Separation of Dilute Binary Gases by Simulated-Moving Bed with

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Separation of Dilute Binary Gases by Simulated-Moving Bed with Pressure-Swing Assist: SMB/PSA Processes Kyle P. Kostroski and Phillip C. Wankat* School of Chemical Engineering, Purdue UniVersity, Forney Hall of Chemical Engineering, 480 Stadium Mall DriVe, West Lafayette, Indiana 47907-1283

Hybrid simulated-moving bed/pressure-swing adsorption (SMB/PSA) cycles are developed to achieve the relatively difficult separation of dilute enantiomeric enflurane in a nitrogen carrier gas. SMB/PSA decreases molar desorbent requirements by taking advantage of gas expansion at low pressure and the desorbent (carrier gas) present in the dilute feed. The most practical SMB/PSA configuration is a pair of two-zone SMB/PSA trains operating 180° out of phase to form a continuous (1,1)-(1,1) SMB/PSA process. With no additional nitrogen carrier gas (D/F ) 0), the (1,1)-(1,1) SMB/PSA produced products having a maximum average purity of ∼86% and a maximum average recovery of ∼86%. By manipulating the recycle ratio (RR) and the bed purge ratio (BPR) at D/F ) 0, either the raffinate or extract product could be produced with purity greater than or equal to 95%. To produce purer products simultaneously, pure nitrogen was added as desorbent. A maximum average purity of 99% and a maximum average recovery of 99% were attainable at D/F ) 15.0. Further decreases in molar desorbent requirements were realized by increasing the number of beds per zone and employing bed coupling during purge. Introduction As much as 40-70% of chemical plant capital and operating costs are due to separation technologies.1 Gas separation processes are an integral part of the chemical processing industry and require continued separations research.2 The major current gas separation processes are cryogenic distillation, adsorptionbased processes, and membrane technologies. The choice of which separation method to use often depends on process economics and required capacity. In terms of capacity, adsorption-based gas separations processes usually fall midrange between large-scale cryogenic distillation and small-scale membrane units. Pressure-swing adsorption (PSA) was initially designed to produce pure nonadsorbed carrier gas (i.e., light product) at high pressure.3,4 The classic Skarstrom stripping-type cycle uses a simple four-step process: (i) bed pressurization with feed, (ii) adsorption and production of product at high pressure, (iii) blowdown/depressurization, and (iv) purge with product at low pressure.5,6 Stripping-type cycles5-8 are generally used commercially to produce high purity light product with moderate to low recovery since some product is used as purge. The solute gas (i.e., heavy product) is usually treated as a waste stream. In contrast, rectifying-type PSA cycles that produce high purity solute gas are essentially the inverse of the Skarstrom cycle: feed is at low pressure and purge is at high pressure, thereby producing a high purity heavy product with moderate to low recovery.9,10 More recently, dual-reflux PSA cycles have been developed to produce high purity light and heavy products by combining the feed and purge steps of the stripping- and rectifying-type cycles to achieve binary gas separations.11-14 Unlike PSA, gas-phase simulated-moving bed (SMB) type adsorption processes have not been widely studied or commercialized. In cases when both gas- and liquid-phase SMB processes are feasible for a given separation, liquid-phase * To whom correspondence should be addressed. Tel.: (765) 4940814. Fax: (765) 494-0805. E-mail: [email protected].

operations have been commercialized since they appear to be more economical. The SMB process was originally developed by Broughton et al.16 The SMB, a derivative of the true-moving bed,4 uses a loop of linked adsorption beds (zones) and a series of port switches to completely retain mass transfer zones within the system and continuously produce products. Fresh feed and desorbent are fed to the SMB while raffinate and extract streams are taken off as products: binary separation is achieved. Recent work in this area has focused on SMB equipment designs,17 two-zone SMBs,18 and bioseparation applications such as insulin purification.19 SMBs are also being studied for chiral separations, such as the separation of the R and S enantiomers of the inhalation anesthetic enflurane; this gas-phase separation uses a nitrogen carrier gas.20-22 In this work, new hybrid SMB/PSA processes have been developed to achieve binary gas separations. For a dilute feed, SMB/PSA has two distinct advantages: (i) regeneration at low pressure and (ii) internal generation of desorbent (carrier gas). SMB/PSA takes advantage of gas expansion at low pressure as well as the desorbent (carrier gas) present in the feed stream to decrease overall molar desorbent requirements. In this way, the SMB/PSA concept can be viewed as an SMB with pressureswing assist: it utilizes the pressure swing of PSA and the mass transfer zone retention of the SMB to achieve binary gas separation while using less desorbent than an isobaric gas-phase SMB. The separation of R- and S-enflurane enantiomers in a nitrogen carrier gas was selected as a model system for evaluating the performance of SMB/PSA.20-22 The gas chromatographic simulated-moving bed (GC-SMB) process was configured as a standard open-loop four-zone SMB with two beds per zone (Figure 1).22 Conversion of the fourzone GC-SMB in Figure 1 to an SMB/PSA process is not trivial because the addition of a pressure swing adds process complexity. If the pressures of all four zones are varied, considerable difficulty in operation arises; it becomes cumbersome to manage pressure profiles and flow rates in all four zones. It is easier to vary only the pressure of the regeneration zone. In the resulting closed-loop four-zone SMB/PSA shown in Figure 2, only the

10.1021/ie071000b CCC: $40.75 © 2008 American Chemical Society Published on Web 04/03/2008

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Figure 1. Schematic of the (2,2,2,2) four-zone GC-SMB process.22 Figure 3. Schematic of the (1,1,1,1,1) five-zone SMB/PSA process.

Figure 2. Schematic of the (1,1,1,1) four-zone SMB/PSA process. Figure 4. Schematic of steps in the two-zone SMB/PSA process.

pressure of the extract-producing bed is varied; the remaining three zones continue to circulate material. Desorbent is added at low pressure to the extract-producing zone, while highpressure desorbent can be optionally added to the active loop of the train at high pressure. A schematic of a five-zone SMB/ PSA is shown in Figure 3. The configuration shown is similar to the five-zone “clean-in-place” SMB configurations that have been studied previously for liquid separations.23-25 Regardless of their potential utility, the four- and five-zone SMB/PSA processes in Figures 2 and 3 have a major limitation. Some degrees of freedom are lost due to the removal of the regeneration zone from the active loop. The addition of highpressure desorbent to the active loop may be necessary to better control flow rates and improve separation. If high-pressure desorbent is added, the four- and five-zone SMB/PSA processes do not realize the benefits of a complete pressure swing. Alternatively, the SMB/PSA process can be developed using the two-zone SMB process as the prototype.26 A key facet of the two-zone SMB/PSA process is that both extract- and raffinate-producing zones are regenerated at low pressure. The steps in the resulting two-zone SMB/PSA process are schematically shown in Figure 4. The operating cycle has four steps: (i) feed, (ii) depressurization/production, (iii) regeneration (or purge)/production, and (iv) repressurization with desorbent. Unlike the four- and five-zone SMB/PSA configurations, the

process in Figure 4 is discontinuous and will require a tank to store desorbent (Figure 5). Alternatively, two out-of-phase twozone SMB/PSA trains can be coupled to form a continuous process (Figure 6); interestingly, this configuration resembles a typical Skarstrom process but with an SMB replacing each column. During the feed step in Figure 6, zone I produces desorbent (carrier gas) for subsequent use while zone II uses carrier gas circulation to push any remaining light product into zone I, thereby purifying the heavy product remaining in the zone. Both light and heavy products are produced during the depressurization and purge steps. During the purge step, the amount of purge gas must be carefully controlled to push out the light product (raffinate) without pushing out the heavy product (extract) which is recovered from zone II in the next cycle. Repressurization with desorbent occurs next, followed by a port switch; the cycle then repeats. In order to create a symmetric cycle, the timing must be adjusted when the SMB/PSA process is configured with two interacting trains. The cycle still has the four steps discussed previously; however, the feed step is separated into three substeps that are symmetric to blowdown, purge, and pressurization. During feed substep 1, feed to train 1 is recirculated through the recycle loop while depressurization occurs in train

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Several key process variables used in designing the SMB/ PSA process are

recycle ratio ) RR )

VII,feed VI,feed

bed purge ratio ) BPR )

VI,purge VT,purge

molar desorbent-to-feed ratio ) D/F

Figure 5. Schematic of the two-zone SMB/PSA process in a (1,1)-tank configuration.

2. Next, during feed substep 2, feed to train 1 is recirculated and a portion of the desorbent produced is used for purging train 2. Finally, during feed substep 3, the feed to train 1 is recirculated and a portion of the desorbent produced is used to repressurize train 2. The raffinate and extract products are produced during the blowdown and purge steps while the desorbent is generated for internal use during feed substeps 2 and 3. The base bed configuration for the two-zone SMB/PSA uses one bed per zone and decoupled regeneration (as shown in Figure 4). Both discontinuous (Figure 5) and continuous (Figure 6) operations are investigated. The following performance measures are used:

purity of raffinate species (R-enantiomer) ) yRaffinate R yRaffinate + yRaffinate R S purity of extract species (S-enantiomer) )

(1)

yExtract S yExtract + yExtract R S (2)

recovery of species i in product stream j )

Nij Nifeed

(3)

productivity (mass of enflurane fed/bed volume‚time) ) MWenfyenf feedFfeedtfeed (4) ncolVcoltcycle In this work, the feed flow rate was scaled so that the productivities of the various SMB/PSA processes were equal to that of the GC-SMB process investigated previously in the literature.22 After making the necessary corrections for its openloop configuration, the GC-SMB process can be compared to the SMB/PSA process in terms of purity, recovery, and molar desorbent consumption. In addition to the component purity defined in eqs 1 and 2, an average purity will be calculated as the mean of the raffinate and extract product purities. An analogous average recovery will also be reported.

(5)

(6) (7)

For a given scenario, the fresh feed flow rate to the SMB/ PSA process was fixed. By mass balance, this also fixes the flow rate of gas flowing through the reducing valve in Figure 4. This gas, which is mainly desorbent (N2) separated from the fresh feed, is used for the purge and pressurization steps. With these two flow rates fixed, the recycle ratio (RR) could be manipulated independently. As defined above, RR is a measure of the amount of circulation during the feed step; this circulation is used to remove any light product heel from zone II during that step. There are competing positive and negative effects of changing RR when no desorbent is added to the process. As RR is increased, the flow rates in zones I and II increase and this increased circulation removes more light product from zone II. However, since the flow rate of internally generated desorbent is fixed, as RR increases there is proportionately less desorbent available for purge. If light product breaks through, the internally generated desorbent will become contaminated with light product, which will contaminate the heavy product during the purge step. Oppositely, if RR is decreased, the flow rates in zones I and II decrease, which decreases the purity of the light product. The bed purge ratio (BPR) defines the split of purge gas between zones I and II. Setting BPR too high will overpurge zone I and cause heavy product breakthrough into the light product. However, setting BPR too low will leave too large a heel of light product in the bed, thereby contaminating zone II after the port switch. The molar desorbent-to-feed ratio (D/F) reflects the amount of fresh desorbent added to the process; internally generated desorbent from the feed is not counted in this measure. Overall, these three critical operating variables (RR, BPR, and D/F) interact and have marked effects on the extract and raffinate product purities. The operating conditions used when evaluating the SMB/ PSA processes were chosen based on the enflurane separation literature as well as heuristics.3,20-22 In this process, the feed is assumed to be available at 25 °C and 4 bar (PH) and only a blower is required in the circulation loop to overcome the pressure drop in the beds. For regeneration, a vacuum pump is used to achieve a low pressure of 0.2 bar (PL) to make the purge more effective and reduce the amount desorbent required. Although the use of vacuum purge does add some cost to the SMB/PSA process, it has proved to be economical on the commercial scale. One such application is the separation of air by vacuum-swing adsorption (VSA).27 In its pure component form, enflurane has a vapor pressure of 175 mmHg at 20 °C.28 A dilute feed was used: 0.0026 mole fraction R-enflurane, 0.0026 mole fraction S-enflurane, and 0.9948 mole fraction N2.20-22 The adsorbent material was a chiral stationary phase based on octakis(3-O-butanoyl-2,6-diO-n-pentyl)-γ-cyclodextrin dissolved in polysiloxane SE-54 and coated on Chromosorb particles.22


Figure 6. Schematic of the two-zone SMB/PSA process in a (1,1)-(1,1) configuration. Table 1. Data for Enflurane Model System20-22 parameter

value

adsorbent L D i p Fs Rp kMTC,fluid Ez isotherm parameters R-enflurane S-enflurane Cp,solid ∆Hads R-enflurane S-enflurane hp Ez,T ap PHtop (PHbottom) PLtop (PLbottom) Tfeed

octakis(3-O-butanoyl-2,6-di-O-n-pentyl)-γ-cyclodextrin 80 cm 15 mm 0.4 0.583 0.38 g/cm3 150 µm 18.52 s-1 R- and S-enflurane: 0.06 cm2/s

∂ci ∂ci |z)0 ) 0, | ) 0, Vf|z)0 ) 0 (9a,9b,9c) ∂z ∂z z)L The boundary conditions for the pressurization, adsorption, and purge steps are3

K0 ) 151, T0 ) 302.15 K K0 ) 230, T0 ) 302.15 K 0.921 kJ/(kg‚K)

EZ

-46 295 kJ/kmol -53 211.3 kJ/kmol estimated with kez ) 0.026 W/(m‚K) estimated with ks ) 0.174 W/(m‚K) 200 cm-1 4 bar (3.9 bar) 0.3 bar (0.2 bar) 25 °C

Theory To describe nonisothermal fixed bed adsorption, we need the mass and energy balance equations, the mass and energy transfer equations, and the equilibrium isotherm. The necessary experimentally determined parameters were taken from the literature on enflurane separations (Table 1).20-22 The normal assumptions of negligible radial gradients, no chemical reactions (other than adsorption), and linear driving force for mass transfer are made. The mass balance is29

e

The necessary boundary conditions on concentration and velocity during blowdown are3

∂qi ∂ci ∂ci,pore + Kdi(1 - e)p + Fs(1 - e)(1 - p) + ∂t ∂t ∂t ∂(Vfci) ∂2ci - eEZ 2 ) 0 (8) e ∂z ∂z

∂ci | ) -Vf|z)0(ci|z)0- - ci|z)0) ∂z z)0 ∂ci | )0 ∂z z)L

(ci|z)0-)purge )

PL (c | ) PH i z)L adsorption

(10a)

(10b)

(10c)

The initial conditions for a clean and saturated bed are,

Vf|z)0 ) V0(P), Vf|z)0 ) Vfeed, Vf|z)0 ) Vpurge (11a,11b,11c) respectively3

ci(z,0) ) 0, qi(z,0) ) 0, ci(z,0) ) c0i , qi(z,0) ) q0i (12a,12b,12c,12d) The adsorption beds were assumed to be initially clean and free of both R- and S-enflurane. It was assumed that N2 does not adsorb.20-22 In writing the energy balance, it is assumed that radial gradients are negligible, the adsorption beds are adiabatic, and heat transfer follows a linear driving force model. The energy balance is29

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∂T ∂T h* + Ff Cp,f p(1 - e) + FsCp,s(1 - p)(1 ∂t ∂t ∂Ts ∂(VT) ∂2T + Ff Cp,f e - Ez,TFf Cp,f e 2 ) 0 (13) e) ∂t ∂z ∂z

FfCp,f e

The boundary condition during the blowdown step is3

∂T ∂T | ) 0, | )0 ∂z z)0 ∂z z)L

(14a,14b)

For the pressurization, adsorption, and purge steps, the boundary conditions are3

EZ,T

∂T | ) -Vf|z)0FfCp,f(T|z)0- - T|z)0) ∂z z)0

(15a)

∂T | )0 ∂z z)L

(15b)

(T|z)0-)purge ) (T|z)L)adsorption

(15c)

The initial condition for temperature is3

T(z,0) ) Tfeed

(16)

For mass transfer, the linear lumped parameter mass transfer equation with fluid-phase driving force is29

∂qi ) kMTC,fluid(ci - ci*) ∂t

(17)

The parameter kMTC,fluid in eq 17 is a lumped mass transfer coefficient. The linear lumped parameter energy transfer equation is29

∆Hads ∂q ap ∂Ts (T - Ts) ) hHTC ∂t Cp,fFp f Cp,f ∂t

(18)

The mass and heat transfer coefficients were assumed to be constant. The heat transfer and thermal axial dispersion coefficients were estimated based on the thermal conductivities of the gas and solid phases, kez and ks, respectively, and the Colburn j-factor:30

hHTC ) jCp,f νf FfPr-2/3

(19)

where j ) 1.66Re-0.51 if Re < 190 and j ) 0.983Re-0.41 otherwise. Both linear and nonlinear isotherms for this system were examined previously.20-22 Juza et al.20 concluded in 1998 that the “actual nonlinear complete separation region is similar to the linear one” for the dilute feed concentration of interest. In 2000, Biressi et al.21 stated that “the complete nonlinear isotherms would not help in this case to develop the separation processes.” The pressure range examined in this work also agrees with that examined by Biressi et al. in 2002.22 Based on these considerations, the use of the following linear, temperaturedependent isotherm in the current work appears to be justified:22

{ [ ( )]}

qi ) K0 exp

∆Hads 1 1 R T T0

ci

(20)

Simulations ADSIM, a product of Aspen Technology, Inc., was used to numerically solve the mass and energy balances, mass and

Table 2. Base Cycle Timing for the SMB/PSA Process step 1. feed - desorbent to circulation loop only 1. feed - desorbent to circulation loop and purge 1. feed - desorbent to circulation loop and repressurization 2. blowdown/pull vacuum 3. purge with desorbent 4. repressurize with desorbent

time (s) at RR ) 0.80 12 4800 7 12 4800 7

energy transfer equations, and equilibrium isotherm in conjunction with their respective boundary and initial conditions. In addition to the assumptions listed previously, ADSIM assumes that the adsorption beds are adiabatic, that the momentum balance follows the Karman-Kozeny equation, and that the gas behavior is ideal. ADSIM solves the governing differential and algebraic equations simultaneously with numerical integration via the method of lines. Each adsorption bed was simulated with 40 grid points in the axial direction; a first-order upwind differencing-type discretization scheme was used and time integration used the implicit Euler technique and a variable step size of 0.01-5 s. Variables expected to change with axial distance, such as concentration, velocity, and density, were discretized over the length of each bed in order to capture the dynamics. Changes in density and velocity occurring as a result of adsorption were accounted for. The pressure drop in the adsorption beds, specified as 0.1 bar, was related to gas velocity by the KarmanKozeny equation. Blowers were added to the flow sheet on the circulation loop to overcome this pressure drop and prevent backflow. All operating conditions are summarized in Tables 1 and 2. The processes were simulated dynamically until cyclic steady state was reached and the material and energy balances converged. Simulations often took approximately 100 cycles to reach cyclic steady state, although more cycles were necessary at high recycle rates. The feed temperature was set at 25 °C; this corresponds to a selectivity (R) of ∼1.5. Although relatively low, this selectivity is reasonable for enantiomeric separations. The high pressure (PH) was set to 4 bar and the fresh feed rate was scaled to maintain the same productivity as the GC-SMB process studied in the literature.22 Specification of the low pressure (PL) and recycle ratio (RR) was more complicated since these quantities are interrelated when D/F ) 0.0. A large value of RR is desirable in this process since it results in better clean-out of any raffinate heel left in zone II during the feed step. However, as RR is increased (with the flow rate of internally generated desorbent constant), proportionately less desorbent is available for purge; also, if RR is increased too much, the desorbent will become contaminated with light product. The circulation vs clean purge tradeoff is an important factor to be considered when operating the SMB/PSA without adding any external desorbent. Results and Discussion For the separation of enflurane enantiomers in a nitrogen carrier gas, the isobaric open-loop (2,2,2,2) GC-SMB process produced products with an average purity of ∼98% at 25 °C and PH of 4 bar.22 This required a D/F value of ∼54.0 after adjusting for recycling that occurs in a closed-loop process. This D/F value is high, probably because of the relatively low adsorbent selectivity and high circulation rate relative to feed rate. The large D/F value can be decreased by using a pressure swing. Two-Zone SMB/PSA with Tank: (1,1)-Tank Configuration. The two-zone SMB/PSA configured as a single train


Figure 7. Average purity vs BPR at various pressure ratios with the two-zone SMB/PSA process in the (1,1)-tank configuration at RR ) 0.8 and D/F ) 0.0.

Figure 8. Average purity vs BPR for various values of RR with the two-zone SMB/PSA process in the (1,1)-tank configuration.

with a tank ((1,1)-tank SMB/PSA) was used during preliminary simulations to find the base cycle timing provided in Table 2. It was also used to ensure that the pressure ratio of PH/PL ) 4/0.2 ) 20 bar was not above the pressure plateau that exists for linear isotherms.3 Figure 7 shows average purity as a function of BPR at a RR of 0.8 and D/F ) 0.0 for two pressure ratios. At a pressure ratio of 20, the (1,1)-tank SMB/PSA process produces products with a maximum average purity of ∼83%. Under the same conditions but with a pressure ratio of 10, the products have a maximum average purity of ∼69%. From these results, it does not appear that a pressure ratio of 20 is above the pressure plateau for the linear isotherm used for this system. Because a pressure ratio of 20 provides more desorbent expansion during purge, better clean-out occurs and higher product purities can be achieved. Feasibility simulations were completed for the (1,1)-tank SMB/PSA process to determine if the SMB/PSA process performs favorably. The average purity

vs BPR results are shown in Figure 8 and Table 3. The maximum average purity attainable at D/F ) 0.0 is ∼83%, which occurs at RR ) 0.8. Clearly, a decent amount of separation is possible without adding any desorbent to the process. Due to its discontinuous nature and the effects of tank mixing, the (1,1)-tank SMB/PSA configuration was not pursued in detail. Two-Zone SMB/PSA with Two Parallel Trains: (1,1)(1,1) Configuration. Extensive simulations for the two-zone SMB/PSA process were done using a configuration with two interacting trains ((1,1)-(1,1) SMB/PSA). First, the effects of BPR and RR were studied for cases when no desorbent was externally added to the process (D/F ) 0.0). A pressure ratio of 20 was used. Results are shown in Table 4 and Figures 9 and 10. When RR was changed, the feed time was scaled linearly to account for changes in velocity, as shown in Table 4; the fresh feed flow rate and the internally generated desorbent flow rate remained constant.

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Table 3. Operating Conditions and Results for Simulations of the Two-Zone SMB/PSA Process in the (1,1)-Tank Configuration with Constant Productivity of 1.90 × 10-11 kg of Enflurane/cm3‚sa RR

BPR

VI,feed (cm/s)

tfeed (s)

raffinate purity (%)

extract purity (%)

average purity (%)

raffinate recovery (%)

extract recovery (%)

average recovery (%)

0.80 0.80 0.80 0.80 0.80 0.80 0.80 0.80 0.80

0.10 0.20 0.30 0.35 0.40 0.45 0.50 0.70 0.90

0.38 0.38 0.38 0.38 0.38 0.38 0.38 0.38 0.38

4800 4800 4800 4800 4800 4800 4800 4800 4800

93.3 90.1 86.3 83.4 81.5 76.3 72.6 61.1 54.8

57.8 66.8 78.0 82.9 85.7 90.2 92.6 96.7 96.9

75.6 78.4 82.1 83.2 83.6 83.2 82.6 78.9 75.8

26.5 53.1 75.3 81.4 85.8 91.2 94.9 98.5 99.4

98.1 93.8 88.1 82.6 79.5 70.1 64.6 39.2 18.0

62.3 73.5 81.7 82.0 82.7 80.7 79.8 68.9 58.7

a

Data correspond to PR ) 20 curve in Figure 7 and RR ) 0.8 curve in Figure 8.

Table 4. Operating Conditions and Results for Simulations of the Two-Zone SMB/PSA Process in the (1,1)-(1,1) Configuration with Constant Productivity of 1.90 × 10-11 kg Enflurane/cm3‚sa RR

BPR

VI,feed (cm/s)

tfeed (s)


extract purity (%)

average purity (%)




0.75 0.75 0.75 0.75 0.75 0.75 0.75 0.80 0.80 0.80 0.80 0.80 0.80 0.80 0.85 0.85 0.85 0.85 0.85 0.85 0.85 0.85 0.85 0.85

0.10 0.20 0.25 0.30 0.35 0.40 0.50 0.10 0.30 0.35 0.40 0.45 0.50 0.70 0.20 0.30 0.40 0.50 0.55 0.60 0.65 0.70 0.75 0.80

0.31 0.31 0.31 0.31 0.31 0.31 0.31 0.38 0.38 0.38 0.38 0.38 0.38 0.38 0.51 0.51 0.51 0.51 0.51 0.51 0.51 0.51 0.51 0.51

6000 6000 6000 6000 6000 6000 6000 4800 4800 4800 4800 4800 4800 4800 3600 3600 3600 3600 3600 3600 3600 3600 3600 3600

92.5 88.2 84.9 80.8 76.6 72.5 65.7 96.1 88.9 85.9 82.7 78.6 74.8 63.0 95.2 92.6 89.1 84.7 81.8 78.8 75.7 72.7 69.9 67.3

62.4 75.3 82.0 87.5 91.8 94.9 98.5 59.0 80.3 85.4 89.6 92.9 95.4 99.6 62.2 68.3 77.2 85.5 88.8 91.6 93.8 95.6 96.9 98.0

77.5 81.8 83.5 84.2 84.2 83.7 82.1 77.6 84.6 85.7 86.1 85.8 85.1 81.3 78.7 80.5 83.2 85.1 85.3 85.2 84.8 84.2 83.4 82.7

41.9 70.3 81.2 88.7 93.7 96.6 99.2 31.5 77.9 85.3 90.6 94.3 96.7 99.8 40.5 55.6 73.0 85.6 89.9 93.1 95.4 97.1 98.1 99.0

96.6 90.6 85.6 78.9 71.4 63.5 48.3 98.7 90.2 86.0 81.1 74.4 67.4 41.2 98.0 95.5 91.1 84.5 80.0 74.9 69.4 63.5 57.7 51.9

69.3 80.5 83.4 83.8 82.6 80.1 73.8 65.1 84.1 85.7 85.9 84.4 82.1 70.5 69.3 75.6 82.1 85.1 85.0 84.0 82.4 80.3 77.9 75.5

a

Data correspond to Figures 9 and 10.

Figure 9 shows that increasing RR yields purer raffinate but less pure extract. Because the flow rates in zones I and II increase with increasing RR, more circulation occurs: this increases the raffinate purity. However, the increased flow rates in zones I and II result in longer penetration distances and the desorbent generated from the feed becomes contaminated with light product. Use of this contaminated desorbent for purge decreases the extract purity. Figure 9 also shows that decreasing RR yields purer extract but less pure raffinate. As RR is decreased, the flow rates in zones I and II decrease; these decreased flow rates result in shorter penetration distances. The desorbent is purer, which increases the extract purity. However, with less circulation the feed velocity becomes relatively larger compared to the recirculating velocity. Because the jump in solute velocities at the feed location is larger, it becomes more difficult to satisfy the fundamental separating mechanism in an SMBsto simultaneously have the velocity of the light component greater than the port velocity in both zones and to have the velocity of the heavy component less than the port velocity in both zones. Thus, the raffinate purity decreases. Figure 9 also shows the effect of BPR, which can be used to fine-tune raffinate and extract purity at a specified RR. Setting BPR too high overpurges zone I by pushing the extract mass transfer zone too far toward the end of the column, which causes

the raffinate product to become contaminated with extract product (the S-enantiomer). On the other hand, underpurging zone I (low BPR) is also troublesome since too much Renantiomer heel remains in the zone and extract purity drops. Zone II cannot be overpurged since it contains only extract product; however, it would be inefficient to use an excessive amount of desorbent in this zone. The results in Figure 9 show that either raffinate or extract (but not both) can be produced with purity >95% with no desorbent (D/F ) 0.0) by adjusting RR and BPR. From the figure, ∼95% pure raffinate can be produced at RR ) 0.85 and BPR of 0.2 while ∼95% pure extract can be produced at RR ) 0.75 and BPR of 0.7. Figure 9 indicates that at D/F ) 0 it is somewhat more difficult to produce pure raffinate than pure extract. An adsorbent with higher selectivity will give higher purities. Figure 10 considers simultaneous production of raffinate and extract by plotting the same data as average purities. The maximum average purity attainable at D/F ) 0.0 is ∼86% at RR ) 0.8 and BPR ) 0.4 with a corresponding maximum average recovery of ∼86%. Although not an extremely high purity, this result shows that the SMB/PSA process performs reasonably well even with no desorbent added; however, the coupling of circulation and purge prevents attainment of high purities at D/F ) 0.0.


Figure 9. Raffinate and extract purity vs BPR for various values of RR with the two-zone SMB/PSA process in the (1,1)-(1,1) configuration at D/F ) 0.0.

Figure 10. Average purity vs BPR for various values of RR with the two-zone SMB/PSA process in the (1,1)-(1,1) configuration.

Desorbent (pure nitrogen) was added to the (1,1)-(1,1) SMB/ PSA process to effectively decouple circulation and purge. This allowed RR to be increased to 0.97 (the feed time was adjusted) while adding desorbent to independently and proportionately supply pure purge gas. In essence, this allowed pure purge gas to be added until the desired average purity was reached. Figure 11 shows the maximum average purity results for the (1,1)(1,1) SMB/PSA process at RR ) 0.97 and various values of BPR. Table 5 contains data corresponding to Figure 11. All data points on the figure correspond to the cases where maximum average purity was attained. At D/F ∼ 5.50, the maximum average purity increased to ∼95%. Further improvements were realized up to ∼99% maximum average purity at D/F ∼ 15.0. Independent specification of the amounts of circulation and purge is necessary to achieve high purities. The desorbent savings are obvious when these results are compared

to the isobaric (2,2,2,2) GC-SMB process in the literature, which achieved ∼98% average purity at D/F ∼ 54.0.22 One may recall that the desorbent savings realized by the SMB/PSA process come at the cost of purging at a vacuum pressure of 0.2 bar. This vacuum pressure is not unrealistic on the commercial scale, and the costs associated with it are likely more than recovered in the desorbent savings. Since the feed is available at 4 bar, no compression costs are incurred since only a small blower is required to overcome the pressure drop in the circulation loop. Further efforts to achieve high purities with low D/F ratios were undertaken. First, the (1,2)-(1,2), (2,1)-(2,1), and (2,2)(2,2) configurations were studied. The length of each bed in the process was fixed and the fresh feed rate was adjusted accordingly to maintain the same productivity. The results (Table 5 and Figure 11) show some improvement: a maximum average purity of ∼99% is achievable at D/F ∼ 12.0 with the (2,2)-

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Figure 11. Maximum average purity vs. D/F with various values of RR and γ for the two-zone SMB/PSA process with different bed configurations. Table 5. Operating Conditions and Results for Simulations of the Two-Zone SMB/PSA Process with Constant Productivity of 1.90 × 10-11 kg of Enflurane/cm3‚sa RR

BPR

D/F

VI,feed (cm/s)

tfeed (s)

bed configb


extract purity (%)

average purity (%)




0.80 0.80 0.80 0.80 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97 0.97

0.40 0.40 0.40 0.40 0.50 0.50 0.50 0.50 0.35 0.35 0.35 0.35 0.25 0.25 0.25 0.25 0.15 0.15 0.15 0.15 0.05 0.05 0.05 0.05

0.00 0.00 0.00 0.00 2.75 2.75 2.75 2.75 5.50 5.50 5.50 5.50 8.75 8.75 8.75 8.75 12.0 12.0 12.0 12.0 15.3 15.3 15.3 15.3

0.38 0.38 0.38 0.38 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52 2.52

4800 4800 4800 4800 730 730 730 730 730 730 730 730 730 730 730 730 730 730 730 730 730 730 730 730

(1,1) (1,2) (2,1) (2,2) (1,1) (1,2) (2,1) (2,2) (1,1) (1,2) (2,1) (2,2) (1,1) (1,2) (2,1) (2,2) (1,1) (1,2) (2,1) (2,2) (1,1) (1,2) (2,1) (2,2)

83.9 82.6 89.8 90.7 88.8 90.0 90.9 94.6 93.9 93.7 96.1 97.5 96.6 96.6 98.1 98.2 98.1 97.8 99.1 99.4 98.9 98.7 99.7 99.8

88.3 90.9 85.2 86.9 92.7 92.7 92.7 90.3 94.6 95.5 93.9 93.1 97.1 97.5 96.9 97.9 98.6 99.3 98.3 99.2 99.1 99.8 99.1 99.3

86.1 86.8 87.5 88.8 90.8 91.3 91.8 92.4 94.2 94.6 95.0 95.3 96.8 97.1 97.5 98.0 98.3 98.5 98.7 99.3 99.0 99.3 99.4 99.7

89.0 91.9 84.3 86.3 93.1 93.0 92.9 89.9 94.7 95.6 93.8 92.8 97.1 97.5 96.8 97.8 98.6 99.3 98.3 99.1 99.1 99.8 99.1 99.3

82.9 80.6 90.4 91.1 88.2 89.6 90.7 94.8 93.8 93.7 96.2 97.6 96.6 96.6 98.1 98.2 98.1 97.8 99.1 99.4 98.9 98.7 99.7 99.8

86.0 86.3 87.4 88.7 90.7 91.3 91.8 92.4 94.3 94.7 95.0 95.2 96.9 97.1 97.5 98.0 98.4 98.6 98.7 99.3 99.0 99.3 99.4 99.6

a

Data correspond to Figure 11. b First and second values correspond to number of beds in raffinate- and extract-producing zones, respectively.

(2,2) SMB/PSA configuration. As expected, the addition of extra beds per zone helps to achieve better separation. The (2,1)(2,1) and (1,2)-(1,2) configurations achieved ∼99% maximum average purities at D/F ∼ 13.0 and D/F ∼ 14.0, respectively. These results imply that adding a bed to the raffinate-producing zone is more favorable than adding it to the extract-producing zone. This is similar to the results obtained for the two-zone SMB for liquid separations25 and is in agreement with the results in Figure 9, which showed that producing pure raffinate is more difficult than producing pure extract. Overall, bed addition decreases the desorbent requirements of the SMB/PSA process to a small, but noticeable, degree. Returning to the (1,1)-(1,1) SMB/PSA configuration, a second option for decreasing desorbent requirements was investigated: bed coupling during regeneration (Figure 12). Essentially, all purge gas is shifted to zone II during purge.

Then, a fraction of the effluent from zone II is taken off as product while the remaining fraction (BPR) is used as purge in zone I. As shown in Table 6, bed coupling is only marginally favorable. At D/F ) 0.0 the maximum average purity was ∼87%, only 1% higher than the analogous uncoupled process. At D/F ∼ 15.0, a maximum average purity of ∼99% was attained; this is no change from the uncoupled process. In general, bed coupling had little effect on decreasing the desorbent requirements of SMB/PSA for enflurane separation, likely due to its low selectivity. The difficulty of the enflurane separation has been evidenced in the performance of the SMB/PSA process developed in this work; it was also discussed in the literature on the GC-SMB process.22 To extract the effects of the adsorbent selectivity (R), it was arbitrarily doubled to a hypothetical value of 3.0. The isotherm parameters for the strongly adsorbed species (S-


Figure 12. Schematic of the two-zone SMB/PSA process with regeneration coupling in a (1,1)-(1,1) configuration. Table 6. Results for Simulations of the Two-Zone SMB/PSA Process with Regeneration Coupling in the (1,1)-(1,1) Configuration with Constant Productivity of 1.90 × 10-11 kg of Enflurane/cm3‚sa

a

D/F (mol/mol)

effective BPR

average purity (%)

0.0 5.5 15.3

0.40 0.35 0.05

87 94 99

Refer to Table 5 for operating conditions.

Table 7. Results for Simulations of the Two-Zone SMB/PSA Process with r ) 3.0 in the (1,1)-(1,1) Configuration with Constant Productivity of 1.90 × 10-11 kg of Enflurane/cm3‚sa

a

D/F (mol/mol)

average purity (%)

0.0 3.5 7.0

91 96 99

Refer to Table 5 for operating conditions.

enflurane) were unchanged, while the parameters of the weakly adsorbed species (R-enflurane) were reduced. With this new selectivity, the feed time was adjusted and the (1,1)-(1,1) SMB/ PSA was simulated at RR ) 0.97 and varying γ and D/F. Table 7 shows that a selectivity of R ) 3.0 reduces desorbent requirements by a significant amount: a maximum average purity of ∼99% can now be attained at D/F ∼ 7.0, less than half the value at R ∼ 1.5. Overall, these results show that adjusting adsorbent properties to increase selectivity is potentially more effective at decreasing desorbent requirements than adding more beds per zone or using bed coupling during regeneration. The two-zone SMB/PSA processes have been able to lessen desorbent requirements and still produce desirable products. Because of their configuration, the two-zone processes were well-suited for SMB/PSA: they allowed complete pressure swings of both the raffinate- and extract-producing zones during the purge step. Thus, the two-zone SMB/PSA process was the focus of this work. Four- and Five-Zone SMB/PSA Processes: (1,1,1,1) and (1,1,1,1,1) Configurations. One may wish to apply the SMB/ PSA process to a four- or five-zone SMB as shown in Figures

Table 8. Results for Simulations of the Four- and Five-Zone SMB/PSA Processes in the (1,1,1,1) and (1,1,1,1,1) Configurations, Respectively, with Constant Productivity of 1.90 × 10-11 kg of Enflurane/cm3‚sa five-zone SMB/PSA D1/F D2/F total D/F raffinate purity (%) extract purity (%) average purity (%) a

0 0 0 69 72 71

0 3.30 3.30 85 89 87

0 8.23 8.23 94 89 92

19.0 3.30 22.3 95 97 96

four-zone SMB/PSA 0 0 0 65 70 68

0 8.23 8.23 88 85 87

19.0 3.30 22.3 92 95 93

Refer to Table 5 and text for operating conditions.

2 and 3. These configurations were also simulated, and results are shown in Table 8. In this case, D1 and D2 refer to highpressure and low-pressure desorbent addition, respectively, and DTotal ) D1 + D2 (Figures 2 and 3 and Table 8). When D1 ) D2 ) 0, Table 8 shows that the five-zone SMB/ PSA process results in a maximum average purity of 71% while the four-zone SMB/PSA process achieves a maximum average purity of 68%. When only low-pressure desorbent is added (D1 ) 0), the four- and five-zone SMB/PSA processes achieve maximum average purities of 87% and 92%, respectively, at D2/F ) D/F ) ∼8.2. Compared to the (1,1)-(1,1) SMB/PSA process at the same D/F, these processes produce products of lower average purities. Further improvement is achieved when high-pressure desorbent is added to the active loops of the fourand five-zone processes. At D1/F ∼ 19.0 and D2/F ∼ 3.0 (DTotal ∼ 22.0), the four- and five-zone SMB/PSA processes achieved maximum average purities of 93% and 96%, respectively. This is a high molar D/F ratio since the desorbent (D2) is being added at high pressure. The five-zone SMB/PSA process achieves higher average purities than its four-zone counterpart at the same value of D/F, confirming the earlier finding that an extra bed gives better separation. Overall, the four- and five-zone SMB/ PSA processes do not perform as well as the two-zone (1,1)(1,1) SMB/PSA process at the same total D/F, even with highpressure desorbent addition. These results imply that SMB/PSA is better-suited with a two-zone configuration, with its complete

3148


pressure swing, than with a four- or five-zone configuration, with their partial pressure swings. Summary and Conclusions SMB/PSA, a new hybrid simulated-moving bed (SMB)/ pressure-swing adsorption (PSA) process, has been developed for separating dilute binary gas species in an inert carrier gas. In comparison with an isobaric gas-phase SMB, SMB/PSA has two key features: regeneration/purge at low pressure and internal generation of desorbent (carrier gas). The SMB/PSA process uses gas expansion at low pressure to fully take advantage of the desorbent present in the feed to decrease molar desorbent requirements. Among the several SMB/PSA configurations developed, the two-zone implementation with one bed per zone and two interacting trains ((1,1)-(1,1) SMB/PSA) appeared to be the best suited for taking advantage of a full pressure swing. As a model system, the relatively difficult enantiomeric separation of the inhalation anesthetic enflurane in a nitrogen carrier gas was chosen. Without adding any external desorbent (D/F ) 0), the (1,1)-(1,1) SMB/PSA configuration achieved a maximum average purity of ∼86% with a concomitant recovery of ∼86%. In addition, by manipulating the recycle ratio (RR) and the bed purge ratio (BPR) at D/F ) 0, either the raffinate or extract product could be produced with purity greater than or equal to 95%, but with lower recoveries. Achieving simultaneous production of raffinate and extract at higher purities required addition of nitrogen carrier gas (desorbent). At a D/F ∼ 15.0, the (1,1)-(1,1) SMB/PSA achieved a maximum average purity of ∼99% and maximum average recovery of ∼99%. For the same separation at the same productivity, an isobaric gas-phase SMB achieved ∼98% average purity at D/F ∼ 54.22 The use of a pressure swing and the internal generation of desorbent enable the SMB/PSA process to have considerably lower molar desorbent requirements while still producing products of desired purity and recovery. All of the newly developed SMB/PSA processes reduce the amount of desorbent required for the relatively difficult separation of R- and S-enflurane enantiomers in a nitrogen carrier gas compared to the isobaric GC-SMB.20-22 These desorbent savings result from taking advantage of the expansion of carrier gas at low pressure for the purge step. Of the four configurations examined (two-zone with tank, twozone with two interacting trains, four-zone, and five-zone), the two-zone configuration with two interacting trains gave the highest purities. The hybrid SMB/PSA process may be an alternative to existing SMB and PSA processes for gas separations. Acknowledgment The authors gratefully acknowledge the technical support staff at Aspen Technology, Inc., for their assistance with ADSIM and the National Science Foundation Grant CTS-0327089 for providing part of the funding for this research. Nomenclature ap ) external surface area per volume, m2/m3 BPR ) bed purge ratio as defined in eq 6, dimensionless ci ) solute i concentration of fluid, kmol/m3 ci,pore ) average solute concentration i in pore cf ) concentration (or molar density) of feed, mol/m3 Cp,f ) fluid-phase heat capacity, J/kg/K Cp,s ) solid-phase heat capacity, J/kg/K dp ) particle diameter, cm

D ) molar amount of desorbent, mol D1 ) high-pressure desorbent, mol D2 ) low-pressure desorbent, mol DTotal ) total amount of high- and low-pressure desorbents, mol D/F ) molar desorbent-to-feed ratio, dimensionless ∆Hads ) heat of adsorption, kJ/kmol E ) activation energy, kJ/mol Ez ) axial dispersion coefficient, cm2/s Ez,T ) thermal axial dispersion coefficient, cm2/s F ) molar amount of fresh feed, mol Ffeed ) molar flow rate of fresh feed, mol/s hHTC ) heat transfer coefficient, W/m2/K K ) linear isotherm parameter kez ) thermal conductivity of gas phase, W/m‚K Ki ) equilibrium constant for species i Kdi ) fraction of interparticle volume species i can penetrate kMTC,fluid ) linear lumped parameter mass transfer coefficient, 1/s ks ) thermal conductivity of solid phase, W/m‚K L ) bed length, m MWenf ) molecular weight of enflurane, g/mol ncol ) number of columns, dimensionless PH ) high (feed) pressure, bar PL ) low (purge) pressure, bar RR ) recycle ratio, dimensionless qi ) amount of solute i adsorbed, kmol/kg adsorbent qj ) average amount of solute adsorbed qi* ) equilibrium amount adsorbed of species i qi ) average amount of solute i adsorbed R ) universal gas constant RR ) recycle ratio as defined in eq 5, dimensionless tbd ) time of blowdown step, s tcycle ) total cycle time, s tfeed ) time of feed step, s tpr ) time of pressurization step, s tpurge ) time of purge step, s T h * ) average equilibrium temperature, K Ts ) solid-phase temperature, K Ts ) average solid-phase temperature, K Vcol ) column volume, cm3 VI,feed ) velocity of stream in zone I during feed step, cm/s VII,feed ) velocity of stream in zone II during feed step, cm/s VI,purge ) velocity of stream in zone I during purge step, cm/s VT,purge ) total velocity of desorbent gas during purge (VT,purge ) VI,purge + VII,purge), cm/s yi ) mole fraction of solute i yenf feed ) mole fraction of enflurane in the fresh feed stream, dimensionless Greek Symbols R ) selectivity ) KS/KR, dimensionless e ) external porosity, m3 of void/m3 of bed p ) internal porosity, m3 of pore/m3 of particle Fs ) solid-phase density, kg/m3 Ff ) fluid-phase density, kg/m3 Literature Cited (1) Humphrey, J. L.; Keller, G. E., II. Separation Process Technology; McGraw-Hill: New York, 1997. (2) Noble, R. D.; Agrawal, R. Separations Research Needs for the 21st Century. Ind. Eng. Chem. Res. 2005, 44, 2887. (3) Ruthven, D. M.; Farooq, S.; Knaebel, K. Pressure-Swing Adsorption; Wiley: New York, 1994.

Ind. Eng. Chem. Res., Vol. 47, No. 9, 2008 3149 (4) Wankat, P. C. Large-Scale Adsorption and Chromatography; CRC Press: Boca Raton, FL, 1986. (5) Skarstrom, C. W. Use of Adsorption Phenomena in Automatic PlantType Gas Analyzers. Ann. N. Y. Acad. Sci. 1959, 72, 751. (6) Skarstrom, C. W. Oxygen Concentration Process. U.S. Patent 3,237,377, 1966. (7) Tondeur, D.; Wankat, P. C. Gas Purification by Pressure Swing Adsorption. Separ. Purif. Methods 1985, 14, 157. (8) Kostroski, K. P.; Wankat, P. C. High Recovery Cycles for Gas Separations by Pressure-Swing Adsorption. Ind. Eng. Chem. Res. 2006, 45, 8117. (9) Diagne, D.; Goto, M.; Hirosi, T. New PSA Process with Intermediate Feed Inlet Position and Operated with Dual Refluxes: Application to Carbon Dioxide Removal and Enrichment. J. Chem. Eng. Jpn. 1994, 27, 85. (10) Ebner, A. D.; Ritter, J. A. Equilibrium Theory Analysis of Rectifying PSA for Heavy Component Production. AIChE J. 2002, 48, 1679. (11) Diagne, D.; Goto, M.; Hirosi, T. Parametric Studies on CO2 Separation and Recovery by a Dual Reflux PSA Process Consisting of Both Rectifying and Stripping Sections. Ind. Eng. Chem. Res. 1995, 34, 3083. (12) McIntyre, J. A.; Holland, C. E.; Ritter, J. A. High Enrichment and Recovery of Dilute Hydrocarbons by Dual-Reflux Pressure-Swing Adsorption. Ind. Eng. Chem. Res. 2002, 41, 3499. (13) Kearns, D. T.; Webley, P. A. Modelling and Evaluation of DualReflux Pressure Swing Adsorption Cycles: Part I. Mathematical Models. Chem. Eng. Sci. 2006, 61, 7223. (14) Kearns, D. T.; Webley, P. A. Modelling and Evaluation of Dual Reflux Pressure Swing Adsorption Cycles: Part II. Productivity and Energy Consumption. Chem. Eng. Sci. 2006, 61, 7239. (15) Broughton, D. B.; Gerhold, G. G. Continuous Sorption Process Employing Fixed Bed of Sorbent and Moving Inlets and Outlets. U.S. Patent 2,985,589, 1961. (16) Broughton, D. B. Continuous Simulated Countercurrent Sorption Process Employing Desorbent Made in Said Process. U.S. Patent 3,291,726, 1966. (17) Depta, A.; Giese, T.; Johannsen, M.; Brunner, G. Separation of Stereoisomers in a Simulated Moving Bed-Supercritical Fluid Chromatography Plant. J. Chromatogr., A 1999, 865, 175. (18) Jin, W.; Wankat, P. C. Two-Zone SMB Process for Binary Separation. Ind. Eng. Chem. Res. 2005, 44, 1565.

(19) Xie, Y.; Mun, S.; Kim, J. H.; Wang, N.-H. L. Standing Wave Design and Experimental Validation of a Tandem Simulated Moving Bed Process for Insulin Purification. Biotechnol. Prog. 2002, 18, 1332. (20) Juza, M.; Di Giovanni, O.; Biressi, G.; Schurig, V.; Mazzotti, M.; Morbidelli, M. Continuous Enantiomer Separation of the Volatile Inhalation Anesthetic Enflurane with a Gas Chromatographic Simulated Moving Bed Unit. J. Chromatogr., A 1998, 813, 333. (21) Biressi, G.; Quattrini, F.; Juza, M.; Mazzotti, M.; Schurig, V.; Morbidelli, M. Gas Chromatographic Simulated Moving Bed Separation of the Enantiomers of the Inhalation Anesthetic Enflurane. Chem. Eng. Sci. 2000, 55, 4537. (22) Biressi, G.; Mazzotti, M.; Morbidelli, M. Experimental Investigation of the Behavior of Gas Phase Simulated Moving Beds. J. Chromatogr., A 2002, 957, 211. (23) Abel, S.; Babler, M. U.; Arpagaus, C.; Mazzotti, M.; Stadler, J. Two-Fraction and Three-Fraction Continuous Simulated Moving Bed Separation of Nucleosides. J. Chromatogr., A 2004, 1043, 201. (24) Paredes, G.; Abel, S.; Mazzotti, M.; Morbidelli, M.; Stadler, J. Analysis of a Simulated Moving Bed Operation for Three-Fraction Separations (3F-SMB). Ind. Eng. Chem. Res. 2004, 43, 6157. (25) Xie, Y.; Chin, C. Y.; Phelps, D. S. C.; Lee, C.-H.; Lee, K. B.; Mun, S.; Wang, N.-H. L. A Five-Zone Simulated Moving Bed for the Isolation of Six Sugars from Biomass Hydrolyzate. Ind. Eng. Chem. Res. 2005, 44, 9904. (26) Jin, W.; Wankat, P. C. Two-Zone SMB Process for Binary Separation. Ind. Eng. Chem. Res. 2005, 44, 1565. (27) Kumar, R. Vacuum Swing Adsorption Process for Oxygen ProductionsA Historical Perspective. Sep. Sci. Technol. 1996, 31, 877. (28) U.S. Occupational Safety & Health Administration Web Site. http://www.osha.gov (accessed December 2007). (29) Wankat, P. C. Rate-Controlled Separations; Kluwer: Amsterdam, 1990. (30) Bird, R. B.; Stewart, W. E.; Lightfoot, E. N. Transport Phenomena; John Wiley & Sons, Inc.: New York, 1960; Chapter 21.

ReceiVed for reView July 23, 2007 ReVised manuscript receiVed January 31, 2008 Accepted January 31, 2008 IE071000B

Separation of Dilute Binary Gases by Simulated-Moving Bed with

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