Vacuum and High-Pressure Distillation

CHAPTER

Vacuum and High-Pressure Distillation

9
Zarko Olujic

Process and Energy Laboratory, Delft University of Technology, Delft, The Netherlands

CHAPTER OUTLINE 9.1 Introduction ..................................................................................................... 295 9.2 Operating pressure ranges and selection criteria ............................................... 298 9.3 Pressure effects ............................................................................................... 301 9.3.1 Stage and reflux requirement........................................................... 301 9.3.2 Physical properties ......................................................................... 304 9.4 Column performance considerations.................................................................. 307 9.4.1 Column capacity ............................................................................ 307 9.4.2 Column efficiency........................................................................... 311 9.4.3 Pressure drop................................................................................. 312 9.5 Equipment design considerations ...................................................................... 313 9.5.1 Trays ............................................................................................. 313 9.5.2 Packings........................................................................................ 315 9.6 Concluding remarks and outlook ....................................................................... 316 References ............................................................................................................. 317

9.1 Introduction By virtue of its nature, continuous multistage distillation involves handling of liquid and vapor mixtures, with corresponding bubble and dew point lines enclosing the distillation space (a coexisting saturated vapor and liquid region). The bubble and dew point temperatures tend to increase (or decrease) with increasing (or decreasing) pressure, while the distance between the dew and boiling-point lines corresponds approximately to the difference in the boiling-point temperatures of two components (in the case of binary mixtures) or two key components (in the case of multicomponent mixtures). Figure 9.1 shows a schematic drawing of a standard, one feedetwo products, tray distillation column, including all ancillary equipment (i.e., a total condenser, a partial kettle-type reboiler, and a reflux accumulator drum, as well as two pumps needed to feed the reboiler and transport the reflux to the top of the column, respectively). Upon startup, the initial amount of the liquid phase is generated by condensing the saturated overhead vapor and returning the required quantity of the liquid (reflux) Distillation: Equipment and Processes. http://dx.doi.org/10.1016/B978-0-12-386878-7.00009-7 Copyright © 2014 Elsevier Inc. All rights reserved.

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FIGURE 9.1 Schematic Representation of a Tray Distillation Column with Ancillary Equipment

into the top of the column. The vapor phase (boil-up), which is introduced into the column at the bottom, is obtained by evaporation of the required quantity of the boiling-point liquid leaving the column at the bottom. An additional amount of liquid and/or vapor is introduced as the feed along the column. Usually it is a single feed, as shown schematically in Figure 9.1, but distillation columns can have a number of feed points receiving liquid and/or vapor. Most frequently, a single feed is a saturated liquid, but also it can be a saturated vapor or a two-phase mixture, and in some cases also a subcooled liquid or an overheated vapor. Distillation columns can have one or more side product draw-offs, mainly as saturated liquid, but also it is possible to arrange a saturated vapor draw-off. Inside a column, from the top to the bottom, there is a continuous transfer of less volatile components from saturated vapor to saturated liquid by condensation and more volatile components from the saturated liquid into the saturated vapor phase by evaporation. If the components of the mixture do not significantly differ in their molar enthalpy of vaporization, the amounts of liquid and vapor generated at each stage will be equivalent (constant molar overflow or equimolar mass transfer). In

9.1 Introduction

case of unequal molar enthalpy of vaporization, as is usually the case when distilling aqueous mixtures, the internal liquid and vapor traffic will be different and needs to be accounted for by including enthalpy balance into stage requirement calculations to arrive at a proper design. The lowest pressure and temperature are established at the reflux accumulator placed below the condenser. In contrast, the reboiler is the place with the highest pressure and temperature for the given system. The difference in top and bottom temperatures may go from 10 K for a close boiling system to hundreds of K as encountered in crude oiledistilling columns. The overhead vapor temperature needs to comply with the availability and temperature of the cooling media, with water and air being the preferred choices. This will set the condenser temperature, and the top of the column temperature is per definition the dew point temperature of the overhead vapor corresponding to the top of the column pressure. This top pressure is the reflux accumulator pressure enlarged by the pressure drop generated by the condenser and the vapor line between the top of the column and the condenser. The top of the column pressure enlarged by the pressure drop of column internals under operation yields the bottom pressure, and the corresponding temperature is the bubble point of the liquid in the column sump. The bottom temperature should generally be below that causing thermal degradation of the material distilled. If the bottom temperature is too high, the operating pressure needs to be reduced accordingly. Therefore, dealing with high-boiling components implies generally an operation under vacuum conditions. The heavier the components are, the deeper the vacuum will be. This means less and less tolerance for pressure drop. In critical cases, an effective measure is to place the condenser inside the top of the column. Indeed, the amount of tolerable pressure drop is an important design consideration and affects the equipment choice and internal configuration of a column. While the pressure drop is just a concern in dealing with above-atmospheric applications, it is a design and operating parameter of primary importance for vacuum applications. When considering a vacuum application, trays are the least preferred choice, followed by random packings. Owing to their relatively lowest pressure drop per stage, structured packings are the preferred choice for distillation under vacuum. Being most effective in this respect, well-established wire gauze packings are found in multistage distillation columns operating with top pressures of around 0.1 bar or 100 mbar (10 kPa). Below this pressure, distillation is performed using wetted-wall (falling film type) columns. For applications at pressures as low as 1 mbar 100 Pa, a “thin-film” or “wiped-wall” column is used. This is a wetted-wall column employing a rotating blade that keeps the film thickness controlled. Under highest vacuum (i.e., at an absolute pressure of a few Pascal), special, molecular, or short-path distillation equipment is used. The present chapter focuses on conventional multistage distillation applications. It addresses and discusses criteria for the selection of the operating pressure, the effects of pressure on separation by distillation, and the overall performance of vapore liquid contacting equipment used in distillation.

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9.2 Operating pressure ranges and selection criteria Design of a distillation column starts with setting the design pressure. In general, a distinction is made between columns designed to operate at above atmospheric pressures and the columns designed to operate under vacuum. The present section addresses the design pressure from a process design standpoint, while distillation columns (cylindrical pressure vessels) operated at vacuum and above-atmospheric pressure conditions have to comply with related mechanical design requirements. This also implies considerations related to the maximum allowable temperature, to guarantee the preservation of mechanical integrity of the chosen construction material under all operating conditions and anticipated extreme situations. One should note that building in large safety margins is expensive and that mechanical engineers also strive for economic designs. The constraints in this respect for industrial distillation applications and a thorough analysis of unfavorable and favorable pressure effects as well as related optimization considerations can be found in books by Kister [1] and by Stichlmair and Fair [2]. Specific features of high-pressure distillation are addressed in detail by Brierley [3]. Atmospheric pressure is a natural choice, but it may be approached in practice to the extent depending on the boiling range of the system to be distilled in conjunction with the availability and costs of cooling and heating media, as well as the thermal degradation sensitivity of the bottom product. Practically, the overhead product temperature should be set high enough to perform condensing in an economic way using the cheapest available cooling media. Water, if readily available, is a standard choice, and the overhead vapor temperature should be 10 to 20 K above the anticipated highest (worst-case) water temperature for the given location or source. If water is scarce or too expensive, the air is used as a cooling medium. The air temperature is subjected to more extreme fluctuations; therefore, a higher temperature approach needs to be considered. If refrigeration is inevitable, then a much smaller temperature approach is chosen. This is done in conjunction with special compact heat exchangers exhibiting a very large heat transfer area per volume. The most sophisticated devices of this kind are employed in the cryogenic distillation of air. On the other side, steam is the most common source of heat used in reboilers. Low-pressure (lowest temperature and cost) steam is a preferred choice. Medium- and high-pressure steam is available at sites with integrated power generation, and it comes at correspondingly higher prices. Hot oil and fired heaters are more expensive options. Where appropriate, hot process fluids are used. A common heat integration practice is to use bottom product to preheat the feed stream. Table 9.1 shows the atmospheric pressure (normal) boiling points of a number of commonly distilled inorganic and organic chemicals. At atmospheric pressure, several important inorganic and light hydrocarbon components have a boiling point well below 0
C (273.15 K). For instance, to get the propylene to the level of temperature suitable for water to be used as cooling medium, an absolute pressure of about 19$bar (19 105 Pa) is required at the top of the column.

9.2 Operating pressure ranges and selection criteria

Table 9.1 Atmospheric Boiling Points (Rounded) of Some Frequently Distilled Inorganic and Organic Chemicals Component

Boiling Point (K at 105 Pa)

Nitrogen (N2) Oxygen (O2) Water (H2O) Methane (CH4) Ethylene (C2H4) Ethane (C2H6) Ethanol (C2H6O) Etylene glycol (C2H6O2) Propylene (C3H6) Propane (C3H8) Isopropanol (C3H8O) Glycerol (C3H8O3) n-Butane (C4H10) i-Butane (C4H10) Pentane (C5H12) Chlorobenzene (C6H5Cl) Hexane (C6H14) Benzene (C6H6) Cyclohexane (C6H12) Cyclohexanone (C6H10O) Heptane (C7H16) Styrene (C8H8) Ethylbenzene (C8H10) Paraxylene (C8H10) Orthoxylene (C8H10)

77.4 (195.8
C) 90.2 (183
C) 373.2 (100
C) 111.2 (162
C) 169.2 (104
C) 184.2 (89
C) 351.2 (78
C) 470.4 (197.2
C) 225.5 (47.7
C) 231.1 (42.1
C) 355.4 (82.2
C) 563.2 (290
C) 272.7 (0.5
C) 261.2 (12
C) 309.2 (36
C) 404.9 (131.7
C) 341.9 (68.7
C) 353.3 (80.1
C) 353.9 (80.7
C) 428.8 (155.6
C) 371.2 (98
C) 418.3 (145.1
C) 409.4 (136.2
C) 411.6 (138.4
C) 417.7 (144.5
C)

Table 9.2 shows typical operating pressures of water- or air-cooled distillation columns. Approximately 35 bar (35$105 Pa) is the upper limit, which is not exceeded in practice because it would require the adoption of unaffordable shell thickness as set by high-pressure vessel design codes. For this reason, demethanizers and deethanizers usually operate at the highest reasonable pressure and use an appropriate refrigerant as cooling medium. In these and other so-called cold-box separations, special refrigeration systems are used, which increases the costs of the separation significantly. The coldest application is the cryogenic distillation of air (temperatures close to 90 K or 183
C), which, depending on the products used (argon, nitrogen, and/or oxygen) and the purity requirements, is carried out at atmospheric and/or at elevated pressures (up to 10 bar or 10$105 Pa).

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Table 9.2 Typical Operating Pressures and Efficiencies of Some Important Tray Column Applications Application Dementanizer De-ethanizer Depropanizer PPP splitter EE splitter Debutanizer De-isobutanizer Benzene–toluene Methanol–water Ethylbenzene–styrene

Pressure (Pa) 5

32$10 Pa (32 bar) 27$105 Pa (27 bar) 19$105 Pa (19 bar) 19$105 Pa (19 bar) 16$105 Pa (16 bar) 5$105 Pa (5 bar) 5$105 Pa (5 bar) 105 Pa (1 bar) 105 Pa (1 bar) 0.2$105 Pa (0.2 bar)

Efficiency (%) 50 60 85 85 85 90 85 75 65 75

Thermodynamically, the upper limit for separation of a component from a mixture by distillation is the critical pressure (i.e., critical temperature). There is no thermodynamic or mechanical limitation on the lower end (columns subject to vacuum usually are designed for full vacuum), where the amount of tolerable pressure drop is a critical design consideration. A typical example of the difficulties encountered with the choice of the operating (vacuum) pressure is a separation of ethylbenzene and styrene, with a polymer-grade styrene monomer as the bottom product. Styrene is a highly reactive molecule that tends to polymerize strongly at temperatures above 363 K. In the past, sieve or valve trays have been used for styrene distillation. Relying on water as cooling medium, the typical pressure at the top of the column was not exceeding 70 mbar (0.07 bar or 7 kPa). To keep the pressure drop low enough, a typical single-shell column usually contained 70 trays and was designed at around 50% of flooding to reduce total pressure drop to an acceptable level. Even with a pressure drop as low as 3 mbar (300 Pa) per tray, the bottom pressure was still around 300 mbar (0.3 bar or 30 kPa), resulting in a bottom temperature of about 378e383 K. At this temperature, the polymerization was a real threat and the only practical remedy was to use a costly polymerization inhibitor, which had to be removed in a subsequent makeup distillation operation. To reduce related costs from adopting the same column design (i.e., pressure drop), an option was to reduce the column’s top pressure by avoiding the pressure drop associated with transport of the vapor from the top of the column to the condenser. Indeed, this was achieved by installing an air-cooled condenser at the top of the column, allowing a lower column top pressure (i.e., a gain of some 25 mbar or 2.5 kPa with respect to conventional water-cooled condensers placed outside the column). It should be realized that such a “small” pressure drop reduction was so significant in this case that a different design approach was chosen

9.3 Pressure effects

involving additional complexities on the mechanical-engineering (construction) side and related costs. It should also be noted that placing condensers, both air and water cooled, at or in the top of the column is an effective engineering solution for demanding vacuum distillations. Regarding the ethylbenzene recovery column mentioned here, due to large vapor density differences as imposed in trayed columns by a pressure drop of about 250 mbar (25 kPa), the rectification section diameter was considerably larger than the stripping section diameter. For example, with feed rates above 30 t/h and high reflux ratios (6e8), the diameter of the rectification section of the trayed columns with top pressures of about 40 mbar (0.04 bar or 4 kPa) has been close to 10 m [4]. A real technology breakthrough in this respect occurred upon implementation of structured packings in these applications (in the late 1970s and early 1980s), which resulted in such a large reduction in the column pressure drop that the bottom temperature dropped below that causing polymerization. Now, there is no more need for using and recovering the antipolymerization additive. This is a specific, highly rewarding additional advantage of the application of corrugated-sheet structured packings in this particular application. A common benefit of a reduction in bottom temperatures is the increase of relative volatility, which often allows a reduction in the operating reflux ratio and consequently leads to reduced investment and/or operating costs. However, the vapor density at the bottom conditions is practically halved, which means dealing with much larger volumetric flows of vapor. This appeared to be a limiting factor in revamps, but in new designs it allows the design of single-diameter packed columns. To avoid extremely large diameters in these applications, two options are considered: increasing the number of installed stages to reduce the operating reflux ratio, and/or increasing the operating pressure accordingly. Both options are possible because of a rather low total pressure drop involved. Certainly, the operating pressure and column pressure drop are two important design considerations because they directly influence the separation via vaporeliquid equilibrium and the performance of the chosen vaporeliquid contacting device.

9.3 Pressure effects 9.3.1 Stage and reflux requirement Relative volatility is a direct measure of the difficulty of separation of a mixture by distillation. Practically, it is expressed as the ratio of equilibrium constants (K-values), which in the case of an ideal mixture reduces to the ratio of saturated vapor pressures of more and less volatile components. For multicomponent mixtures, this is the ratio of light to heavy key components. Figure 9.2 shows atmospheric pressure equilibrium curves for some common test systems. Methanolewater (M/W) is a nonideal system, with relative volatility increasing from 2 at the upper end to nearly 7 at the lower end. Chlorobenzenee ethylbenzene (CB/EB) is an ideal system with a constant relative volatility

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1,0

0,8

0,6

y [-]

302

0,4 M/W,

1.013 bar

CB/EB, 1.013 bar C6/C7, 0.17 bar

0,2

C6/C7, 0.33 bar C6/C7, 1.03 bar C6/C7, 4.14 bar

0,0 0,0

0,2

0,4

0,6

0,8

1,0

x [-]

FIGURE 9.2 Equilibrium Curves of Different Test Systems (1 bar [ 1.105 Pa)

(a ¼ 1.13) over the whole range of compositions. The cyclohexaneen-heptane system (C6/C7) is nearly ideal and is frequently employed for testing the performance of trays and packings at above and below atmospheric pressures [5]. Corresponding relative volatilities are given in Table 9.3, together with other relevant properties of this test system. When the relative volatility becomes 1 over the whole range of compositions, the equilibrium curve covers the diagonal (y ¼ x) and the distillation space diminishes. However, in practice there are many mixtures exhibiting nonideal behavior, in which the equilibrium curve crosses the diagonal at a certain composition, and this point (i.e., interception of the equilibrium line and the diagonal in a McCabeeThiele plot) is generally known as the azeotropic point. Ethanolewater (E/W) is a wellknown example of an azeotropic system. A relative volatility of 1 implies that there is no separation possible by distillation at this condition. However, in the composition range below and above the azeotropic point, it is possible to distill, and in both cases one product stream will have composition corresponding to that of the azeotropic point. Indeed, the azeotropic point is a thermodynamic barrier for separation by distillation, which, however, in some cases can be alleviated by changing the operating pressure. Fundamental and practical aspects of the vaporeliquid equilibrium and azeotropic mixtures separation can be found in Chapters 6 and 7, respectively. As indicated in the yex plot shown in Figure 9.2, for the C6/C7 system, assumed to behave ideally, the relative volatility decreases with increasing pressure. In general, this is true as long as the vapor pressure is the controlling factor in the

9.3 Pressure effects

Table 9.3 Pressure Effect on Relevant Properties of the Cyclohexane–n-Heptane (C6/C7) Mixture (Average at bottom Temperature) Used as a Total Reflux Distillation Test System at the Separations Research Program of University of Texas at Austin, Texas, USA Pressure (Pa) Property

0.17$105

0.33$105

1.03$105

4.14$105

Average temperature, K Liquid density, kg/m3 Liquid viscosity, Pa s Liquid diffusivity, m2/s Vapor density, kg/m3 Vapor viscosity, Pa s Vapor diffusivity, m2/s Surface tension, N/m Relative volatility Slope of equilibrium line Liquid load, m3/m2h, at F-factor ¼ 2 m/s (kg/m3)0.5

322 (49
C) 659 4.67 E4 2.31 E9 0.66 6.67 E6 13.30 E6 0.018 1.94 1.54 8.9

334 (61
C) 657 4.31 E4 2.72 E9 1.19 6.94 E6 11.40 E6 0.017 1.86 1.50 12.0

370 (97
C) 625 2.97 E4 4.44 E9 3.53 7.78 E6 4.17 E6 0.014 1.64 1.35 21.6

427 (154
C) 561 1.61 E4 9.17 E9 13.14 9.17 E6 1.39 E6 0.008 1.42 1.32 46.5

equilibrium behavior of mixtures. Practically, this means that the equilibrium line comes closer to the diagonal, which indirectly indicates that more stages are required in a McCabeeThiele plot to meet the same top and bottom product specifications. This also means that a larger minimum reflux ratio and consequently a larger operating reflux ratio are required. Inevitably, more stages and a higher reflux will increase equipment and operating costs. To initialize column design calculations, an assumption is made regarding the column pressure drop. When the number of equilibrium or ideal stages (theoretical plates) at optimum reflux is fixed, these stages are converted to actual stages (i.e., trays or a certain packed height) using the overall tray (column) efficiency or the height equivalent to a theoretical plate (HETP), respectively. For trays of known configuration as well as chosen packing type and size, reliable estimates of column pressure drop can be obtained using well-established predictive models. Adding the estimated pressure drop to the top pressure represents new column bottom pressure, and a new iteration is performed by calculating the corresponding number of stages and reflux ratio. Repeating calculations ends when subsequent values are within a given tolerance. In high-pressure columns, a very high absolute pressure drop (about 1 bar or 105 Pa) may be relatively small (i.e., less than 10% of the operating pressure drop). Typical examples are so-called propyleneepropane (PP-) splitters, in which such a high absolute pressure drop is usually not a concern from a fluid dynamics point of view but, if underestimated, could be highly detrimental to thermodynamics (i.e., vaporeliquid equilibrium). In other words, a rather small change in the relative

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volatility of a close boiling system (a < 1.2) strongly influences both the number of stages and the reflux ratio. This is illustrated in the following example. Considering a typical debutanizer column in an ethylene plant, the column separates propane and other lighter components from butane and heavier components. For this column, the relative volatility can be treated as constant because an error of about 10% (e.g., taking 2.5 instead of 2.25) could lead to an underestimate of the number of equilibrium stages for a stage or two, which is usually within the design margin. However, if we consider the next column in the separation sequence, a PPsplitter with polymer-grade product purity, then an assumption of constant relative volatility is not valid anymore. In this case, due to the pressure drop and corresponding temperature increase, the relative volatility decreases from the top to the bottom (e.g., from 1.1 to 1.07). Making the same error in the relative volatility would lead to a deficit in the number of equilibrium stages larger than that corresponding to the total number of stages encountered in the debutanizer column. In other words, some situations tolerate a certain degree of design sloppiness and/or equilibrium data inaccuracy, which, however, in other situations could have catastrophic consequences. Therefore, particular care and the utmost precision are required when dealing with design or rating of the columns with close boiling systems. However, if a heat pump (vapor recompression) assisted design is considered, a much lower operating pressure can be chosen. For a PP-splitter, this means that the operating pressure can be reduced from 19 bar (19$105 Pa) to 10 bar (10$105 Pa), which leads to a significant increase in relative volatility and consequently in much lower stages and reflux ratio requirements. This resulted in practice in a significant added benefit (i.e., placing all the required stages in one shell instead of using two columns connected in series).

9.3.2 Physical properties Stage and reflux requirement calculations rely on the availability and accuracy of vaporeliquid equilibrium data, which can be found in the form of numerous models and related parameters in commercial process simulation packages. The next step in column design, column sizing or dimensioning, requires knowledge of a number of relevant physical (transport) properties (i.e., densities, viscosities, and diffusivities) of two phases as well as surface tension at process conditions (pressure and temperature profiles) as established along the column. Therefore, distillation column designers, as well as operators, must be aware of the nature and extent of the effects of the pressure on all relevant physical properties of both phases, to be able to understand what happens or could occur in a system under consideration. Table 9.4 shows the effect of increasing pressure on important physical properties affecting the design and operation of a distillation column. This information, based on a complex graphical plot from Lockett’s book [6], can be used as a qualitative guide. Detailed calculations need to be performed to quantify properly the sensitivity of each variable to the changes caused by variations in the operating pressure (see Table 9.3 for the C6/C7 system at four operating pressures).

9.3 Pressure effects

Table 9.4 Trends in Physical Properties Changes of Saturated Liquids and Vapors with Increasing Pressure Property

Trend

Temperature Liquid density Liquid viscosity Liquid diffusivity Surface tension Vapor density Vapor viscosity Vapor diffusivity Enthalpy of vaporization

Increases Decreases Decreases Increases Decreases Increases Increases Decreases Decreases

Pressure has no significant effect on density of the liquid in distillation operations. Since the temperature of a saturated liquid increases with increasing pressure, the liquid density will decrease slightly. The density of a vapor increases nearly proportionally to pressure increase. Therefore, the vapor density is an important dimensioning variable directly influencing the capacity of existing columns and the choice of the column diameter in new designs. The liquid viscosity, which is related to liquid diffusivity and was used to directly correlate tray efficiency, tends to decrease with increasing pressure, while the diffusivity increases at approximately the same rate. The hydrocarbons distilled at high pressure are the lightest ones and therefore the absolute values of the viscosity are rather low, while vacuum distillation deals with heavy components having a relatively much higher viscosity. According to the well-known O’Connell tray efficiency correlation [7]: EOC ¼ 51  32:5$logða  hÞ

(9.1)

where EOC (%) is the column efficiency, a () is the average relative volatility of the light key component, and h (cp or mPa s) is the molar average dynamic viscosity of liquid feed mixture; an increase in viscosity leads to a decrease in tray efficiency, which is in fact a consequence of a corresponding decrease in liquid diffusivity. The same occurs with increasing relative volatility, which means that with increasing pressure (i.e., decreasing liquid viscosity and decreasing relative volatility), the tray efficiency tends to increase. Although not including surface tension, the O’Connell correlation provides a good indication of the effect of processing conditions on (a generic) tray efficiency and tends to give conservative estimates. Interestingly, the vapor viscosity increases slowly with the increasing pressure. A similar but opposite trend is with the vapor diffusivity. This fully compensating effect is the reason that the Schmidt number of the vapor is practically insensitive to the increase in the operating pressure, while the liquid Schmidt number shows a

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pronounced decreasing trend. In other words, as mentioned here, the liquid density also decreases with increasing pressure, but this is not as pronounced as in the case of viscosity. The vapor diffusivity tends to decrease while the liquid diffusivity tends to increase with increasing pressure. However, the effect of the liquid and vapor diffusivities on the mass transfer efficiency is much less than that of the interfacial area changes caused by a significant change in the operating pressure. For the trays operating in mixed or froth regime and emulsion regime, this is directly related to the size and the residence time of bubbles. For packings, an increased specific liquid load leads generally to an increase in the effective area. However, if the liquid load is too high, it exhibits a deteriorating effect on packing efficiency, and this is certainly the case if the specific surface area of chosen packing is not large enough to ensure liquid film flow. Surface tension of hydrocarbons decreases with increasing pressure. At high pressures, the surface tension can be very low. The size of bubbles, as appearing in froth and emulsion regimes, tends to decrease with increasing pressure. This leads to an increase in interfacial area and vaporeliquid contact time. That is the main reason why the tray efficiency in high-pressure distillations is generally higher than that observed in applications around atmospheric pressure where a mixedfroth (larger bubbles) regime prevails. However, low surface tension and high pressure may make bubbles so small that they cannot escape from liquid, thus getting saturated (reaching the equilibrium) and not participating anymore in mass transfer. Even worse, vapor may be entrained by liquid flowing into the downcomer and entered into the tray below. This is usually called vapor backmixing. This similarly occurs with small liquid droplets in low-pressure systems, which, entrained from liquid, are taken by the vapor and hit the tray above, where they get partly aspirated (i.e., returned to a tray they left before). This kind of liquid and/or vapor recirculation (backmixing) is detrimental to the efficiency and depends on the amount of backmixing. Unfortunately, pressure effects on relevant physical properties of the liquid and the vapor, including some partly compensating effects, are difficult to quantify properly and may be influenced to a different extent by tray hydraulics. Without a proper understanding of underlying phenomena, in our attempts to quantify tray efficiency properly, we continue to rely upon empirical or semiempirical correlations, based on data obtained from total reflux distillation experiments conducted with wellestablished test systems. Therefore, the best sources of efficiency data are similar applications involving the same or similar trays. These, however, are not available in the public domain. A rough idea about the efficiencies experienced in various services can be obtained from Table 9.2, which contains the kind of generic information found in open literature (see Chapter 2), where also typical operating pressures are indicated. It should be noted that, similar to densities, the diffusivities of the liquid and vapor phases come closer to each other at the highest pressure distillation. Thermodynamically, the critical pressure is the upper operating limit for distillation, because at this condition two phases lose their distinctive characteristics (i.e., become a supercritical fluid).

9.4 Column performance considerations

9.4 Column performance considerations Above- and below-atmospheric situations differ considerably with respect to the effect of pressure on vapor density. For the PP-splitter case, a pressure drop of around 1 bar (105 Pa) does not affect the vapor flow significantly. On the other hand, the column pressure drop of 250 mbar (0.25 bar or 0.25 105 Pa), as encountered in a trayed ethylbenzeneestyrene column [4], is much smaller in the absolute value; however, it is nearly four times larger than the column top absolute pressure. Consequently, the vapor density at the bottom is about four times larger than that at the column top. A practical consequence was that the diameter of the rectifying section of these columns had to be much larger than that of the stripping section. New ethylbenzene recovery columns, equipped with structured packings, exhibit a rather low pressure drop (less than 100 mbar or 0.1 105 Pa) and, consequently, all are designed as one diameter column. Since the liquid density is less affected by the pressure, a practical consequence is that vacuum columns operate with a relatively low specific liquid load. In other words, since the column diameter depends strongly on the vapor density, the increasing operating pressure results in the same internal mass flow in a smaller column diameter accompanied by a correspondingly increased specific liquid load. Expressed as superficial liquid velocity per hour (m3/m2h), the specific liquid load of distillation columns varies generally from 1 m3/m2h in low vacuum to nearly 70 m3/m2h in high-pressure applications.

9.4.1 Column capacity Liquid and vapor densities in conjunction with superficial vapor velocity determine the operating range of a distillation column and, most importantly, normalize it in a way that limits the whole operating range to a relatively narrow range of absolute numbers, which are easy to remember and place accordingly. Figure 9.3 shows a

FIGURE 9.3 Capacity Factor Correlation for Sieve Trays

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so-called capacity plot for sieve tray columns, with the capacity factor, CV (m/s), as a function of the flow parameter 4LV (), with the tray spacing as a parameter. These two important design and operation parameters are defined as: rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi rV (9.2) CV ¼ uV;max $ rL  rV And: 4LV ¼

rffiffiffiffiffiffi rffiffiffiffiffiffi rffiffiffiffiffiffi M_ L V_ L rV rL uL rL $ $ $ ¼ ¼ rL rV uV rV M_ V V_ V

(9.3)

where uV,max (m/s) is the maximum allowable (flooding) superficial vapor velocity; rV (kg/m3) and rL (kg/m3) are densities; M_ V (kg/s) and M_ L (kg/s) are mass flow rates; V_ V (m3/s) and V_ L (m3/s) are volume flow rates; and uV (m/s) and uL (m/s) are superficial velocities of the vapor and the liquid, respectively. The capacity factor is effectively the maximum allowable vapor velocity, and the curves shown in Figure 9.3 represent the corresponding flooding line at different tray spacings. Larger tray spacing allows a higher vapor velocity (i.e., a leaner column or more capacity). At the lower end of the operating flow parameter range, this line coincides with the so-called entrainment or jet flooding line. This implies flooding caused by excessive entrainment. This type of flooding dominates in typical distillation applications. At the high flow parameter end, which is characteristic of highpressure distillation, downcomer and/or system flood is usually a limiting factor. A thorough discussion of the system limit (i.e., the ultimate capacity of distillation columns) can be found in a recent paper by Fitz and Kunesh [8]. Similar capacity factorevseflow parameter plots are available for both random and structured packing columns, with the packing size as a parameter (see Chapters 3 and 4). The flow parameter (see Eqn (9.3)) represents the ratio of the squares of kinetic energies or momentum fluxes of the two phases. At total reflux conditions, the molar flow rates of two phases are equal, and in many, mainly close boiling distillations, their ratio is close to unity, which means that the flow parameter is practically proportional to the square of the ratio of vapor and liquid densities. Another frequently used measure of column capacity is the so-called F-factor, FV. It represents the square of the vapor kinetic energy (i.e., a product of the superficial vapor velocity and the square root of vapor density) and is related to the capacity factor by the square root of the density difference of two phases: pffiffiffiffiffiffi pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi F V ¼ uV $ rV ¼ C V $ rL  rV (9.4) Due to a strong compensating effect (i.e., a decrease in the superficial vapor velocity accompanied by an increase in vapor density with increasing pressure), the whole operating range of both tray and packed columns varies within a relatively narrow range of F-factor values, such as between 1 and 4 m/s (kg/m3)0.5 or (Pa0.5). In general, the operating F-factor increases with decreasing pressure. At

9.4 Column performance considerations

the same operating conditions, a larger F-factor means more capacity, and vice versa. Superficial velocity of vapor or the liquid is the average empty column velocity, resulting from the volumetric vapor flow rate divided by the column cross-sectional area. This simple manifestation of the (continuity) law of mass conservation, applied to closed conduits, is the basis for column diameter determination. sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4$M_ V 4$M_ V ¼ DC ¼ (9.5) pffiffiffiffiffiffi rV $0:8$uv;max 0:8$FV;max $ rV Here, factor 0.8 represents a typical design value, implying that the operating vapor velocity is placed 20% below the anticipated flooding limit. One should note that Eqn (9.5) takes the empty column cross-sectional area as its basis, while designers prefer to base the column diameter on the so-called net area or, more conservatively, on the active (bubbling) area [7]. For a conventional crossflow tray, the former is the total column cross-sectional area minus the downcomer area, and the latter is the total column cross-sectional area minus two times the downcomer area. In other words, depending on the chosen capacity factor correlation, the maximum superficial vapor velocity uV,max (m/s) in the denominator of Eqn (9.5) should be multiplied by either (1  (ADC/AC)) or (1  (2ADC/AC)). Here, ADC (m2) and AC (m2) represent the downcomer area and the column crosssectional area, respectively. In both cases, the column diameter will be correspondingly larger than that based on the empty column cross-sectional area, to the extent depending on the chosen downcomer area. This in turn depends on the specific liquid load (i.e., the operating pressure). In cases where pressure drop is critical and there is no other option to reduce it to affordable levels, the design point may be chosen to be 0.5 of the flooding, resulting in a considerably larger column diameter. However, when designing large-diameter columns, tighter design is the preferred choice to reduce high capital costs. Indeed, it requires some daring for column designers to take 0.85 or even 0.9 of the flooding as the design value. Equation (9.5) directly suggests that the capacity of an existing column can be increased by increasing the operating pressure, which is often done in practice but within a limited range, that is, a margin on the upper side that accounts for normal operation pressure fluctuations and certainly should not be pushed beyond the value that would force the pressure relief valve to open. One should note that in some cases, even a small increase in pressure could have detrimental effects on separation due to a significant decrease in relative volatility. Sometimes, a certain tolerable loss in separation quality is acceptable to achieve a temporary profitable increase in capacity. Another potential problem with increasing column pressure is the increase of bottom temperature (i.e., the possibility that the existing reboiler could be insufficient to perform accordingly at a reduced temperature driving force). This, however, is partially compensated by a positive effect of increased pressure (i.e., a corresponding

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decrease of the evaporation enthalpy, which is reflected in a certain reduction of reboiler duty). Using an appropriate capacity plot, which can be extended to account for the effects of surface tension, the liquid viscosity, and/or contactor geometry, the corresponding flooding velocity can be obtained from Eqn (9.2). Critical locations are the top and bottom stages in each column section (i.e., those of the top and the bottom of the column and the stages above and below the feed stage). Normally, one diameter design is preferred, but if the column diameters at the critical locations differ considerably, two- or three-diameter column designs can be considered and arranged if necessary. Smaller diameter differences can be smoothed out by choosing different tray spacing or tray design, or different size of a packing. Certainly, packings can be combined with trays and any combination of internals is possible if it brings a benefit to the application under consideration. Since a column is basically a large tube, the variations in velocities and/or diameters can easily be quantified using another simple manifestation of the continuity law: uV1A1 ¼ uV2A2. Here, A (m2) represents the cross-sectional area, which in the case of trays can be that of the shell, which serves as the basis for determination of the superficial (linear) vapor velocity, or active (bubbling) area, or free area. The active area is the total area minus the area occupied by downcomers, whereas the free area is that available for passage of the vapor through the trays. The above expression, with the column cross-sectional area and superficial velocity fixed, allows the direct estimation of effective or local (hole) velocities based on the chosen value of the bubbling or free area. Another practical equation is related to the vapor and liquid loads at total reflux conditions: uVrV ¼ uLrL, which allows simple calculation of the superficial liquid velocity from the given F-factor values: uL ¼ uV ðrV =rL Þ ¼ FV $r0:5 V =rL . Equipment performance is usually tested under total reflux distillation conditions, which means that with increasing F-factor, both the vapor load and the liquid load increase accordingly. Table 9.3 shows the extent of the increase in the superficial liquid velocity (specific liquid load) with increasing operating pressure, as experienced in total reflux distillation experiments carried out with the C6/C7 system. The trend of the maximum tray capacity curves shown in Figure 9.3 indicates that in the range of flow parameter values from 0.01 to 0.1, covering vacuum to atmospheric-pressure operations, the capacity curves tend to flatten or even decrease at the lowest end. This indicates that in this region of column loads, the capacity is insensitive to an increase in flow parameter. However, above a flow parameter of 0.1, there is a decreasing capacity trend with increasing flow parameter. Here, high liquid loads, which require installation of the larger downcomers at the cost of active area, are accompanied by a high vapor density in conjunction with a low liquid density and a strongly decreasing surface tension (light hydrocarbons), indeed a combination of the properties that can be detrimental (in structured packings) or beneficial (in trays) for mass transfer performance.

9.4 Column performance considerations

9.4.2 Column efficiency At this point, it is interesting to mention that with increasing flow parameter (i.e., operating pressure), the flow regime on a sieve tray changes from a spray flow regime at the lowest end, via a froth (mixed) regime in the middle part of the plot, to an emulsion regime at the higher end of the flow parameter. This is indicated in the capacity factor plot shown in Figure 9.3, and these three major types of vaporeliquid interface are schematically illustrated in Figure 9.4. In other words, under vacuum conditions, trays operate mainly in the spray (drop) regime; around the atmospheric pressure, the mixed-froth regime prevails; and a high-pressure operation is usually associated with the emulsion regime. The latter is essentially a bubble flow regime with the liquid as a continuous phase, characterized by the presence of a relatively large number of small bubbles creating a large interfacial area. The other extreme in this respect is the spray (drop) or jet regime, where liquid loads and weir heights are at the lowest end and vapor velocity at the highest end. Here, the vapor is in continuous phase and the interfacial area is the surface area of dispersed droplets. With increasing relative vapor load (the ratio of the operating F-factor and the maximum F-factor), there is a progressive reduction in the size of the droplets, resulting in a strong increase of the interfacial area; this, however, is always much lower than that occurring in the bubble flow regime. It should be noted here that there are some complex proprietary devices that maximize the capacity by allowing nearly all the liquid from a tray to be entrained and consequently separated from vapor in axial cyclones placed above the tray deck [9]. This means that a relatively very large interfacial area is created; however, the residence time of the vapor is so short that this must be at some expense of efficiency. Certainly, in the spray regime the high vapor velocity is responsible for a partial breakage and dispersion of the liquid stream into droplets. However, at high Vapor continuous

Spray

Froth Emulsion

Vapor load

Droplets Bubbles

Liquid continuous

Liquid load FIGURE 9.4 Schematic Illustration of Flow Regimes on a Sieve Tray Depending on Vapor and Liquid Loads (For color version of this figure, the reader is referred to the online version of this book.)

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pressures (i.e., high liquid loads and high aerated liquid heights on trays), the relatively much larger liquid cross-flow velocities contribute to the breakage of larger into smaller bubbles. As a result, both the interfacial area and vapor residence time tend to increase, but some compensating effects mentioned here, such as saturation and entrainment of bubbles, usually do not allow a significant enhancement of the mass transfer. The mixed-froth regime lies between the drop and bubble flows and is a very complex case. The simplest modeling approach is to interpolate between models describing the spray regime on one side and the bubble regime on the other side. Certainly, modeling and predicting the efficiency of trays are everlasting challenges, but for nonproprietary sieve trays adequate knowhow can be found in open literature. A thorough description of basic principles and design practices, including working equations for the estimation of vapor and liquid phase mass transfer coefficients as well as the interfacial area, can be found in the book by Stichlmair and Fair [2]. Various well-founded correlations allow observation of potential operating conditions effects. A graph in Ref. [2], showing the effect of increasing relative vapor load on trend and the magnitude of the interfacial area of three typical distilling systems, provides a good quantitative indication of the effect of the surface tension (the size of droplets and bubbles). In summary, the lowest tray efficiency is observed in the spray (drop or jet) regime, whereas the highest efficiency is associated with the bubble flow regime. The mixed-froth regime is somewhere in between. Interestingly, distillation of an ethylbenzeneestyrene system is a typical low-vacuum (around 100 mbar or 10 kPa) operation where the flow conditions are such that the spray regime prevails. However, as mentioned in this chapter, at that time, due to the need to minimize the pressure drop, lower than usual superficial vapor velocities were chosen by increasing column (section) diameters accordingly. Hence, these trays actually operated in the froth regime, achieving a relatively higher efficiency. Figure 9.5 shows the mass transfer efficiency (HETP) and pressure drop curves of Montz-Pak B1-250 measured under total reflux conditions as a function of the F-factor, illustrating the effect of the operating pressure on the performance of a well-established structured packing [5]. This plot suggests that the packing efficiency tends to improve with increasing operating pressure. This, however, is something specific to the test system and the pressure range employed.

9.4.3 Pressure drop The pressure drop is an important design and operating variable. It is usually fixed at the highest affordable value by the designer. High-pressure tray columns have larger downcomers and weir height, which implies both a higher clear liquid and a higher froth layer and consequently a higher pressure drop for the given vapor load. As suggested in Figure 9.5, with increasing pressure (i.e., the specific liquid load), the pressure drop of a structured packing increases and the capacity decreases accordingly. This occurs with all test systems and at the extent of the relative

9.5 Equipment design considerations

FIGURE 9.5 Effect of the Operating Pressure (1 bar [ 1.105 Pa; 1 mbar [ 100 Pa) on the Hydraulic and Mass Transfer Performance of the Montz-Pak B1-250 at Total Reflux (C6/C7, d [ 0.43 m, h y 3.3 m)

capacity gain depends on the vapor density. In general, the pressure drop tends to increase with increasing specific liquid load.

9.5 Equipment design considerations The wall thickness of column shell increases with increasing pressure. The diameter decreases, but a taller column may be necessary. Therefore, the total weight is not increasing proportionally with the pressure, but tall columns with a high lengthto-diameter ratio are often exposed to wind pressure, which means additional wall thickness at the column base. This, in turn, means more weight (i.e., an increased capital cost), which may further increase if additional provisions are needed to fix such a column shell properly. More complex is the situation with the choice between packings and trays as well as with the choices and compromises required to arrive at an optimum design within each of these categories of vaporeliquid contacting devices.

9.5.1 Trays A typical tray is a so-called cross-flow (single-pass) tray, with liquid coming from the downcomer passing under the downcomer apron (clearance area), and flowing over the bubbling (active) area toward the outlet weir. Weir height affects the height of the froth that flows over the weir into the downcomer. An active area is provided with holes (sieve tray), valves, bubble cups, or any other tray types of vaporeliquid contactors.

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Tray dimensions depend to some extent on the operating pressure. Regarding the relative free area (i.e., the area available for vapor flow), it decreases with increasing pressure because of a larger downcomer area, which is needed to handle appropriately large liquid loads. For a high-pressure application, the tray open area can be as low as 5%, which implies that hole (local) velocities are 20 times larger than the superficial vapor velocity, which is rather low. Much higher vapor velocities are encountered in vacuum services, where up to 15% of the open area is available for vapor flow. Due to a relatively larger open area, the hole spacing (pitch) is smaller in vacuum than in high-pressure applications. The hole diameter is not directly affected, which depends more on the fouling tendency of the systems. In general, for the same open area, smaller holes are preferred, because they generate less entrainment and thus are beneficial for capacity, and by generating a larger interfacial area, smaller bubbles may also be beneficial for efficiency. For example, in the case of very clean systems, such as cryogenic air distillation, the hole diameters down to 1 mm are employed in industrial columns. If fouling is a potential threat, then larger hole diameters are chosen. For instance, until the 1980s, ethylbenzeneestyrene distillation columns that were prone to styreneemonomer polymerization-based fouling were equipped with trays containing holes with a 12 mm diameter. Even larger holes are used in some services with severe fouling. Weir length increases with increasing pressure (i.e., liquid load) as well as the weir height. In high-pressure applications, weir heights from 5 to 10 cm are common. Downcomer opening or clearance is slightly lower than the weir height. In vacuum columns, weir heights smaller than 3 cm are preferred. Due to a rather low liquid load, the trays employed in vacuum distillation are usually single-pass trays. To ensure sufficient hydraulic gradient in the large diameters, the slotted-sieve tray design was introduced [6], with specially designed openings in the tray allowing the ascending vapor to direct and push the liquid toward the downcomer. The shape and orientation of these flow-directing devices were optimized to approach nearly plug flow of the liquid and, more importantly, to minimize the strong detrimental effect of the stagnant zones at the sides of a tray, which tends to increase with increasing diameter. In general, with increasing operating pressure, at a certain point the liquid load becomes so high that it cannot be handled properly by a single-pass tray. If the weir loading exceeds 60 m3/h per meter weir length, two or more passes are required to handle the liquid appropriately. Therefore, high-pressure distillations are generally associated with multipass trays. The well-known UOP MD tray is the most prominent representative of the highly effective proprietary equipment of this kind [6]. Due to a rather short liquid flow path length and the fact that both phases are well mixed, the efficiency of multipass or multidowncomer trays is effectively approaching point efficiency. Therefore, multidowncomer trays are generally less efficient than common cross-flow single-pass trays. In practice, where appropriate, this can be compensated by installing more trays with small tray spacing. Although considerable effort has been spent in the past on understanding the mass transfer on trays, knowledge of tray fluid dynamics and its relation to tray

9.5 Equipment design considerations

efficiency is still insufficient. We still get confronted with difficulties when interpreting situations regarding actual tray efficiencies that are lower than point efficiency or larger than 100%. In any case, point efficiencies can be estimated with some confidence and can be used as safe design values for industrial columns, because, as experienced in the practice, carefully designed trays usually exhibit somewhat higher overall efficiencies [2].

9.5.2 Packings The ethylbenzeneestyrene separation by distillation under vacuum is the most typical example of the full utilization of favorable performance characteristics of structured packings. However, with increasing flow parameter (i.e., increasing operating pressure), the relative advantages of structured packings fade away. As experienced in the early years of implementation, it appeared that in some cases even uniformly irrigated structured packings were not able to perform accordingly under elevated (de-isobutanizers) and high-pressure (PP-splitter) distillation conditions. It should be noted, however, that these applications were revamps, undertaken with the idea to have a certain capacity gain. However, in given situations, this was possible only by using structured packings with specific geometric areas of 250 m2/m3 or less, and in some cases a steeper corrugation inclination angle (60
) instead of the common 45
was chosen to maximize capacity gain. Unfortunately, such a choice resulted in insufficient efficiency, and the overwhelming enthusiasm around the overall performance of structured packing cooled down. Nevertheless, this “failure” stimulated a breakthrough in tray design. During the last 15 years, a number of new high-capacity trays have been developed and successfully implemented [9,10]. Dedicated high-pressure testing conducted by Fractionation Research Inc. with the first generation of sheet metal structured packings has indicated the existence of a strong performance deterioration effect in the loading region, a so-called efficiency hump [11]. Afterward, considerable effort was undertaken to understand and eventually solve the problem [12]. Soon, it became clear that high liquid load accompanied by a low density difference and very low surface tension created conditions under which film flow is practically impossible and a high degree of entrainment on both the liquid and the vapor sides was responsible for a pronounced backmixing in both phases. In fact, structured packing is a falling film device, and the surface should be covered by a thin flowing film to utilize the interfacial area and mass transfer kinetics in a favorable way. Confronted with high liquid loads (e.g., in excess of 50 m3/m2h), structured packings with a specific geometric area of 250 m2/m3 simply do not have enough installed area to “absorb” easily detachable low-surface-tension hydrocarbon liquid. Under such conditions, the liquid that flows mainly freely partly chocks the flow channels, which forces vapor to choose other less loaded channels to be able to flow upward. Packing becomes less effective in radial spreading of both the liquid and the vapor. This leads to a decrease in the effective contacting area, and is accompanied by more or less pronounced backmixing in both phases, due to a rather low ratio of liquid and vapor phase densities,

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which results in an appreciable loss of efficiency. This could be alleviated to some extent by more frequent redistribution of the liquid and vapor, but this requires additional height to accommodate devices used for this purpose. Certainly there are some examples of the successful operation of structured packings under high pressures. However, this occurred in less demanding separation services [13]. This also indicates that geometric features of corrugated sheet packings could be of some influence on the capability of packing to handle large liquid loads. In general, common size structured packings tend to perform well if the specific liquid load is smaller than 30 m3/m2h. On the other hand, third-generation random packings, which allow spreading of liquid and vapor laterally, in all directions, proved to be capable of performing reasonably well under high-pressure conditions [14]. The capacity gain of the most efficient Norton packing (IMTP 25) was small compared to that of smaller surface versions, which however exhibits an HETP value of 0.5 m, corresponding approximately to the trays with a tray spacing of 0.6 m. Frequent liquid redistribution (a section occupies 1.5e2 m height) adds significantly to the column shell height. This is most probably the main reason why random packings found only limited application in high-pressure services. Interestingly, better efficiencies were obtained in close boiling than in wide boiling applications. Since all relevant physical properties remain similar in the whole column, low relative volatility is the only one that makes a difference, confirming the well-known experience gained with trays that low relative volatility is favorable for efficiency. In general, structured packings are the best choice at the low flow parameter end (i.e., in vacuum), and trays dominate high-pressure applications. Random packings cover a wide (middle) range of sub- and above-atmospheric pressures, but they bring less capacity gain. The newest development, the so-called fourth-generation random packings, could bring some change to the status quo regarding the application window of random packings [15,16]. Adequate quality of initial distribution and redistribution of the liquid and vapor is a prerequisite for a reasonable performance of large-diameter columns equipped with both random and structured packings, in below- and above-atmospheric pressure services.

9.6 Concluding remarks and outlook Pressure affects distillation in various ways, and these have been addressed explicitly. Owing to their distinctive benefits, structured packings and trays are well established in vacuum and high-pressure applications, respectively. The newest developments indicate that the trays may lose ground in some applications to fourth-generation random packings, which exhibited an increased capacity and efficiency in combination with an affordable pressure drop at higher pressure. The problem is that high-capacity trays

References

cause the relatively highest pressure drop in distillation applications, which makes close boiling separations carried out in trayed columns unsuitable for the implementation of energy-conserving heat pump systems. More capacity, even with the most advanced devices, generally implies more pressure drop. Therefore, striving for more capacity in revamps or reduced diameters in new designs is not in line with globally proclaimed goals of making process industries more sustainable. More detailed information on the performance characteristics of trays, random packings, and structured packings can be found in Chapters 2, 3, and 4, respectively. Chapter 1 gives an overview of standard trays and random and structured packings, and it provides a detailed analysis of relative performances under various conditions as encountered in distillation practice. Equipment performance testing is elaborated in greater detail in another book [17], where also the most important industrial applications (refining, petrochemicals, air distilling, and specialty chemicals) are addressed and discussed in certain detail.

References [1] H.Z. Kister, Distillation Design, McGraw-Hill, New York, 1992. [2] J.G. Stichlmair, J.R. Fair, Distillation Principles and Practice, McGraw-Hill, New York, 1998. [3] R.J.P. Brierley, High-pressure distillation is different, Chem. Eng. Prog. 90 (7) (1994) 68e77. [4] R. Billet, Energieeinsparung bei thermischen Stofftrennverfahren, Hu¨thig Verlag, Heidelberg, 1983.
Olujic, Distillation: high and low pressure distillation, in: I. Wilson, C. Poole, [5] Z. M. Cooke (Eds.), Encyclopedia of Separation Science, Elsevier, Amsterdam, 2007 online update 1. [6] M.J. Lockett, Distillation Tray Fundamentals, Cambridge Press, 1986. [7] R.K. Sinnott, Chemical Engineering Design (Volume 6 in Coulson & Richardson’s Chemical Engineering Series), Fourth ed., Elsevier, Amsterdam, 2005. [8] C.W. Fitz, J.G. Kunesh, Column hydraulics: system limit/ultimate capacity, Chem. Eng. J. 88 (2002) 11e19. [9] P. Wilkinson, E. Vos, G. Konijn, H. Kooijman, G. Mosca, L. Tonon, Distillation trays that operate beyond the limits of gravity by using centrifugal separation, Chem. Eng. Res. Des. 85 (2007) 130e135.
Olujic, M. Jo¨decke, A. Shilkin, G. Schuch, B. Kaibel, Equipment improvement [10] Z. trends in distillation, Chem. Eng. Process. 48 (2009) 1089e1104. [11] C.W.E. Fitz, J.G. Kunesh, A. Shariat, Performance of structured packings in a commercial-scale column at pressures of 0.02-27.6 bar, Ind. Eng. Chem. Res. 38 (1999) 512e518.
Olujic, J.G. Kunesh, Liquid backmixing in structured packing in high [12] F.J. Zuiderweg, Z. pressure distillation, IChemE Symp. Ser. No 142 (2) (1997) 865e872. [13] G.W. Hausch, P.K. Quotson, K.D. Seeger, Hydrocarbon Process. 71 (4) (1992) 67e70.

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[14] R.F. Strigle Jr, Packed Tower Design and Applications: Random and Structured Packings, Gulf Publishing Company, Houston, 1994. [15] M. Schultes, Raschig super-ringda new fourth generation packing offers new advantages, Chem. Eng. Res. Des. 81 (2003) 48e57. [16] I. Nieuwoudt, C. Corio, J. Degarmo, Improvement in random packing performance, Pet. Technol. Q. (PTQ) 4 (2010) 67e75. [17] A. Gorak, H. Schoenmakers, Distillation Operation and Applications, Elsevier.