. Item 4 in the list above is the key feature in an inert-venting system. For a given pressure, as more inert is accumulated in the reflux drum, the temperature decreases. Venting more material removes more inerts. Therefore the temperature controller must have reverse action (when temperature goes up, too much is being vented, so compressor power should be decreased). Standard flow, level and pressure control parameters are used. The two temperature controllers have 1-minute deadtimes and are tuned by running relayfeedback test and applying Tyreus-Luyben tuning rules. The mass ratio of reflux to feed is 0.563. In some chemical plants manipulation of cooling water flowrate is avoided because of issues with heat-exchanger fouling and corrosion due to high cooling water return temperatures. In this case a control degree of freedom is lost. Pressure is then controlled by manipulating the flowrate of the vapor vent stream. This structure only
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works when there is a significant vent flowrate, which would not be the case for very small inert compositions in the feed.
4.3 Dynamic Results: The control structure is tested by making step disturbances in either throughput and feed composition. Dynamic simulation results for 20% changes in the setpoint of the feed flow controller at 0.5 hours are given in Figure 7. Solid lines are increases in feed; dashed lines are decreases. As expected, an increase in feed flowrate causes increases in the flowrates of steam and cooling water. Both temperature control loops bring the temperatures back to the desired values without about 1 hour. Bottoms water purity (xB) and liquid distillate methanol purity (xD) come to new steadystate values that are close to specification. In Figure 8 changes in feed composition are imposed on the column. Solid lines are when the CO2 concentration of the feed is increased from 2 to 3.5 mol% with the composition water reduced appropriately. Dashed lines are when the CO2 concentration of the feed is decreased from 2 to 0.5 mol% with the composition water increased appropriately. Stable regulatory control is achieved. The flowrate of the vapor vent (DV) changes drastically for these large changes in the inert that must be removed. These very large changes in inert feed composition put a severe rangeability load on the compressor that is pulling the vent stream out of the reflux drum. As the lower left graph in Figure 8 shows, the flowrate of the vent DV changes by a factor of about 7 for changes in inert feed composition from 0.5 to 3.5 mol% CO2. If you look closely at the 20% increase in feed given in Figure 7, you can see a large transient dip in Stage 28 temperature due to the large step change in the cold feed stream. The effect is also felt in the bottom purity, which drops to about 97 mol% for a brief time. Figure 9 illustrates the effectiveness of using a steam-to-feed ratio with the Stage 28 temperature controller resetting the ratio. This combination of feedforward and feedback control works well for attenuating transient deviations. The mass steam-to-feed ratio is 0.563 in the 2 mol% case used in the control study.
5. Conclusion
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The presence of inert components in the feed to a distillation column necessitates the use of inert venting from the top of the reflux drum. There is an inevitable loss of valuable components in this vent stream. For a fixed minimum reflux-drum temperature, these losses can be reduced by operating at higher pressure. However higher pressure normally reduces volatility, which results in higher energy consumption. We have illustrated this important engineering trade-off to find the optimum column pressure. A control structure has been developed that effectively handles large disturbance. The heart of this structure is a reflux-drum temperature controller that adjusts the vent flowrate to prevent the accumulation of inert in the overhead system.
References (1) Turton, R., Bailie, R. C., Whiting, W. B., Shaelwitz, J. A. Analysis, Synthesis and Design of Chemical Processes 2nd Edition, 2003, Prentice Hall. (2) ICIS; https://www.icis.com/explore/commodities/chemicals/methanol/
Figure Captions Figure 1 – Flowsheet; design with 2 mol% CO2 Figure 2 – Effect of pressure with 0.5 mol% CO2 feed composition Figure 3 – Effect of pressure; 2 mol% CO2 feed Figure 4A – Economics; 0.5 mol% CO2 Figure 4B – Economics; 1 mol% CO2 Figure 4C – Economics; 2 mol% CO2 Figure 4D – Economics; 3 mol% CO2 Figure 4E – Economics; 5 mol% CO2 Figure 5 – Effect of feed composition Figure 6 – Aspen Dynamics control structure Figure 7 – 20% feed rate disturbances Figure 8 – Feed CO2 composition disturbances Figure 9 - +20% increase in feed with and without steam-to-feed ratio
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Table 1 – Effect of Design Pressure; 0.5 mol% CO2 Pressure QR (bar)
DV
(kW) (kmol/h)
yD
Vent
Energy
Capital/3
(mf
Loss
(106
(106 $/y)
MeOH) (106 $/y) $/y)
0.6
819
6.015
0.919
0.4939
0.1834
0.2843
0.65
824
3.081
0.846
0.2334
0.186
0.2872
0.68
810
2.411
0.808
0.174
0.181
0.287
0.7
829
2.113
0.785
0.1487
0.188
0.2883
0.8
839
1.324
0.687
0.0816
0.188
0.289
1
859
0.7544
0.550
0.0372
0.192
0.291
1.2
881
0.491
0.458
0.0202
0.197
0.294
12
13
14
15
16
17
18
19
20
21
22
23
24
25
26
Conflict of Interest The author has no conflict of interest. Author Statement No conflict
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