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The separation is carried out in a six-port vapor phase SMB pilot plant operated at 448 K and 350 kPa, using n-heptane as desorbent and 5A-zeolite as ...
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Ind. Eng. Chem. Res. 1996, 35, 2313-2321

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Vapor-Phase SMB Adsorptive Separation of Linear/Nonlinear Paraffins Marco Mazzotti,† Renato Baciocchi, Giuseppe Storti,‡ and Massimo Morbidelli* Laboratorium fu¨ r Technische Chemie LTC, ETH Zentrum, CAB, Universita¨ tstrasse 6, CH-8092 Zu¨ rich, Switzerland

The separation of various mixtures of the C5-C6 paraffin fraction is considered, being an important stage of the process aimed to gasoline octane enhancement. The separation is carried out in a six-port vapor phase SMB pilot plant operated at 448 K and 350 kPa, using n-heptane as desorbent and 5A-zeolite as adsorbent. The experimental runs deal with the separation of both a binary mixture (70% isopentane and 30% n-pentane) and a multicomponent C5-C6 stream. The results of these experiments illustrate the role and the effect of the different operating parameters on the process performance and demonstrate the possibility of achieving high performance separation. Moreover, they nicely assess the reliability of the known criteria for selecting robust and optimal operating conditions for SMB units. Introduction Large scale adsorptive separations based on the simulated moving bed (SMB) technology (cf. Ruthven and Ching, 1989) have been widely applied to mixtures of hydrocarbons since the development of the Sorbex process by UOP (Broughton and Gerhold, 1961; Johnson and Kabza, 1993). This technique exploits the principle of displacement chromatography; hence, in addition to a suitable adsorbent (typically some kind of synthetic zeolite) a proper desorbent is needed. Adsorption-based separations exhibit two advantages with respect to distillation. First, isomers with very close boiling points can be separated, such as, for example, in the Parex process where p-xylene is separated from a mixture of the alkylaromatic C8 fraction. Secondly, a feed stream constituted of many components can be fractionated based on molecular properties other than differences in boiling points. Examples are given by the separation of paraffins from olefins in the Olex process, based on the presence of the unsaturated double bond, and the separation of linear paraffins from branched and cyclic ones in the Molex process (Molex, Olex, and Parex are trademarks of UOP), based on steric factors. In general, industrial SMB separation units are operated in the liquid phase, with 12 or 24 ports, i.e., adsorption columns, properly arranged in one or two large cilindrical vessels. However, in the last few years the feasibility of vapor phase SMB separations has been demonstrated at the laboratory scale. In particular, vapor phase SMB units with a small number of ports (six in the reported cases) achieve high purity performances in the separation of xylene isomers (either por m-xylene), provided that the column size and the operating conditions are properly designed (Storti et al., 1992, 1993b). A productivity of the order of 0.15 g per gram of zeolite per hour with high purity of the outlet * To whom correspondence should be addressed. Tel.: +411-6323034. Fax: +41-1-6321082. E-mail: morbidelli@ tech.chem.ethz.ch. † Present address: Dipartimento di Chimica, Politecnico di Milano, Via Mancinelli, 7, 20131 Milano, Italy. E-mail: [email protected]. ‡ Present address: Dipartimento di Ingegneria Chimica e Materiali, Universita` degli Studi di Cagliari, Piazza d’Armi, 09123 Cagliari, Italy. E-mail: [email protected].

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streams and high recovery of the feed components has been achieved. It is evident that the practical and economical applicability of this mode of operation must be evaluated case by case, considering, for example, the possibility of integrating the energy utilization of the vapor phase SMB unit within the whole industrial plant. The vapor phase operation appears particularly promising from the applicative point of view when considering the separation of components more volatile than xylenes, such as the so-called light naphta fraction C5C6. In this case, a lower operating temperature can be adopted, with important advantages in terms of energy consumption and zeolite deactivation, which may become important at high temperatures. The fractionation of light naphta, which typically consists of a mixture of pentanes and hexanes, (together with a small fraction of butanes, heptanes and aromatics) with a molar ratio between linear and nonlinear paraffins equal to about 1, is one of the steps of a process aimed to gasoline octane enhancement. The feed to the separation unit is the outlet of the catalytic isomerization reactor, where the nonlinear content is raised to about 70%. The objective of the separation is to produce a nonlinear paraffin rich stream, whereas the linear rich fraction is recycled. This can be achieved by adsorptive separation using 5A-zeolite as adsorbent solid, since linear paraffins are adsorbed but neither the branched nor the cyclic ones are (Barrer, 1978). This feature of the adsorbent solid represents a remarkable difference with respect to other separations, such as those involving xylenes, where all components of the feed stream are adsorbed with very similar saturation loading capacities but different selectivities. In this work, the separation of various mixtures of the C5-C6 paraffin fraction is considered, in the context of the light naphta treatment mentioned above. In particular, we study an alternative process which, rather naturally originating from the two commercially available processes licensed by UOP and Union Carbide (cf. Raghuram and Wilcher, 1992), combines the high mass transfer efficiency peculiar of the continuous simulated moving bed separation of the UOP process, i.e., the liquid phase Molex process, with the reduced nonselective hold-up peculiar of the vapor phase cyclic operation of the Union Carbide technology. The separa© 1996 American Chemical Society

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tion is carried out in a six-port vapor phase SMB pilot plant operated at 448 K and 350 kPa, using n-heptane as desorbent and 5A-zeolite as adsorbent. Two cases are considered: the first one deals with the separation of a binary mixture of 70% isopentane and 30% npentane. The second one is concerned with the separation of a multicomponent stream constituted of seven linear and nonlinear C5-C6 paraffin species. Even though the study of binary separations is of key importance from the conceptual point of view, the scale-up of its results to the separation of multicomponent mixtures of industrial interest is not straightforward. Two important issues must be accounted for when considering such multicomponent feed streams: the presence and role of impurities on one side and the change in the separation performance when going from a binary to a multicomponent feed on the other side. The former issue requires the treatment of true industrial feedstocks and long-lasting operation, since impurities may often have a long-term effect. The latter issue can be addressed simply using a synthetic multicomponent mixture with no impurities, since the effect on separation performance and process dynamics of the multicomponent feed with respect to the binary feed can be fully investigated at the laboratory scale. In this work the multicomponent separation mentioned above has been considered with the aim of analyzing this last aspect. In addition to the development of new separations or new modes of operation for known separations, a considerable theoretical effort has been devoted to the definition of criteria for the choice of optimal and robust operating conditions following an approach based on the modeling of adsorptive separation units through equilibrium theory (Storti et al., 1993a, 1995; Mazzotti et al., 1994). The cited results typically refer to systems characterized by either the linear adsorption equilibrium isotherm or the constant selectivity stoichiometric Langmuir isotherm. Neither can be used to describe the adsorption behavior of the systems considered in this work. For these we need to introduce the constant selectivity nonstoichiometric Langmuir isotherm, which allows us to account for widely different saturation loading capacities of the chemical species involved. Criteria for the optimal design of countercurrent separation units involving systems of this kind have been recently developed by Mazzotti et al. (1996). Thus, the experimental investigation developed in this work provides a nice opportunity of validating these theoretical findings. Experimental Setup In this section the adopted unit configuration is first discussed and compared with the classical four-section configuration. Then, the SMB pilot plant is described, analyzing separately the SMB separation section, the feed and withdrawal zones, and the control system of the pilot unit. Finally, the different separations examined in this work, including the adopted desorbent and adsorbent, are illustrated. Unit Configuration. In principle, adsorptive separations are most conveniently operated in a continuous unit where the solid and the fluid phases move countercurrently (Ruthven, 1984). The scheme of a true countercurrent (TCC) unit of this type in the classical closedloop four-section configuration is illustrated in Figure 1. There are two inlet streams, i.e., the mixture to be separated and the desorbent makeup, respectively. The

Figure 1. Scheme of a four-section true countercurrent unit for adsorption separation.

two outlet streams are the extract, where the most adsorbable components are preferentially collected, and the raffinate, which instead contains the least adsorbable species. Each one of the four sections plays a specific role in the separation process. Sections 2 and 3 perform the fractionation, by conveying the strong components downward toward the extract outlet and the weak components upward toward the raffinate outlet. Section 1 regenerates the adsorbent by desorbing the most adsorbable components, whereas section 4 regenerates the desorbent by adsorbing the least adsorbable ones. The presence of these two sections makes the recycle of both the adsorbent and the desorbent possible. The practical problems related to the actual movement of the solid phase can be overcome by simulating this countercurrent contact by moving the fluid phase and the column in a fixed bed unit. The continuous column movement is simulated by a discrete movement obtained by shifting the feed and withdrawal nodes at discrete time along the column axis in the same direction of the fluid flow. In order to closely mimic the countercurrent apparatus the sections of the separation unit are divided in several ports or subsections. Each one of these is constituted of a fixed bed which is fed with the fluid stream leaving the previous subsection after addition or withdrawal of an external stream when required. The two-unit configurations are related by a set of geometric and kinematic equivalence relationships, which state that: (i) the relative velocity of the two phases must be the same whether the adsorbent beds are fixed or not; (ii) the solid velocity in the TCC unit corresponds in the SMB configuration to the average velocity given by the ratio between the port length, L, and the switch time, t* (which is the time period between two successive port switches); (iii) the section length in the TCC unit, Lj, is equal to the port length times the number of ports per section in the SMB configuration, i.e., Lj ) Lnj. This equivalence allows one to use countercurrent models, which are simpler, to simulate the behavior of SMB units (cf. Storti et al., 1993a).

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Figure 2. Scheme of the six-port three-section open-loop simulated moving bed laboratory unit.

The closed-loop four-section configuration described above is adopted in the Sorbex process, and an openloop modification of it has been used for the vapor phase separation of xylenes (Storti et al., 1992, 1993b). However, this would not be the optimal process configuration for the system considered in this work. Here, in fact, the weak components are indeed very weak, with almost negligible adsorptivity, whereas the desorbent is actually stronger than all the components to be separated. Thus, the task to be performed by section 4, i.e., to adsorb all the weak components and regenerate the desorbent to be recycled, becomes rather difficult. In this case it is convenient not to regenerate the desorbent and adopt an open-loop three-section configuration. The equivalent SMB configuration is illustrated in Figure 2, which corresponds to the actual configuration of the pilot plant used in this work. As anticipated, this is constituted of six adsorption columns, which are equally distributed among the three sections, thus yielding a 2-2-2 configuration. The whole outlet of section 3 is collected as the raffinate stream. SMB Pilot Plant. The SMB pilot plant used in this work is schematically illustrated in Figure 3. It consists of the columns where the separation takes place and the tubes and valves used to convey fluid streams and perform port switches. The above-mentioned apparatus is contained inside a thermostatic chamber where temperature values not larger than 573 K can be kept uniform and constant with a maximum error of about 1 K. The chamber is maintained under nitrogen flow to prevent air from entering the chamber and to keep its atmosphere inert. Each adsorption bed is a 1.2 m long stainless steel column (0.015 m internal diameter). The adsorbent solid fills 1 m of each column, the other 0.2 m being stuffed with glass wool so as to properly distribute the inlet flow and retain the solid particles. As an extra protection a filter, with a sintered element of 7 µm size, is located at the outlet of each column. As shown in Figure 3, the port switching, which simulates the fluid-solid countercurrent movement, is performed by three 6 + 1 port six-position valves and one 12 + 1 port six-position valve. Two of the 6 + 1 port valves are used to convey the feed and the desorbent streams to each of the six column inlets, while the other one connects each column outlet to the extract reservoir. On the other hand, the 12 + 1 port valve is used to connect each column outlet to the inlet of the following column, except for the column connected to the raffinate reservoir. The four multiposition valves are chromatographic valves with maximum operating temperature and pressure of 573 K and 1400 kPa, respectively. Each valve is electrically actuated, being connected to the outside actuator by a specially designed

0.45 m long stand-off, passing through a hole in the wall of the thermostatic chamber. Outside the chamber the stand-off is cooled by means of a water coil, so as to keep the actuator below its maximum operating temperature, i.e., about 320 K. All the lines in the hot section of the plant are made of 1/8 in. stainless steel tubes, connected by stainless steel fittings. Liquid reservoirs are used to store the desorbent and the feed mixture; for each feedstock there is one main and one backup tank. Both are kept under constant pressure with nitrogen and equipped with a relief valve. The flow rate of the fresh desorbent is controlled in the liquid state through a mass flow meter and controller. This control system is virtually independent of pressure and temperature, but it needs a minimum pressure drop of about 100 kPa for the control valve to operate properly. The desorbent and the feed stream, which comes directly from the feed tank, are first vaporized and then enter the thermostatic chamber containing the SMB unit where each of them is conveyed to the relevant multiposition valve. Two streams come out from the chamber. These are first condensed in simple tube-tube heat exchangers and then metered and controlled by two liquid flow control systems identical to that indicated above. Since the liquid flow control systems have a pressure drop of about 100 kPa and the collecting tanks are at atmospheric pressure, the pressure in the unit must be at least 200 kPa. Sampling points (not shown in Figure 3 are positioned after the mass flow meters and controllers so as to avoid any perturbation of the flow rates in the separation section during sampling. Particular care is taken in keeping the sampling vials at low temperature with dry ice, so as to avoid partial evaporation of the sample. All lines in the cold section of the pilot plant are made of 1/8 in. copper tubes, connected by brass fittings. Control of the Unit. Let us refer to the open loop configuration shown in Figure 2, where five mass flow rates must be assigned in order to uniquely define the steady state operating conditions of the SMB unit. These are the mass flow rates in sections 1 and 3, which coincide with those of desorbent and raffinate, respectively, the mass flow rate in section 2, and the mass flow rates of extract and feed streams. There are only two relationships among these five flow rates, i.e., the overall material balances at the extract and feed nodes. Hence, three degrees of freedom are left and must be constrained in order to fully determine the system and keep the desired operating conditions. For this only three mass flow rates have been controlled, those of the desorbent, extract, and raffinate streams, while leaving free the feed stream. In addition to flow rates, it is necessary to control the pressure level in the unit, which, as mentioned above, should exceed the minimum value of 200 kPa required by the proper operation of the flow control system. This can be achieved by simply imposing this pressure value to the feed tank, since this is directly connected to the SMB unit, with negligible pressure drop in between. A number of experiments with different process schemes has shown that this strict constraint on the pressure level in the unit produces a very effective stabilization of the operating regime of the process. Experimental Runs. The experiments performed in this work deal with the separation of binary and multicomponent mixtures of pentanes and hexanes. In

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Figure 3. Scheme of the experimental SMB laboratory unit.

all cases the mole percentage of nonlinear components is about 70%, hence similar to that obtained in the industrial processes after the catalytic isomerization stage. Binary experimental runs deal with the separation of a mixture of 70% isopentane and 30% n-pentane. Five multicomponent experiments have also been performed, indicated by the letters T, U, V, X, and Y, with feed streams constituted of seven species. The exact feed composition in runs T, U, and V is (expressed as weight percentage): isopentane 39%, 2-methylpentane 10%, 3-methylpentane 9%, methylcyclopentane 7%, cyclohexane 1%, n-pentane 25%, and n-hexane 9%. In runs X and Y, isopentane, cyclohexane, and n-hexane concentrations have been changed to 32%, 7%, and 10%, respectively. The adsorbent is constituted of 1/16 in. pellets of commercial 5A-zeolite manufactured by Union Carbide. In all the experiments n-heptane has been used as desorbent due to the following arguments: it is a linear hydrocarbon with a high adsorptivity with respect to the chosen adsorbent; it is heavier than the species to be separated, thus easy to remove from the extract and raffinate streams by distillation; it is the lightest species having the two aforementioned characteristics, thus minimizing the increase in operating temperature due to the presence of a heavier component than the ones contained in the feed stream. The experimental samples have been analyzed by means of a Carlo Erba Strumentazione gas chromatograph, equipped with a FID detector. Design of the Operating Conditions of the SMB Unit. The criteria for designing the operating conditions of SMB units have been derived using a simplified model, based on the equilibrium theory (Storti et al., 1993a, 1995; Mazzotti et al., 1994, 1996). In this frame both axial mixing and mass transport resistances are neglected. The process is isothermal, and the fluid flow rate is constant. Therefore, thermodynamic equilibrium conditions are assumed to be established everywhere in the column at any time.

The criteria above are available for at least three adsorption equilibrium models: the linear model, the constant selectivity stoichiometric Langmuir model, and the constant selectivity nonstoichiometric Langmuir model. Only the latter one is able to describe, at least approximately, the system under consideration here. In the case of a mixture constituted of NC adsorbable species this model is given by the following relationship

θi )

(Γ∞i /Γ∞)Kici

Γi Γ

) ∞

1+

NC Klcl ∑l)1

(Γ∞i /Γ∞)FfKiyi ) 1 + Ff

NC K l yl ∑l)1

(1)

where Γi and ci are the adsorbed phase and fluid phase concentration of component i, respectively; θi and yi are the corresponding dimensionless quantities; Γ∞ is the reference adsorbed phase concentration, whereas Γ∞i is the saturation loading capacity of component i, which is different for each component; Ff is the overall fluid phase molar concentration; and Ki is the equilibrium constant of the i-th component. In the frame of equilibrium theory the key operating parameters of countercurrent adsorptive separation units (i.e., TCC and SMB) are the flow rate ratios mj in each section of the unit. In the case of a SMB unit, these are defined as follows

mj )

Gjt*/Mav - V*Ff VFsΓ∞(1 - *)

(j ) 1, ..., NS)

(2)

where, with reference to section j, Gj is the fluid mass flow rate; t* is the switch time; Mav is the average molar mass of the fluid phase; V is the column volume; Ff is the overall fluid phase molar density; Γ∞ is the reference adsorbed phase concentration; * ) ( + (1 - )p), where  and p are the inter- and intraparticle void fractions, respectively; Fs is the adsorbent solid mass density; and NS is the number of sections of the unit, which in the case under examination is 3.

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The criteria for the choice of the operating conditions which lead to the desired separation can be expressed in the form of constraints on the operating parameters mj. These constraints are very simple as far as section 1 is concerned since m1 must be greater than a critical value (Mazzotti et al., 1996). On the other hand, the operating parameters in sections 2 and 3 are tightly coupled. Different regions in the operating parameter plane (m2, m3) can be identified, each one corresponding to operating conditions which in the frame of equilibrium theory achieve a certain separation performance, provided that m1 fulfills the constraint reported above. Each of these regions will be identified in the following for the specific separations investigated experimentally, using the procedure described in detail elsewhere (Mazzotti et al., 1996). Characterization of the Adsorption Equilibria. Before considering the separation process, let us characterize the adsorption equilibria of the systems under examination. In particular, we will evaluate the parameters appearing in eq 1, i.e., the equilibrium constant, Ki, and the saturation loading capacity, Γ∞i , for each component. It is worth noting that the selected model is certainly not the best possible choice, and other more advanced thermodynamic models are available that would indeed provide a better description of the adsorption equilibria of the system under examination. However, for these models no equilibrium theory results for describing fixed bed adsorbers are available, and then no design criteria can be derived. Thus, we select the best equilibrium model which is compatible with a simple description of the separation unit behavior. This approach is even more justified when the multicomponent separation problem involving eight components is considered. Even the use of model (1) is in this case prohibitive since it would involve the estimation of 16 equilibrium parameters. For this we will have to adopt an appropriate shortcut procedure. Binary Mixtures. In order to characterize the system used in the binary separation experiments both single component literature data and two-component experimental results have been used. The latter information have been provided by experiments carried out in the apparatus described by Paludetto et al. (1987). All the reported data have been either measured or estimated at 448 K, i.e., the operating temperature of the SMB experiments. The saturation loading capacity of n-heptane on 5A∞ zeolite pellets that we have measured is Γn-C7 ) 0.7 mol/kg. This value is consistent with those measured by Haber et al. (1973) on pellets (0.8 mol/kg) and by Doetsch et al. (1974) on pure crystals (1.2 mol/kg). The differences between pellets and pure crystal data are due to the presence of about 25% of a clay binder and to the incomplete removal of water. In the latter work it was also proved that a satisfactory interpretation of the equilibrium data could be obtained using a nonideal model isotherm based on either energetic heterogeneity or sorbate-sorbate interaction, but not using a simple Langmuir isotherm. In particular, Figure 1 of the cited paper shows that the equilibrium constant of the single component Langmuir model depends on the coverage fraction, instead of being constant as it is assumed by the Langmuir model. However, from the same figure the order of magnitude of the equilibrium constant at the temperature of our experiments can be roughly estimated as Kn-C7 ) 1 Torr-1 ) 7.5 × 10-3 Pa-1.

With reference to n-pentane, on the basis of direct measurements, the value of the saturation loading ∞ capacity has been set to Γn-C5 ) 0.85 mol/kg. In order to determine the value of Kn-C5, the measured selectivity between n-heptane and n-pentane is used. This selectivity is concentration dependent, with an average value of about 2.5. This indicates that in this separation the desorbent is strong; i.e., it is more adsorbable than both components of the mixture to be separated. Since selectivity between components i and j is defined as Sij ) (Γi/Γj)/(ci/cj), in the case of the nonstoichiometric Langmuir isotherm (1) it is given by Sij ) Γ∞i Ki/ ∞ (Γ∞j Kj). Thus, using the known values of Γn-C7 , Kn-C7 ∞ and Γn-C5, for the measured value of selectivity the equilibrium constant Kn-C5 can be calculated, thus obtaining Kn-C5 ) 2.6 × 10-3 Pa-1. No literature data are available on isopentane, other than the trivial information that it is nonadsorbable on 5A-zeolite crystals (Barrer, 1978). As a matter of fact, a very small adsorptivity is measured on zeolite pellets, probably due to the presence of the clay binder. Precise measurements are very difficult due to the very low values of its adsorbed phase concentration. Selectivity between n- and isopentane, which is notably a very sensitive parameter, varies from one experiment to the other between 20 and 100. Thus, only rough estimates are possible, and the following ones have been se∞ lected: Ki-C5 ) 0.75 × 10-3 Pa-1 and Γi-C5 ) 0.03 mol/ kg. Multicomponent Mixtures. In the case of the separation of a multicomponent feed stream a shortcut procedure should be adopted to design the operating conditions to achieve the desired performance (Mazzotti et al., 1994). This is based on the concept of strong-key and weak-key components, which is derived from the theory of multicomponent distillation. Since the separation between the two key components is more difficult than the real multicomponent separation, the complete separation region in the operating parameter space can be determined with reference to this binary separation and then applied to the real case. Through breakthrough experiments we have determined that the components of the feed mixture can be put in increasing order of adsorptivity as follows: isopentane, 2- and 3-methylpentane (which have very similar adsorptivities), methylcyclopentane, and cyclohexane (which are similar too), n-pentane, n-hexane. In principle, the shortcut binary separation should involve cyclohexane and n-pentane as weak- and strong-key components, respectively. In practice, due to the difficulty in obtaining precise equilibrium measurements for components with very low adsorptivity such as cyclohexane, the binary separation discussed above is taken as reference separation and the regions of separation determined for this are used as approximate regions also for the multicomponent case. Experimental Results In this section the separation performances achieved during the experimental runs in the SMB pilot unit are reported and discussed. Particular attention is devoted to the interpretation of the experimental findings on the base of the design criteria derived from the equilibrium theory model mentioned above. During all the experiments the temperature has been kept constant and equal to 448 K. On the contrary, small variations in pressure are exhibited from run to run, with values ranging from 300 to 400 kPa. In all

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the experimental runs a switch time of t* ) 180 s was adopted. Moreover, the values used for the physical properties in eq 2 are as follows: V ) 176.7 × 10-6 m3, * ) 0.53, Fs ) 1665 kg/m3, Γ∞ ) 0.8 mol/kg, and Ff ) 100 mol/m3. During these experiments, dealing with linear/ nonlinear paraffin separation, the pilot unit has exhibited a dynamic behavior which is qualitatively similar to that described by Storti et al. (1992) in the context of xylene separation. A key aspect is that the stationary regime of the unit is actually a cyclic steady state, and therefore, its correct experimental identification is mandatory. This is done according to the following criteria (cf. Storti et al., 1992): (1) average compositions of extract and raffinate streams must be constant; (2) overall and single component mass balances must be fulfilled; (3) instantaneous composition profiles in extract and raffinate must exhibit a periodic behavior. When all these criteria are satisfied it is assumed that the cyclic steady state has been achieved. The average compositions of the outlet streams are used to calculate appropriate performance parameters (Storti et al., 1995). In the case of the binary separation of a mixture constituted of the strong component A and the weak component B, these are defined as follows:

a

b

purities in the extract and raffinate: PE )

PR )

cEA cEA + cEB cRB cRA + cRB

(3)

(4)

recovery of A in the extract and of B in the raffinate:

RA )

RB )

cEAE cFAF cRBR cFBF

(5)

(6)

desorbent and adsorbent requirement: m1 Dr ) m3 - m2 Ar )

V* + (V/2)σ(1 - *)(m3 + m2) Vσ(1 - *)(m3 - m2)

(7)

(8)

productivity, which is defined as the ratio between the feed mass flow rate and the overall mass of adsorbent in the SMB unit (Nac is the number of adsorption columns, Nac ) 6 in this case):

PR )

F NacV(1 - *)Fs

(9)

The theoretical analysis developed in previous works (Storti et al., 1995; Mazzotti et al., 1996) allows us to identify different regions in the operating plane (m2, m3), each of them characterized by the achievement of different separation performance. With reference to Figure 4a, the central triangle-shaped region is the complete separation region, corresponding to operating

Figure 4. Relative position of the experimental operating points (b) in the (m2, m3) plane with respect to the calculated regions of separation. Points F and U (O) are the only ones for which the constraint on m1 is not fulfilled. (a) Separation of n- and isopentane. (b) C5-C6 multicomponent separation. Operating conditions and measured separation performances are reported in Tables 2 and 3, respectively.

conditions for which complete separation is achieved, i.e., PE ) PR ) RA ) RB ) 100%. On the right-hand side of this, in the pure extract region, 100% purity in the extract is obtained whereas the raffinate is polluted with the strong component A. The recovery of A, RA, decreases from 100% to 0% while going from the leftto the right-hand boundary of the pure extract region. On the left-hand side of the complete separation region the pure raffinate region is located, where PR ) 100% while PE is smaller. The recovery of component B, RB, decreases while moving from the boundary with the complete separation region, where RB ) 100%, toward the vertical axis. In the region at the upper left corner of the (m2, m3) plane neither outlet stream is pure. For constant values of the flow rate ratio m1 and the switching time, t*, desorbent requirement, Dr, and productivity, PR, are constant along straight lines with unitary slope in the (m2, m3) plane. On the other hand, adsorbent requirement, Ar, is constant along straight lines passing through the origin of the axes. All these parameters improve, i.e., Dr and Ar decrease and PR increases, while moving from the diagonal toward the vertex of the complete separation region. It follows that

Ind. Eng. Chem. Res., Vol. 35, No. 7, 1996 2319 Table 1. Separation of n-Pentane and Isopentane: Experimental Operating Conditions and Composition of the Outlet Streams run

F

H

L

M

N

O

P

Q

R

S

n-CF5 (wt %) i-CF5 (wt %)

32.5 67.5 107 29 55 117 3.6 21.4 9.6 1.3

33.0 67.0 151 34 53 92 5.1 8.5

32.0 68.0 150 81 95 95 5.9

32.0 68.0 154 90 110 109 9.3

30.5 69.5 151 38 53 95 4.3 6.9

33.3 66.7 162 115 125 91 2.7

29.7 70.3 149 41 144 87 29.1 10.2

29.6 70.4 151 70 176 105 37.7

30.4 69.6 154 80 145 98 27.4

30.2 69.8 159 98 165 92 2.8

3.6

9.8

11.9

4.5

0.5 42.2

30.6

1.7 27.6

G1 (g/h) G2 (g/h) G3 (g/h) Ff (mol/m3) n-CE5 (wt %) i-CE5 (wt %) n-CR 5 (wt %) i-CR 5 (wt %)

1.6 5.7

41.8

Table 2. Separation of n-Pentane and Isopentane: Experimental Flow Rate Ratios and Values of the Separation Performance Parametersa run

F

H

L

M

N

O

P

Q

R

S

m1 (m1,cr) m2 m3 PE (%) PR (%) Rn-C5 (%) Ri-C5 (%) Dr Ar PR (1/h)

0.419 (0.386) 0.042 0.166 14 12 35 4 4.1 1.6 0.031

0.632 (0.377) 0.082 0.170 38 100 100 16 7.9 2.3 0.023

0.622 (0.395) 0.297 0.361 100 100 100 100 10.7 6.4 0.016

0.638 (0.422) 0.336 0.427 100 100 100 100 7.7 5.2 0.023

0.628 (0.373) 0.095 0.166 38 100 100 23 10.1 3.0 0.018

0.675 (0.450) 0.454 0.503 100 79 37 100 16.2 11 0.013

0.707 (0.502) 0.137 0.692 74 100 100 86 1.4 0.9 0.124

0.703 (0.530) 0.280 0.838 100 99 97 100 1.4 1.1 0.128

0.695 (0.489) 0.321 0.645 100 100 100 100 2.4 1.7 0.078

0.722 (0.520) 0.418 0.751 100 94.6 86 100 2.4 2.0 0.080

a

The critical value of the flow rate ratio in section 1, m1,cr, is reported in parentheses below the operating value, m1.

the vertex represents the optimal operaing point for a given separation. Separation Performance in the Case of Binary Mixtures. Let us first consider the experimental runs dealing with the binary separation of iso- and n-pentane, whose operating conditions are summarized in Table 1. These include the concentrations in the feed stream, nCF5 and i-CF5 , the mass flow rates in the three sections of the unit, Gj, and the fluid density, Ff, whose variations are due to pressure changes. In the same table the concentrations measured in the extract, n-CE5 and iCE5 , and in the raffinate, n-CR5 and i-CR5 , are also reported. Note that no numerical value indicates that the corresponding concentration is below the least analytically detectable limit. From the analysis of the data in Table 1 it can be seen that in six runs out of 10, i.e., in runs H, L, M, N, P, and R, the raffinate is pure; that is, no n-pentane has been detected in the raffinate stream. Moreover, in runs L, M, and R complete separation has been achieved, since the extract is also found to be pure, i.e., free of isopentane. Depending upon the operating conditions, the concentrations of iso- and n-pentane in the outlet streams are rather low, such as in runs F, H, L, M, N, O, and S, or else it reaches pretty high values, as is the case of runs P, Q, and R. Run Q is especially noticeable, since the concentration of n-pentane in the extract is about 20% higher than in the feed stream. In Table 2, the information about the operating conditions in the experimental runs is summarized in terms of flow rate ratios mj, calculated using eq 2, where the average molar mass of the fluid phase has been computed as the weighted average of the two inlet streams. In the same table the values of the performance parameters defined above are also indicated. In Figure 4a, the relative position in the (m2, m3) plane of the points corresponding to the operating conditions of the experimental runs is compared with the regions of separation calculated following the approach outlined above.

By analyzing Figure 4a together with the purity values in Table 2, it can readily be noted that all the operating points, with the exception of points F and Q, fall in the correct separation region. In particular, points H, N, and P are in the pure raffinate region, points O and S are in the pure extract region, and points L, M, and R are in the complete separation region. The operating point Q is very close to the vertex of the triangle-shaped complete separation region, and in particular, it lies very close to the right hand boundary with the pure extract region. It is apparent that the vertex represents the optimal operating point, but also the least robust one, where even the slightest change in the operating conditions may produce a drastic variation in the separation performance (Storti et al., 1993a). This explains why in run Q only 99% purity in the raffinate has been obtained, due to a 0.5 wt % impurity of n-pentane, while 100% purity has been achieved in the extract. In run F no pure raffinate is obtained, contrary to what can be predicted by considering the position of the corresponding operating point in the (m2, m3) plane. This apparently anomalous behavior can be explained by also considering the constraint on the operating parameter m1. The calculation of the critical purge ratio, m1,cr, for each run is not straightforward, since it requires a one-dimensional numerical search on the straight line spanned by the parameter m1. The calculated values are reported between brackets in Table 2, thus showing that in all experimental runs, not excluding run F, the operating parameter m1 is greater than the corresponding critical purge ratio. However, the flow rate ratio m1 in run F is 30% smaller than in all the other runs; furthermore, it is only 10% larger than the critical value. Under these conditions the effect of experimental disturbances and model uncertainties is significant and the predicted separation performance is spoiled. Thus summarizing, in run F the constraint on the purge ratio m1 is not fulfilled. Therefore, in this run the adsorbent is not properly

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Table 3. C5-C6 Multicomponent Separation: Experimental Flow Rate Ratios and Values of the Separation Performance Parametersa run

T

U

V

X

Y

m1 (m1,cr) m2 m3 PE (%) PR (%) RL (%) RNL (%) Dr Ar PR (1/h)

0.64 (0.429) 0.34 0.44 99.3 93 86 99.6 7.3 4.8 0.026

0.62 (0.394) 0.04 0.19 32 53 76 16 4.7 1.3 0.038

0.81 (0.400) 0.08 0.19 58 97 99 22 8.2 2.0 0.028

0.86 (0.475) 0.21 0.66 92 99.7 98 96 2.1 1.1 0.110

0.86 (0.508) 0.38 0.78 99.7 94 89 99.8 2.3 1.7 0.098

a The critical value of the flow rate ratio in section 1, m 1,cr, is reported in parentheses below the operating value, m1.

regenerated, and in agreement with the theoretical results the prediction of the separation performance based on the position of the operating point in the (m2, m3) plane is not valid. With reference to recovery, the values reported in Table 2 are in very good agreement with the prediction that can be made by considering the position of the operating points in the (m2, m3) plane. In particular, it is consistent with model predictions that in run O the recovery in the extract is much lower than in run S, whereas it is almost complete in run Q, which is very close to the complete separation region. Similarly, the recovery in the raffinate is much higher in run P than in runs H and N (between these last two runs, the latter is better than the former because it is closer to the complete separation region). Finally, the effect of changes in the operating parameters mj on the performance parameters representing desorbent requirement, adsorbent requirement, and productivity also corresponds to the prediction of the theory. As expected, all these parameters increase while moving the operating point in the (m2, m3) plane away from the diagonal toward the vertex of the triangle-shaped region. Accordingly, the best values for these are achieved in runs P and Q. Note that in these runs only 1.4 g of desorbent is required for each gram of feed mixture separated and that a productivity larger than 0.12 g per gram of zeolite per hour has been achieved. Separation Performance in the Case of Multicomponent Mixtures. In the case of the five experimental runs dealing with the separation of a sevencomponent feed stream, the same experimental setup, bed characteristics, and operating conditions (i.e., temperature, pressure, and switching time) as in the binary experiments have been adopted. The most important difference concerns the desorbent purity which in this set of experiments was rather poor. In particular, the analysis showed that this contained 0.08% (by weight) of isopentane, 0.24% of n-pentane, and 0.05% of nhexane. These impurities are necessarily found in the wrong outlet stream: the isopentane present in the desorbent pollutes the extract, whereas n-pentane and n-hexane end up in the raffinate, thus spoiling the separation. The flow rate ratios, mj, and the separation performance parameters (which are now referred to the whole linear and nonlinear fractions) are summarized in Table 3. In Figure 4b the operating points corresponding to the experimental runs and the same regions of separation considered in Figure 4a are shown in the (m2, m3) plane. By analyzing the relative position of the operat-

ing points with respect to the complete separation region and considering the separation performances reported in Table 3 it can be concluded that also in this case the operating points fall in the correct separation region, as discussed in detail in the following. The operating point corresponding to run T is just outside the boundary of the complete separation region on the pure extract side. This agrees well with the experimental purities achieved, particularly when it is considered that 99.3% purity in the extract is due to the isopentane impurity in the desorbent. Point T is rather close to the straight line m3 ) m2 and then as expected the other performance parameters are rather poor (i.e., desorbent and adsorbent requirements and productivity). The same happens in the case of run V, which is in the pure raffinate region in agreement with the observed purity values. Also in this case linear impurities in the desorbent hinder raffinate purity. In the case of run U the constraint on the flow rate ratio m1 is not fulfilled, even though its value is larger than the critical purge ratio predicted by the model. It is worth noticing that the critical values reported in Table 3 are calculated with reference to the separation of the strong-key and weak-key components, thus being less accurate than in the binary case analyzed in the previous section. Thus, in run U the situation is similar to that discussed above in the case of run F for binary separations. Experimental runs X and Y correspond to highperformance separations (in terms of desorbent and adsorbent requirements and productivity), as expected by the position of the operating points in the plane (m2, m3), which are closer to the optimal vertex of the triangle than to the diagonal. In particular, slightly more than 2 g of desorbent per gram of feed mixture are consumed and a productivity of about 0.1 grams per g of zeolite per hour is achieved. In both runs X and Y, 99.7% purity is achieved in the raffinate and in the extract, respectively, in good agreement with theoretical predictions. Finally, let us focus on runs V and X, which are both located in the pure raffinate region. As expected, the purity in the extract is poor, whereas that in the raffinate is high, even though it is not 100% as it happens in the binary case under similar conditions (run V can be compared with run H and run X with run P). By looking at the composition of the raffinate, it can be readily concluded in both cases that the linear component impurities in the raffinate are only those which are present in the fresh desorbent itself. However, since in run X the desorbent requirement is low, about four times lower than in run V, it follows that the purity in the raffinate is larger, i.e., 99.7% instead of 97%. Similar observations can be made if the extract purities in runs T and Y are considered. This analysis confirms that whenever the desorbent has some impurities, these are carried directly to the wrong outlet streams. Thus, the important conclusion with respect to applications is that in these cases minimizing the desorbent consumption is also a mean to increase purity. This is a further advantage of countercurrent or SMB separation processes which have the peculiarity of strongly reducing desorbent requirement with respect to any other adsorptive separation technique. Concluding Remarks In this work we have investigated the separation of several mixtures of the C5-C6 paraffin fraction in a

Ind. Eng. Chem. Res., Vol. 35, No. 7, 1996 2321

simulated moving bed laboratory unit. The process is operated in the vapor phase at 448 K and about 350 kPa. The interest of the process from the applicative viewpoint stems from two considerations: the low operating temperature makes the vapor phase operation particularly attractive with respect to liquid phase operation in terms of heat integration, with no shortcomings due to zeolite deactivation. Moreover, this operation fits rather well in processes for gasoline octane enhancement, where this separation stage follows the isomerization reaction which is carried out in a gas phase catalytic reactor. On the basis of the analysis of the adsorption equilibrium behavior of the system under examination, constituted of the mixtures to be separated, n-heptane as desorbent and 5A-zeolite as adsorbent solid, a threesection open-loop configuration has been adopted. Experimental runs have dealt with the separation of a binary feed stream (n-pentane 30% and isopentane 70%) and a multicomponent C5-C6 mixture. In all cases the separation performances achieved are rather promising from the applicative point of view, in terms both of purity of the outlet streams and of various performance indices, such as recovery of the desired components, desorbent and adsorbent requirement, and productivity. The experimental results agree nicely with the theoretical predictions on which the design criteria used in developing the separations are based. Acknowledgment We gratefully acknowledge the financial support of CNR Progetto Finalizzato Chimica Fine II. Nomenclature Ar ) adsorbent requirement, defined by eq 8 ci ) fluid phase molar concentration of component i Dr ) desorbent requirement, defined by eq 7 E ) extract mass flow rate F ) feed mass flow rate Gj ) fluid mass flow rate in section j of a SMB unit Ki ) adsorption equilibrium constant of component i L ) column length Lj ) section length, TCC unit mj ) flow rate ratio in section j Mav ) average molar mass of the fluid phase nj ) number of ports in section j, SMB unit Nac ) overall number of columns NC ) number of components NS ) number of sections PE ) desorbent free extract purity, defined by eq 3 PR ) desorbent free raffinate purity, defined by eq 4 PR ) productivity, defined by eq 9 R ) raffinate mass flow rate Ri ) recovery of the i-th component, defined by eqs 5 and 6 Sij ) selectivity between components i and j t* ) switching time V ) column volume yi ) fluid phase dimensionless concentration of component i, yi ) ci/Ff

θi ) adsorbed phase dimensionless concentration of component i, θi ) Γi/Γ∞ Ff ) fluid phase reference molar density Fs ) bulk solid mass density σ ) capacity ratio, σ ) Fs Γ∞/Ff Subscripts and Superscripts A,B ) components to be separated (A stronger than B) D ) desorbent E ) extract F ) feed i ) component index j ) section index l ) component index R ) raffinate

Literature Cited Barrer, R. M. Zeolites and clay minerals as sorbents and molecular sieves; Academic Press: London, 1978. Broughton, D. B.; Gerhold, C. G. Continuous sorption process employing fixed beds of sorbent and moving inlets and outlets. U.S. Patent 2,985,589, 1961. Doetsch, I. H.; Ruthven, D. M.; Loughlin, K. F. Sorption and diffusion of n-heptane in 5A zeolite. Can. J. Chem. 1974, 52, 2717. Haber, J.; Najbar, J.; Pawelek, J.; Pawlikowska-Czubak, J. Thermodynamics and kinetics of the adsorption of n-heptane on type 5A molecular sieves. J. Colloid Interface Sci. 1973, 45, 252. Johnson, J. A.; Kabza, R. G. SORBEX: industrial-scale adsorptive separation. In Preparative and Production Scale Chromatography; Ganetsos, G., Barker, P. E., Eds.; Marcel Dekker: New York, 1993; p 257. Mazzotti, M.; Storti, G.; Morbidelli, M. Robust design of countercurrent adsorption separation processes: 2. Multicomponent systems. AIChE J. 1994, 40, 1825. Mazzotti, M.; Storti, G.; Morbidelli, M. Robust design of countercurrent adsorption separation processes: 3. Nonstoichiometric systems. AIChE J. 1996, in press. Paludetto, R.; Storti, G.; Gamba, G.; Carra`, S.; Morbidelli, M. On multicomponent adsorption equilibria of xylene mixtures on zeolites. Ind. Eng. Chem. Res. 1987, 26, 2250. Raghuram, S.; Wilcher, S. A. The separation of n-paraffins from paraffin mixtures. Sep. Sci. Technol. 1992, 27, 1917. Ruthven D. M. Principles of adsorption and adsorption processes; John Wiley: New York, 1984. Ruthven, D. M.; Ching, C. B. Counter-current and simulated counter-current adsorption separation processes. Chem. Eng. Sci. 1989, 44, 1011. Storti, G.; Mazzotti, M.; Furlan, L. T.; Morbidelli, M.; Carra`, S. Performance of a six port Simulated Moving Bed pilot plant for vapor-phase adsorption separations. Sep. Sci. Technol. 1992, 27, 1889. Storti, G.; Mazzotti, M.; Morbidelli, M.; Carra`, S. Robust design of binary countercurrent adsorption separation processes. AIChE J. 1993a, 39, 471. Storti, G.; Mazzotti, M.; Furlan, L. T.; Morbidelli, M. Analysis of a six port Simulated Moving Bed separation unit. In Proceedings of the Fourth International Conference on Fundamentals of Adsorption; Suzuki, M., Ed.; Kodansha: Tokyo, Japan, 1993b; p 607. Storti, G.; Baciocchi, R.; Mazzotti, M.; Morbidelli, M. Design of optimal operating conditions of simulated moving bed adsorptive separation units. Ind. Eng. Chem. Res. 1995, 34, 288.

Received for review December 19, 1995 Revised manuscript received April 9, 1996 Accepted April 10, 1996X

Greek Letters Γi ) adsorbed phase molar concentration of component i Γ∞i ) adsorbed phase saturation molar concentration of component i Γ∞ ) adsorbed phase reference molar concentration  ) external void fraction p ) intraparticle void fraction * ) overall void fraction, * )  + p(1 - )

IE950766L

Abstract published in Advance ACS Abstracts, May 15, 1996. X