Design Guidelines for Solid-Catalyzed Reactive Distillation Systems

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Ind. Eng. Chem. Res. 1999, 38, 3696-3709

Design Guidelines for Solid-Catalyzed Reactive Distillation Systems Hoshang Subawalla† and James R. Fair* Separations Research Program, University of Texas at Austin, Austin, Texas 78712-1062

In this paper we discuss design guidelines for solid-catalyzed reactive distillation systems. The guidelines are used to generate initial estimates for column pressure, reactive zone location, catalyst mass, reactant feed location, reactant ratio, reflux ratio, column diameter, number of equilibrium stages, and packed height. They form a part of a methodical design procedure that makes extensive use of both nonequilibrium (rate-based) and equilibrium-stage simulation models. Important choices prior to design include selection of reliable thermodynamic and reaction kinetic models. We tested the guidelines for two etherification systems and validated them experimentally for a hydration reaction. The results from a case study, the manufacture of tertamyl methyl ether, are shown here. Superimposing reaction on separation leads to unique design trade-offs. Thus, column diameter depends both on maximum vapor velocity and on packing catalyst density, reactant ratios are a function of conversion and azeotrope formation, the operating pressure affects the relative volatility, chemical equilibrium, and reaction rate (reactive zone temperature), and the reflux ratio impacts both separation and conversion. The guidelines and procedures presented here simplify the detailed reactive column design considerably. Introduction Reactive separation processes such as reactive distillation, sorption-enhanced reaction, reactive absorption, and reactive crystallization combine the essential tasks of reaction and separation in a single vessel. In reactive distillation and sorption-enhanced reaction, the separation of products from reactants enhances conversion, while superimposing reaction on separation enhances the mass-transfer rate in reactive absorption and crystallization. Combining reaction and distillation is possible when reaction rates are comparable to those in a reactor at pressures suitable for distillation. Reactive systems where reactants have similar relative volatilities, and where product and reactant volatilities differ considerably, are ideally suited for reactive distillation. Equilibrium-limited reactions are also excellent candidates for reactive distillation because conversion is enhanced with product removal and the resultant shift in equilibrium. Under normal circumstances these reactions would require a large excess of one reactant with resultant recycle. Reactive distillation is also useful when the desired product undergoes further reaction because product removal reduces product degradation. The equipment used for homogeneous reactive distillation processes consists of bubble-cap or sieve trays with high weirs that provide the necessary liquid holdup and residence time needed for reaction. Heterogeneous reactive distillation processes use a solid catalyst, where most of the reaction takes place within the catalyst particle. The equipment consists primarily of catalystcontaining packing which allows for simultaneous reaction and mass transfer between vapor and liquid phases. Heterogeneous reactive distillation is often termed catalytic distillation to differentiate between the two processes (heterogeneous and homogeneous). Packing * To whom all correspondence should be addressed. Email: [email protected]. Fax: (512) 471-7060. † Present address: Air Products and Chemicals Inc., Chemicals Division (Escambia Plant), P.O. Box 467, Pensacola, FL 32592.

and tray designs used for catalytic distillation are substantially different from conventional packing and tray geometries. These novel designs have two important features: they maximize reaction efficiency by providing intimate contact between vapor, liquid, and solid catalyst and simultaneously minimize pressure drop through the catalyst bed. Two commercially used packing types are catalytic bale1 and structured packing.2 Other catalytic packing types such as structured carpet packing3 and ring-shaped ion-exchange resins4 have also been used. Currently, the largest users of reactive distillation technology are fuel-ether producing units. A variety of ethers can be produced by reacting iso-olefins having four, five, or six carbon atoms with either methanol or ethanol. Two of the most popular ethers are methyl tertbutyl ether (MTBE) produced by the reaction of isobutylene and methanol and tert-amyl methyl ether (TAME) produced by reacting a mixture of isoamylenes (2methyl-2-butene and 2-methyl-1-butene) with methanol. Other applications include methyl acetate manufacture from methanol and acetic acid5 and hydrogenation of butadienes to butenes.6 Processes under consideration include the decomposition of ethers to high-purity olefins, alkylation of aromatics, trans-esterification, dimerization, and hydroisomerization. Details of these and several other processes are given by Stadig.7 Previous Work Research in reactive distillation design is mainly limited to synthesis rather than detailed design. Transformed variables have been developed and used extensively for the synthesis and design of reactive distillation systems.8-10 The transformed variables have been extensively used to develop residue curve maps11,12 and distillation line diagrams13 for both equilibrium-controlled and kinetically-limited systems and for chemicalequilibrium-controlled reactive column design.14 Geometric fixed-point methods have also been used for the design of kinetically-controlled reactive columns.15,16

10.1021/ie990008l CCC: $18.00 © 1999 American Chemical Society Published on Web 08/31/1999

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While these tools are very useful for preliminary screening and design, they cannot be used for detailed design because of some of their limiting assumptions (constant molar liquid holdup on each stage and vapor-liquid equilibrium). Besides, these tools do not account for the specialized nature of reactive column internals. Detailed design (postsynthesis) is discussed by Sneesby et al.17 for ether systems and by Jakobsson et al.18 for the TAME reaction. Sneesby et al. discuss design trade-offs specifically for the ethyl tert-butyl ether (ETBE) system. However, their design strategy does not account for the effect of column internals on design. Jakobsson et al. provide useful guidelines for selecting the correct type of packing (nonreactive packing, catalytic packing with high separation efficiency, and catalytic packing with low separation efficiency). While they do account for the effect of column packing on design, they do not discuss the effect of other design parameters such as the number of stages, reactant ratio, column pressure, or reflux ratio. In this paper we present design guidelines for solidcatalyzed reactive distillation systems. They are used to generate initial estimates for column pressure, reactive zone location, catalyst mass, reactant feed location, reactant ratio, reflux ratio, column diameter, number of equilibrium stages, and packed height. They can also be used to ascertain the need for an upstream reactor. The guidelines constitute a methodical design procedure and are useful for screening viable designs for a particular reactive system. We do not discuss reactive distillation synthesis in this paper. Instead, we focus our efforts on the postsynthesis step, i.e., the detailed design of a process featuring a reactive column for which reactive distillation is judged to be both feasible and advantageous. We make extensive use of simulation tools such as nonequilibrium (rate-based) and equilibrium-stage models. While the traditional equilibrium-stage model assumes that vapor leaving a stage is in equilibrium with liquid on that stage, the rate-based model assumes equilibrium only at the vapor-liquid interface and uses mass- and heat-transfer correlations to determine molar flux across the interface. The rate-based model uses the actual number of trays or packed height as opposed to ideal equilibrium stages. Important choices made prior to design include an accurate vapor-liquid-equilibrium (VLE) model and a reliable reaction kinetics model. The kinetic model should accurately predict reaction rates under reactive distillation conditions, particularly for systems where reactive column conditions differ considerably from the typical test or operating conditions. For example, cumene is typically synthesized by reacting benzene with propylene in a vapor-phase fixed-bed reactor at high temperatures and pressures. The same reaction when performed in a reactive column19 requires lower pressures and temperatures. Besides, one reactant (benzene) is mainly present in the liquid phase. Thus, not only is there a change in the reaction conditions but also a possible change in the reaction mechanism. Both VLE and kinetic model parameters are normally regressed from an experimental data set. To ensure accuracy in design, the uncertainty in model parameters must be evaluated. Pilavachi et al.20 provide a systematic framework to evaluate parametric uncertainty for reactive column simulation and design. While we did not perform an exhaustive parametric analysis, we did

Figure 1. Intrinsic MTBE reaction rate as a function of isobutene conversion at different temperatures and stoichiometric reactant ratio (methanol/isobutene ) 1).

compare thermodynamic and kinetic models and parameters from different sources prior to their use for the TAME case study. This paper has three parts. Initially, we present and discuss the guidelines from a generic standpoint. We recommend that the reader follow the guidelines in the order defined. Next, we show how these guidelines are used to design a TAME manufacturing process that consists of a reactor and a reactive column. Last, we present results from this case study and discuss use of the guidelines for other systems. Our experience at applying these guidelines for three systems of industrial interest, two etherification (MTBE and TAME) systems and a proprietary trans-esterification system, have been favorable. We also experimentally tested our models with data from a hydration reaction.21 However, it would be presumptuous of us to claim that the guidelines are universally applicable; each system has certain characteristics that make the use of one or more these guidelines difficult. In particular, we recommend caution for systems that are not equilibrium-limited and/ or where one reactant is a vapor or gas (e.g., hydrogenation of butadienes). For equilibrium-limited systems, a clear advantage of reactive distillation is conversion enhancement with product removal. For systems not limited by chemical equilibrium, the advantages are less obvious and include increased selectivity to the desired product, reduced equipment cost, and increased catalyst life. Design Guidelines Use of a Prereactor. For many reactive systems that are not severely equilibrium-limited, reaction rates are high for a large fraction of the conversion range. They decline abruptly only when compositions and temperatures approach equilibrium. For such cases a plug-flow reactor (PFR) that handles a substantial part of the reaction duty should be used upstream of the reactive column. Figure 1 shows MTBE formation rates as a function of equilibrium conversion and temperature. The rates were calculated from kinetics provided by Rehfinger and Hoffmann.22 The results indicate that rates decline rapidly when conversion exceeds 65-70%. Therefore, a PFR operating in the 70-80 °C temperature range could profitably be used to achieve a majority of the conversion. Many MTBE processes in industry

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use two reactors (operating at different temperatures) in series. The effluent from the first reactor (normally adiabatic) approaches equilibrium within 1-5% and is cooled before being fed to a second (normally isothermal) reactor operated at a lower temperature. Typical conversions for a two-reactor system range from 90 to 95%.23 When a prereactor is used in conjunction with a catalytic distillation column, reactor conversion is typically 88-92%, and the total conversion of the reactorcolumn pair approaches 100%. Operating Pressure. The operating pressure range depends on the condenser coolant and reboiler heating medium temperatures. Normally, condenser pressure is set such that a sufficiently large temperature driving force exists for water to be used as the coolant. The appropriate heating medium (normally steam at a fixed pressure) is selected after accounting for the column pressure drop. Within this range the operating pressure depends on the reaction temperature, relative volatility, and the effect of pressure on azeotropes. The column and reactive zone temperatures increase with increasing pressure while the relative volatility declines. Since the reaction rate increases exponentially with temperature, a small increase in pressure will result in a substantial increase in rate when the system is kinetically-limited. For intraparticle diffusion-limited systems, there is a trade-off. While the intrinsic rate increases with temperature, the diffusion resistance within the catalyst pellets also increases and the effective rate is often much lower than the intrinsic rate. Temperature changes impact the reaction rate only marginally when either vapor-liquid or liquid-solid transport is rate-limiting. The effect of temperature on the reaction equilibrium is also very important. For exothermic reactions equilibrium shifts toward the reactants (product decomposition) with increasing temperature. Thus, a temperature increase often results in a decline in conversion. Departure from chemical equilibrium also decreases with increasing temperature, thereby decreasing the reaction driving force and eventually causing a decline in the reaction rate.24 For endothermic reactions operation at the maximum possible pressure is beneficial, both from a chemical equilibrium and reaction rate standpoint. Changes in the relative volatility with operating pressure are more gradual and small variations do not affect reactant zone concentrations or rate substantially. However, pressure-sensitive azeotropes may form or disappear with changes in pressure and the exact effect depends on the type of azeotrope (minimum or maximum boiling), azeotrope composition (reactant/product or reactant/reactant), and the effect of pressure on azeotrope composition. Pressure also alters the boiling points and volatilities of components. Thus, methanol has a higher boiling point than MTBE at pressures below 3 bar. Above this pressure MTBE is the highest boiler. From an economic standpoint the vessel wall thickness increases with increasing pressure. However, the column diameter decreases because of the increase in vapor density. Excess reactant is required to maintain conversion when the composition of that reactant in an azeotropic product increases with pressure. For such cases a pressure increase concentrates the reactant away from the reactive zone, thus causing substoichiometric reactant ratios. For example, the methanol content of the C4-methanol (olefin and paraffin-methanol) azeotropes in the MTBE system increases with increasing pressure,

thus requiring the use of reactant ratios (methanol/ isobutene) in the range 1.1-1.25 (stoichiometric reactant ratio ) 1). For most systems the operating pressure is driven by temperatures and reaction rates such that we maximize the rates and minimize the residence times and catalyst mass. However, there is an upper pressure limit beyond which rates actually decrease because of reactant depletion due to relative volatility and/or azeotropic effects, and a decrease in the departure from chemical equilibrium. Reactive Zone Location. The reactive zone location depends on the reactant and product volatilities. The reactive section should be located where the concentration of at least one reactant (preferably the limiting reactant) is the maximum. It is located toward the top of a column when the limiting reactant is the most volatile component in the system and the product is the least volatile component in the system. Examples of such reactive systems include benzene alkylation where propene, the limiting reactant, is the most volatile component and cumene, the product, is the heavy component. Other examples include methyl acetate hydration where acetic acid (a coproduct with methanol) is one of the heavier components while methyl acetate is the most volatile. Stringent purity specifications on either the distillate or bottoms product may often require the use of a rectification or stripping section above or below the reactive zone to remove trace amounts of products. When the product is the most volatile component, the procedure is reversed. For a four-component system with two intermediate-boiling reactants, the reactive zone should be located at the center of the column. The less volatile product is removed at the bottom and the more volatile product leaves with the distillate stream. Jakobsson et al.18 provide useful guidelines for appropriate reactive zone location. Reactant Ratio and Feed Location. Stoichiometric reactant amounts in the reactive zone favor high conversion and are maintained by ensuring that the quantity of reactants present is sufficient to meet both reaction and azeotropic requirements. For systems that are reaction-rate or chemical-equilibrium-limited, vaporliquid mass-transport rates exceed reaction rates considerably.25 For such cases a reactant-containing azeotrope may force the reactant away from the reactive zone prior to its use in the reaction, thereby causing substoichiometric reactant ratios. Reactant requirements can be calculated from product specifications, desired conversion, and azeotropic compositions at different operating pressures. We recommend that special attention be given to azeotropes as they have a profound effect on column reaction and separation performance. For example, even though methanol is one of the high boilers in a MTBE column, its azeotropes with the C4 (olefin and paraffin) components are all minimumboiling azeotropes. Therefore, thermodynamic nonideality forces methanol to behave like a low-boiling component. A judicious choice of feed location guarantees high concentrations of reactants in the reactive zone. If reactants are the most volatile components in the system, then the feed should be introduced at the bottom of the reactive zone. Relative volatility effects force reactants to move toward the condenser and enrich the reactive zone. When products are more volatile than reactants, the reactants should be fed at the top of the


reactive section. If the volatilities of the reactants are very different, a subsidiary feed location may be required to ensure stoichiometric quantities of reactants in the reactive zone. Fresh (unreacted) feed streams should be fed directly to the reactive zone, to maximize reactant concentrations in this region. Feed location is adjusted to minimize product decomposition when a prereacted product-containing stream is fed to a reactive column. Such a stream should enter the column at some distance from the reactive zone, thereby ensuring that separation between products and reactants occurs on stages between the feed point and the reactive zone. Catalyst Mass. Selecting an appropriate amount of catalyst is an important aspect of reactive column design. Inadequate catalyst volume reduces the residence time and gives poor conversion. For equilibriumlimited systems, excess catalyst, particularly when it is inappropriately located, may result in product decomposition. For such cases simultaneous reaction and separation leads to compositions and temperatures that favor the reverse reaction. This phenomenon is different from conventional plug-flow and stirred-tank reactors where temperatures and compositions remain constant once equilibrium is attained. Equilibrium in conventional reactors can only be disturbed by changing temperatures and/or pressures, or by separating products from reactants. At a microscopic level any point in the reactive zone may be represented by a combination of an infinitesmally small ideal reactor (no diffusion resistance) and an ideal separator. Thus, a reactive stage or segment is a summation of infinite such reactor-separator blocks. For a given temperature and desired conversion a plugflow reactor (PFR) requires less catalyst than a mixedflow reactor. An ideal separator would completely separate reactants from products. Therefore, a PFRideal separator pair represent an ideal reactive stage with reaction and separation duties decoupled. A similar concept termed HETES (height equivalent of a theoretical stage) is used by D’Amico et al.26 We determine minimum catalyst requirements by simulating a series of isothermal PFRs and ideal separators in series. The feed to the first PFR contains the necessary quantity of unconverted reactants. For equilibrium-limited reactions each isothermal PFR contains just enough catalyst to achieve 99% equilibrium conversion at that temperature. For systems that are not chemical-equilibrium-limited ideal PFR conversion is based on considerations such as product selectivity. The partially reacted effluent from the first PFR is fed to an ideal separator that completely separates products from reactants, such that the feed to the next PFR is devoid of products. We repeat this procedure until the desired conversion or selectivity (yield) is attained for different temperatures within the permissible range. The optimum operating temperature enables us to meet the conversion or selectivity (yield) targets while utilizing the minimum amount of catalyst. For either exothermic or endothermic reactions we would like to operate at the highest possible temperature so as to maximize the reaction rates and minimize the catalyst mass. However, other factors such as chemical equilibrium (specifically for exothermic reactions), catalyst life, and column pressure (higher temperatures require higher operating pressures) impact the final operating temperature choice. Realistically, each reactive stage does not operate isothermally nor

does it completely separate products from reactants. Adiabatic operation and imperfect separation are the norm. Therefore, the actual catalyst mass is 20-30% greater than the minimum amount calculated using the ideal reactor-separator train. Number of Theoretical Stages. The number of theoretical (equilibrium) stages in each column section (rectification, reaction, and stripping) may be estimated once the reactive zone location and catalyst mass are known. Short-cut methods, such as the Fenske27Underwood28 method, give reasonable estimates for the nonreactive stripping and rectification sections. When the reactants are the light components, the reactive zone is located toward the top of the column and feed enters either below or at the bottom of the reactive zone. The stripping zone below the reactive section and the feed point separates unconverted reactants and inerts from the desired (heavy) product. Therefore, the stripping section of a reactive column performs a similar task to that of a nonreactive column because no reaction takes place below the feed point. A conservative estimate of the number of stripping and rectification stages is obtained as follows: (1) Assume that the column feed stream consists of a prereacted mixture of products and reactants at a desired conversion (equilibrium or some specified value) and temperature. (2) Specify desired top and bottom product purities and a reflux ratio 20% above the minimum. (3) Determine the theoretical number of stages and feed location by using the Fenske-Underwood method. (4) The theoretical stages below the feed represent the stripping stages. (5) Assume a conversion that is half the initially assumed value and repeat steps 1-3. The total number of stages should remain the same, since the top and bottom compositions are unchanged. However, the feed location changes and the number of rectification stages decrease because the feed is richer in reactants. The theoretical number of rectification stages obtained using this method represents a conservative estimate. The reactive zone is located near the bottom of the column when the reactants are the least volatile components and one or more products are more volatile. The section above the reactive zone that removes unconverted reactants from product(s) functions as a rectification section. Feeds are introduced above the reactive zone and no reaction occurs above the feed point. A procedure similar to that outlined in steps 1-5 can be used to obtain rectification- and stripping-stage estimates for this case: (1) Repeat steps 1-3 from above. (2) The number of rectification stages is the theoretical number of stages above the feed point. (3) Repeat step 5 from above. The number of stripping stages is the theoretical number of stages below the feed point and represents a conservative estimate. Reactive Zone Height and Column Diameter. For nonreactive columns the diameter is based solely on flooding velocity and pressure drop considerations. For reactive columns a key additional constraint is packing catalyst density (catalyst mass per unit column volume). Catalytic bale or structured packing densities range from 6 to 10 lb/ft3 (96-160 kg/m3). Reactive zone height depends on the catalyst mass, packing catalyst density, and column diameter. Pressure drop constraints limit the amount of catalyst that can be accommodated per

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Table 1. Procedure to Estimate the Reactive Zone Height, Reflux Ratio, and Column Diameter (1) Assume that the column feed stream consists of a prereacted mixture of products and reactants at a desired conversion. (2) Specify the distillate and bottom product compositions and determine minimum reflux requirements using the Underwood method. Use a reflux ratio 20% greater than the minimum reflux ratio as an initial estimate to determine column vapor and liquid velocities. (3) Estimate the diameter from the maximum allowable vapor velocity (80% of flood velocity). (4) Estimate the catalyst volume from the catalyst mass and maximum packing catalyst density.

Vcat )

mcat Fpacking

(5) Determine the reactive zone height from the calculated catalyst volume and estimated diameter.

hcat )

Vcat π d 2 4 col

(6) Determine the number of reactive stages by dividing the total reactive zone height by the catalytic packing HETP.

Ncat )

hcat HETPcat

(7) Simulate a reactive column with the calculated number of rectification, reaction, and stripping stages. (8) Increase the reflux ratio by 10%. If the conversion decreases, go to step 9. If the conversion increases, calculate the new maximum vapor velocity and column diameter and repeat steps 4-7. Keep increasing the reflux ratio in 10% increments and repeating steps 4-7 (if necessary) until there is no change in conversion. Go to step 10. (9) If the conversion decreases with increasing reflux, decrease the reflux ratio by 10%, calculate a new maximum vapor velocity and column diameter, and repeat steps 4-7. Keep decreasing the reflux ratio in 10% increments and repeating steps 4-7 (if necessary) until there is no change in conversion. Go to step 10. (10) If the desired conversion is attained, then we have reliable estimates for the reactive zone height, reflux ratio and column diameter. If the desired conversion is not attained, increase the catalyst mass and repeat the procedure starting with step 3.

unit column volume. When residence time limitations occur (inadequate catalyst mass), either the column diameter or reactive zone height has to be increased to accommodate the additional catalyst mass. Increasing the reactive zone height is not always an option because changes in the packing height affect column separation performance, thereby changing compositions in the reactive zone. For equilibrium-limited systems this could lead to temperatures and concentrations that favor the reverse reaction, thus causing product decomposition. Reflux Ratio. Reflux rate and reflux ratio impact both reaction and separation performance in reactive columns. For equilibrium-limited systems, a high reflux rate increases the separation of products from reactants, thereby increasing the reaction driving force. For example, for the MTBE system reactive column reflux ratios range from 1.5 to 2.0.29 A nonreactive column fed with a prereacted equilibrium mixture requires a reflux ratio of 0.6-0.8 for the same separation. The same holds true for the TAME system, although reflux ratios for this system are more dependent on the amount of excess methanol. For low-boiling reactants, an increase in reflux increases the reactant recycle across the reactive zone. This is particularly useful when azeotrope formation depletes the reactive zone of one or more reactants. For such cases we compensate for reactant depletion by increasing the reflux ratio, or by feeding excess reactant at an auxiliary feed location. The trade-off between excess reactant and increased reflux is dictated by economics. The reflux ratio also significantly impacts the reactive zone residence time for both equilibrium- and nonequilibrium-limited systems. Excessive reflux leads to operating problems and insufficient reaction holdup, thus causing incomplete conversion.5 For ether systems (MTBE and TAME) excess reflux causes incomplete conversion of iso-olefins (isobutene and isoamylene) and leads to product contamination. Thus, reactive columns, like extractive distillation columns, have a minimum and maximum reflux ratio. For kinetically-limited sys-

tems Doherty and co-workers15,16 include the effect of reflux ratio on residence time by defining a dimensionless Damkohler number (ratio of a characteristic liquid residence time to a characteristic reaction time). Conventional short-cut methods for determining minimum reflux cannot be easily extended to reactive columns because the reflux ratio impacts both reaction and separation. Boundary-value procedures14 and fixedpoint methods15,16 are more useful for determining minimum and actual reflux requirements. Reactive column diameters depend on the vapor-liquid traffic and packing catalyst density, and the reflux ratio directly impacts vapor and liquid velocities. Thus, we use an iterative procedure (Table 1) to estimate the reflux ratio, reactive zone height, and column diameter. Actual Trays/Packed Height. For nonreactive columns theoretical (equilibrium) stages are converted to either actual number of trays or packed height by using tray efficiencies or the height equivalent of a theoretical plate (HETP). We recommend the use of experimental tray efficiency or packing HETP data wherever possible for the stripping and rectification sections. If packing geometry is known, HETP may be estimated using reliable models or correlations. Superimposing reaction on separation makes the use of Murphree efficiencies and/or height of transfer units for the reactive zone extremely difficult. Experimental HETP data can be used to back-calculate the number of reactive stages from total reactive zone height. However, data for catalytic packing are rarely available and should be used with caution, particularly when the physical properties of the test system and the reaction mixture differ considerably. For such cases it may be beneficial to calculate mass-transfer coefficients from experimental data by using physical properties and operating conditions from the test system. We recommend that the rate-based (nonequilibrium) approach be used for simulation and design for reactive columns, particularly when thermophysical and transport property data and reliable column hydraulic models are available.

Ind. Eng. Chem. Res., Vol. 38, No. 10, 1999 3701 Table 2. Product and Byproduct Reactions in TAME Synthesis

Table 3. Kinetic Rate Expressions for TAME Synthesis

[

]

[

]

(a)

K1aMa1B K2aMa2B KT KT - k4 aT 1 -k2 aT 1 KM aT KM aT rM ) KT a + aM KM T

[

]

(b)

KT K1aMa1B k2 aT 1 KM aT 1 a2B r1B ) - k5a1B 1 KT K3 a1B a + aM KM T

]

(c)

K2aMa21B KT k4 aT 1 KM aT 1 a2B r2B ) + k5a1B 1 KT K3 a1B a + aM KM T

(d)

KT KT K1aMa1B K2aMa2B - k4 aT 1 -k2 aT 1 KM aT KM aT rT ) KT a + aM KM T

[

[

[

[

[

[

]

[

]

]

]

]

[

]

[

]

Table 4. Kinetic Parameters for a TAME Synthesis Reaction

Case Study: Manufacture of tert-Amyl Methyl Ether (TAME) System Characteristics. The commercial ether 2-methoxy-2-methylbutane (tert-amyl methyl ether or TAME) is manufactured by reacting isoamylenes (2methyl-1-butene and 2-methyl-2-butene) with methanol. Overall conversion of isoamylenes is restricted by equilibrium to 56% at 60 °C when stoichiometric quantities of isoamylene and methanol are used, but increases with increasing methanol-to-isoamylene ratios. There are three main reactions: two etherification reactions (one for each isoamylene) and isomerization between isoamylenes. Both etherification reactions are exothermic and equilibrium conversion decreases with increasing temperature. Isomerization equilibrium favors the formation of 2-methyl-2-butene, the dominant isoamylene present in the reaction mixture. From a kinetic standpoint, this is undesirable because fast-reacting 2-methyl-1-butene (2M1B) is replaced with slow-reacting 2-methyl-2-butene (2M2B). Possible side reactions include the formation of dimers and co-dimers of 2M1B and 2M2B, methanol condensation to dimethyl ether, and isoamylene hydration to tert-amyl alcohol (TAA) when water is present.30 High-temperature and low methanol content promote dimer formation while dimethyl ether is formed when methanol is present in excess. The hydration reaction is strongly equilibriumlimited and water must be present in large excess to form significant amounts of TAA. The key product and byproduct reactions are shown in Table 2. Thermodynamics and Reaction Kinetics. Kinetics for the three reactions are described by a LangmuirHinshelwood formalism.31 The mechanism which fits the data best assumes that alcohol and ether are strongly adsorbed on active sites, with the surface reaction between adsorbed alcohol and isoamylene in the bulk liquid being the rate-limiting step.31 The proposed mechanism, a modification of the Eley-Rideal mechanism commonly used for gas-phase reactions, results in the intrinsic rate expressions31 reproduced in

parameter

value

activation energy for reaction r2 (E2) pre-exponential factor for reaction r2 (k02) activation energy for reaction r4 (E4) Pre-exponential factor for reaction r4 (k04) rate constant for reaction 5 (k5) activation energy for KT/KM ratio (ER) pre-exponent for KT/KM ratio (k0R) k2/k4 at 354 K k4KT/KM at 354 K k5 at 354 K KT/KM at 354 K

100.0 kJ/mol 1.7054 × 1014 mol h-1 g-1 95.1 kJ/mol 1.3282 × 1014 mol h-1 g-1 3.1767 mol h-1 g-1 -4468.71 kJ/mol 0.02811 0.2437 0.1615 mol h-1 g-1 3.1767 mol h-1 g-1 0.1283

Table 5. Equilibrium Conversion as a Function of Temperature for the TAME Reactiona

a

temperature (°C)

conversion

50 60 70 80

0.618 0.560 0.491 0.446

Stoichiometric reactant ratio (methanol/isoamylene ) 1).

Table 3. The four parameters (k2, k4, k5, and KT/KM) are expressed as functions of temperature with activation energies and pre-exponential factors shown in Table 4. Also shown are the values calculated by Rihko and Krause31 at 354 K. They report standard deviations ranging from 6 to 39% for the etherification reaction parameters, and an activation energy of 39.1 kJ/mol with a standard deviation of 74.4 kJ/mol for the isomerization reaction. Since the uncertainty in estimating isomerization activation energy is considerable, we use the rate constant reported at 354 K. Rihko and Krause used activities instead of mole fractions and concentrations because of the nonideality of the reaction mixture. We calculated activity coefficients using the UNIQUAC model. Temperature-dependent equilibrium conversions calculated using experimental data and a fixed reactant ratio (methanol/isoamylene ) 1) are shown in Table 5. Rihko et al.32 used the same data to express equilibrium constants as a function of temperature for the three

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Table 6. C5 Feed Stream Composition36 component

mole fraction

mass fraction

boiling point (°C)

isopentane 1-pentene 2-methyl-1-butene n-pentane 2-pentene (cis) 2-methyl-2-butene

0.4821 0.0366 0.0823 0.0850 0.1555 0.1585

0.488 0.036 0.081 0.086 0.153 0.156

27.84 29.94 31.14 36.04 36.94 38.54

Table 7. Azeotropes Present in the TAME Reaction Mixture P ) 2.5 bar component component 1 2 methanol methanol methanol methanol methanol methanol methanol

2M1B 2M2B n-pentane isopentane 1-pentene 2-pentene TAME

x1

T (°C)

0.21 0.28 0.295 0.21 0.22 0.265 0.763

53.76 58.69 58.64 51.22 53.64 56.73 87.56

P ) 4 bar x1

T (°C)

0.243 69.24 0.31 73.78 0.328 73.96 0.252 66.61 0.267 69.06 0.301 72.29 0.793 102.45

P ) 5.5 bar x1

T (°C)

0.268 80.07 0.331 84.69 0.347 85.20 0.280 77.85 0.283 80.75 0.322 82.72 0.802 113.20

main reactions. They observed that the experimental equilibrium constant at a given temperature deviated from a constant value at very low methanol mole fractions (xMeOH < 0.01). They concluded that the uncertainty in activity coefficient calculations caused some discrepancies in equilibrium and rate calculations at these conditions. A change in the reaction mechanism may also occur at very low methanol mole fractions, resulting in a change in the reaction order with respect to methanol activity.33 Jakobsson et al.18 use the kinetic and equilibrium expressions provided by Rihko and coworkers31,32 with encouraging results. The mixed C5 feed stream consists of isopentane, n-pentane, 2-pentene (cis and trans), 1-pentene, and cyclopentene. It contains approximately 25% isoamylene and is a product from either a steam or fluid catalytic cracking unit. Typical feed stream compositions and atmospheric boiling points are shown in Table 6. All the C5 components form minimum-boiling azeotropes with methanol and contribute significantly to the nonideality

of the reaction mixture. The methanol content of the C5-methanol azeotropes increases with increasing pressure. Azeotropes and azeotropic compositions at three pressures (2.5, 4, and 5.5 bar) are shown in Table 7. Binary pairs involving methanol are highly nonideal and interaction parameters for these pairs were estimated from experimental data when possible. The TAME-C5 binary systems are mildly nonideal. Interaction parameters for these pairs were calculated from experimental data when available or estimated using the UNIFAC group contribution method. Binary pairs involving only C5 components (e.g., n-pentane-isopentane) are mainly ideal with activity coefficients varying from 1 to 1.4 for the entire composition range. Interaction parameters for these pairs were estimated using UNIFAC. Binary interaction parameters for the UNIQUAC Gibbs energy model and their sources are shown in Table 8. Basis and Assumptions. The reactor-column configuration produces 200000 metric ton/year (236.25 kmol/h at 8300 h/year) of TAME. We accounted for two etherification reactions (etherification of 2M1B and 2M2B) and isoamylene isomerization while neglecting side reactions. The absence of kinetic and thermodynamic data for the dimers and trimers makes it difficult to include these components in the model. While reaction kinetics are available for the TAA reaction,34 TAA formation depends strongly on the amount of water present in the reaction mixture. Since only trace water amounts are present in the feed, we did not include the TAA reaction in the model. Detailed design for this system would necessarily include the side reactions. We chose design specifications based on typical industrial conditions for a catalytic distillation unit with a fixedbed prereactor;35 they are shown in Table 9. The simulated process (Figure 2) consists of an isothermal plug-flow reactor operating at 70 °C and 4 bar and a packed reactive column. The mixed feed stream to the reactor contains 25% excess methanol and conversion approaches equilibrium (within 0.5%) when the catalyst mass exceeds 9550 kg. The equilibrium

Table 8. UNIQUAC Binary Parameters for TAME Simulation and Design component 1

component 2

b12 (K-1)

b21 (K-1)

source

methanol methanol methanol methanol methanol methanol methanol 2M1B 2M1B 2M1B 2M1B 2M1B 2M1B 2M2B 2M2B 2M2B 2M2B 2M2B TAME TAME TAME TAME n-pentane n-pentane n-pentane isopentane isopentane

2M1B 2M2B TAME n-pentane isopentane 1-pentene 2-pentene 2M2B TAME n-pentane isopentane 1-pentene 2-pentene TAME n-pentane isopentane 1-pentene 2-pentene n-pentane isopentane 1-pentene 2-pentene isopentane 1-pentene 2-pentene 1-pentene 2-pentene

54.40 28.82 75.15 -30.5 -9.34 -10.49 403.0 -6.85 -41.65 -6.81 82.98 62.57 64.84 -16.61 -3.89 63.71 52.74 63.15 -85.47 107.92 18.98 20.78 -110.03 59.39 28.04 -48.75 15.36

-788.85 -757.37 -511.92 -553.32 -670.42 -612.24 -716.53 3.80 20.71 -5.30 -98.81 -69.18 -71.61 1.042 -6.752 -81.44 -58.81 -69.79 -60.35 -144.72 -39.71 -41.85 98.07 -84.01 -33.93 40.45 -18.27

Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) UNIFAC estimate Dortmund Data Bank (ASPEN) UNIFAC estimate Dortmund Data Bank (ASPEN) UNIFAC estimate UNIFAC estimate UNIFAC estimate Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) Dortmund Data Bank (ASPEN) UNIFAC estimate UNIFAC estimate experimental data regression UNIFAC estimate UNIFAC estimate UNIFAC estimate UNIFAC estimate Dortmund Data Bank (ASPEN) UNIFAC estimate UNIFAC estimate UNIFAC estimate


Figure 2. Reactive distillation process for TAME production from methanol and a mixed C5 stream. Table 9. Design Specifications for the TAME Case Study description

value

minimum desired conversion maximum top product TAME impurity minimum bottom product TAME purity minimum TAME recovery (molar) maximum percent flood maximum catalyst packing density

92% 70 ppm (mass) 98.6 wt % 99.6% 80% 10 lb/ft3(160 kg/m3)

mixture is subsequently fed to a reactive column operated at 4 bar (condenser pressure). The reactive section is packed with catalytic structured packing (approximate surface area ) 160-180 m2/m3 and void fraction ) 0.65-0.7). The rectification and stripping sections are packed with sheet metal structured packing (approximate surface area ) 220-250 m2/m3 and a void fraction ) 0.9-0.93), hereafter referred to as type II structured packing. Examples of type II structured packing include Flexipac-2Y, Gempak-2A, and Mellapak-250Y. Temperatures in the reactive zone of the column range from 77 to 80 °C and the overall process yield (conversion of isoamylenes to TAME) is 92.2%. The bottom product from the reactive distillation column has a TAME mass fraction of 0.988. A process material balance is shown in Table 10 and design and operating parameters for the prereactor and reactive column are included in Table 11. Stages/segments are numbered from the top. Application of Design Guidelines. The reaction rates decrease to half their initial value when conversion exceeds 32% for any temperature (Figure 3). At 70 °C and a methanol-to-isoamylene ratio of 1.25, the reaction approaches equilibrium within 0.5% at 62% conversion. Isothermal operation at 70 °C ensures high reaction rates and minimizes intrapellet diffusion resistance. Although rates are approximately twice as large at 80˚C (methanol/isoamylene ) 1.25), equilibrium conversion is limited to 52%, and diffusion resistance increases considerably (effectiveness factor η ≈ 0.68) at this

temperature. We recommend the use of a prereactor (preferably isothermal) operating with exit conversions within the 55-65% range. We examined operating pressures between 2.5 and 5.5 bar and used three factors, reactive zone temperature, separation requirements (minimum reflux ratio and number of stages), and azeotropic methanol requirements, to select an optimum pressure. Table 12 shows separation, catalyst, and azeotropic methanol requirements at three pressures. We chose an operating pressure of 4 bar on the basis of the results shown in Table 12. Although the rate increases when the pressure is increased to 5.5 bar, the reduction in catalyst mass is not sufficient to offset the increase in methanol and separation requirements. Two important factors were considered for reactive zone location: 2M2B concentration and top product purity (less than 70 ppmw TAME). First, we simulated a nonreactive column with a feed at equilibrium (compositions and temperature corresponding to conditions at reactor exit). We found that the concentration of 2M2B (least volatile C5 component) reached its maximum value between equilibrium stages 12 and 23, while high methanol concentrations prevailed between equilibrium stages 2 and 13. Consequently, we located the reactive zone between stages 6 and 24 (nonequilibrium segments 9 and 30). The conversion increases when the reactive zone shifts toward the condenser (stages 3-21), although there is an increase in TAME top product impurity (4000 ppm). The optimum reactive zone location represents the best compromise between conversion and top product purity. Methanol concentrations in the lower half of the reactive zone were increased by injecting methanol on stage 24 (segment 30), the last stage/segment of the reactive zone. The amount of methanol required at 4 bar for both reaction and azeotrope formation is shown in Table 13. We simulated the column using different reactant ratios (methanol/isoamylene) and found that conversion decreased when the amount of methanol fed to the column was less than that shown in Table 13. Amounts exceeding the calculated amount did not significantly increase conversion, although additional methanol did decrease the amount of reflux required. When excess methanol is supplied to the reactive zone, only a few stages near the top of the reactive section have reactant (methanol/ isoamylene) ratios greater than unity; most stages are still richer in isoamylene. Substoichiometric reactant ratios favor dimer formation and increase intrapellet

Table 10. Material Balance for TAME Reactive Distillation Process Stream Identity catcrk

meoh

feed

prod

exmeoh

top

bot

phase temperature (°C) pressure (bar)

Quantity

liquid 79.20 4.1

liquid 104.79 4.1

liquid 70.27 4.1

liquid 70.0 4.0

liquid 104.79 4.1

liquid 68.75 4.0

liquid 138.66 4.05

methanol 2M1B 2M2B TAME n-pentane isopentane 1-pentene 2-pentene total flow (kmol/h) total flow (kg/h) total flow (L/min)

0.0 85.59 164.8 0.104 88.40 501.2 38.06 161.7 1040.0 74132 2195.4

313.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 313.0 10029 240.7

215.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 215.0 6889 165.3

297.81 2.95 16.50 0.047 87.68 500.05 37.97 160.57 1103.6 67241 1851.1

trace 0.03 0.78 230.24 0.72 1.23 0.09 1.15 234.24 23809 627.0

Flow Rates (kmol/h) 313.0 156.0 85.92 9.55 164.8 83.94 0.104 157.0 88.40 88.40 510.2 501.2 38.06 38.06 161.7 161.7 1353.0 1196.1 84161 84161 9935.0 2293.7

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Ind. Eng. Chem. Res., Vol. 38, No. 10, 1999 Table 12. Design and Operating Parameters as a Function of Pressure P ) 2.5 bar

description

P)4 bar

P ) 5.5 bar

minimum reflux ratio 0.286 0.329 0.364 minimum no. of stages 12 13 14 no. of equilibrium stages 26 27 29 reactive zone temperature (°C) 62-65 76-78 89-92 catalyst mass for equilibrium 30929.1 9543.8 4142.3 conversion (kg) reflux ratio 0.40 0.46 0.51 azeotropic methanol 245.37 302.15 341.18 requirements (kmol/h) Table 13. Methanol Requirements for Reaction and Azeotrope Formation purpose Figure 3. Formation rates for TAME synthesis as a function of isoamylene conversion at different temperatures and a reactant ratio (methanol/isoamylene ) 1.25). Table 11. Design and Operating Parameters for the TAME Reactive Distillation Scheme equilibrium-stage model

description no. of segments/stages (total) total packed height (m) no. of segments/stages (rectification) rectification section packed height (m) no. of segments/stages (reactive) reactive section packed height (m) no. of segments/stages (stripping) stripping section packed height (m) column diameter (m) reflux ratio bottoms-to-feed (B/F) ratio column feed segment(s)/stage(s) catalyst density (kg/m3) total catalyst mass (kg) catalyst mass in column (kg) percentage flood reboiler duty (MW) condenser duty (MW) total isoamylenes conversion (%) isoamylenes conversion in column (%) TAME mass fraction (bottom) TAME mass fraction (top) reactive zone average temperature (°C) reactor diameter (m) reactor length (m) catalyst mass in reactor (kg) reactor operating temperature (°C)

rate-based model

35

45

16.36 4

23.39 7

1.64 @ HETP ) 0.41 m

4.09

19

22

10.62 @ HETP ) 0.56 m

12.29

10

14

4.10 @ HETP ) 0.41 m

7.01

4.0 1.5 0.167

4.05 2.0 0.166

24 (methanol) and 29 156.59 30444 20900

30 (methanol) and 36 156.63 33743 24200

80.0 19.281 19.095 93.08

77.06 23.128 22.919 91.92

30.41

29.25

0.99

0.988

15 ppm 74.96

71 ppm 75.6

1.5 6.0 9543.8

1.5 6.0 9543.8

70

70

diffusion resistance. Thus, additional methanol is often required to minimize dimer formation and the actual methanol quantity is larger than that shown in Table 13. Two feed streams enter the column: the equilibrium mixture from the prereactor (mixed feed) enters on stage 29 (segment 35) and subsidiary methanol enters on stage 24 (segment 30). If the mixed feed stream is fed at a point higher than stage 29 (segment 35), it results

reaction (XIA ) 0.935) azeotrope with n-pentane azeotrope with isopentane azeotrope with 1-pentene azeotrope with 2-pentene azeotrope with unreacted 2M1B azeotrope with unreacted 2M2B azeotrope requirements (total) total methanol requirement (reaction + azeotrope formation) methanol supplied to prereactor unconverted methanol at exit of prereactor methanol added to reactive distillation column

amount (kmol/h) 234.15 43.15 168.88 13.87 69.64 1.79 4.82 302.15 536.30 313.04 165.29 223.38

in TAME decomposition on reactive stages/segments below the feed. Some separation capacity should exist between the mixed feed point and the reactive zone to recover TAME from the prereacted feed and prevent its decomposition. We selected an initial mixture feed point between the last reactive stage/segment and the reboiler, simulated column performance with this feed location, and improved initial estimates by matching column and feed TAME compositions. The final mixed feed location selected gave the highest conversion. Packing between the feed point and the last reactive segment removes much of the TAME from the reactants, while the packed section below the feed removes reactants from the bottom product. The stripping section below the feed point should be large enough to prevent 2M2B (the highest boiling C5 component) from contaminating the bottom product. Additional methanol fed directly to the reactive zone increases methanol concentrations in the reactive zone and improves conversion. The subsidiary feed location at the bottom of the reactive section ensures a steady supply of methanol to this zone because the minimumboiling C5/methanol azeotropes force methanol toward the condenser. The excess methanol feed location depends on several factors, including mixed feed composition. Jakobsson et al.18 feed excess methanol at the top of the reactive zone and maintain a high feed rate such that methanol concentration in the upper part of the column is at the azeotrope composition of the hydrocarbon-methanol mixture. We simulated the same configuration (methanol injection at the top of the reactive zone) and observed that methanol requirements increased by approximately 20% (for the same conversion level). We calculated the catalyst volume by simulating a series of plug-flow reactors and ideal separators (HETES concept). Three reactors containing 25450 kg (28.28 m3) of catalyst were required for 93% conversion. An equilibrium-stage simulation of a reactive column with a

Ind. Eng. Chem. Res., Vol. 38, No. 10, 1999 3705 Table 14. Estimated and Calculated HETP Values for the TAME System section rectification reaction stripping (above feed) stripping (below feed)

estimate iteration iteration final (m) # 1 (m) # 2 (m) value (m) 0.406 0.559 0.406 0.406

0.61 0.635 0.559 0.483

0.558 0.584 0.533 0.457

0.523 0.546 0.495 0.437

prereactor indicated a need for 30444 kg, or 20% more than the amount calculated using the HETES concept. A rate-based simulation of the same process required 31500-33700 kg of catalyst, an amount approximately 25-30% in excess of the ideal quantity (25450 kg). The results show that the catalyst volume can be estimated with reasonable accuracy using the HETES concept. A short-cut Fenske-Underwood calculation with 2M2B as the light key component and TAME as the heavy key gave 15 stripping stages and 13 rectification stages. A rigorous nonreactive column simulation confirmed that 15 stripping and 10 rectification stages were required. Available HETP estimates for type II structured packing vary from 0.36 to 0.46 m (14-18 in.). We used an initial HETP of 0.41 m (16 in.) for the rectification and stripping sections and a value of 0.56 m (22 in.) for the reactive zone. We used a hydraulic model developed by Subawalla et al.36 in conjunction with a rate-based simulation to update our initial estimates. The hydraulic model integrates the column capacity (pressure drop and percent flood) and mass-transfer efficiency (liquid holdup, mass-transfer coefficients, and interfacial area) calculations. Results from the iterative procedure are shown in Table 14. Comparison of the initial estimates with the final results shows that nonreactive, catalytic packing HETP values calculated with a different physical system are not entirely appropriate for catalytic distillation conditions. We used the Fenske-Underwood method to estimate initial reflux requirements for a nonreactive column fed with the same prereacted equilibrium mixture as a reactive column. The minimum reflux ratio predicted using this method was 0.33. We used a reflux ratio of 0.46 (20% greater than the minimum) for our initial estimate and determined a preliminary column diameter (2.5 m) at 80% flood (maximum permissible vapor velocity). We calculated the catalyst volume from the initial catalyst mass (15900 kg) and catalytic packing density (160 kg/m3) estimates and used it to estimate the reactive zone height (7.91 m). Since catalytic packing HETP values range from 0.46 to 0.61 m (18-24 in.), there is a corresponding variation in the number of reactive stages (13-18). We used initial estimates for rectification, reactive, and stripping stages to simulate a reactive column and found that the conversion achieved (80%) was much lower than our target conversion (92%). We followed the procedure shown in Table 1 and kept increasing the reflux ratio until there was a negligible increase in conversion at a reflux ratio of approximately 1.5 (for a fixed methanol feed rate). The catalyst mass in the column (4.05-m diameter) ranges from 19000-24160 kg and is accommodated using 19 reactive stages (24 reactive segments), and a reactive packed height of 12.29 m. The final estimates for the column diameter and height were obtained after 5 iterations of the design procedure. The estimates were verified using an equilibrium-stage simulation which indicated that 35 (11 stripping, 5 rectification, and 19 reactive) stages were

Figure 4. Reactant and product liquid composition profiles in an isothermal plug flow reactor preceding the reactive column. (Operating temperature ) 70 °C; inlet methanol-to-isoamylene ratio ) 1.25).

Figure 5. Reactant and product liquid composition profiles in a TAME reactive column with prereacted feed.

required. A part of the rectification load (the equivalent of 5 stages) was transferred to the reactive section. The maximum vapor velocity (77% of flood velocity) corresponding to this design occurs at the top of the reactive zone. Results and Discussion Reactor composition profiles are shown in Figure 4. Isomerization equilibrium favors 2-methyl-2-butene (2M2B) formation, thus, the increase in 2M2B formation and sharp decrease in 2-methyl-1-butene (2M1B) composition near the reactor entrance. Very little change in composition occurs for lengths greater than 4.5 m, which indicates that all three reactions approach equilibrium beyond this point. Figures 5-12 show column composition, temperature, flow, HETP, reaction, and mass-transfer rate profiles. Important observations and conclusions from these profiles are summarized below: (1) The methanol/C5 azeotropes behave as light key components and methanol is enriched in the rectification section, although it is not the lowest boiling component (Figure 5). Flat temperature and composition profiles in the reactive zone indicate the absence of significant vapor-liquid mass transfer. However, some rectification does occur and isopentane (the most volatile

3706


Figure 6. Liquid composition profiles of nonreactive C5 components in a TAME reactive column with prereacted feed.

Figure 8. Molar liquid and vapor flow profiles in a TAME reactive column with prereacted feed.

Figure 9. Column HETP profiles in a TAME reactive column with prereacted feed. Profiles for each section are shown separately. Figure 7. Liquid and vapor temperature profiles in a TAME reactive column with prereacted feed.

C5) is stripped from the remaining C5 components (Figure 6). The stripping section below the feed point is primarily responsible for separating TAME from the reactants and inerts. The rise in temperature reflects the increase in TAME (heavy key) composition (Figure 7). The stripping section between the bottom of the reactive zone and the feed segment, besides stripping TAME from the reactants and inerts, also separates 2M2B (the least volatile C5) from the intermediateboiling C5 components (n-pentane and 1-pentene). (2) Substoichiometric reactant ratios prevail over much of the lower half of the reactant zone. Since 2M2B is the highest boiler among the C5 components, it is present in large amounts near the feed point and the bottom half of the reactive zone, while methanol is forced upward on account of azeotrope formation (Figure 5). The addition of excess methanol does not guarantee stoichiometric reactant ratios (methanol/isoamylene); however, it does increase methanol concentrations and conversion within the reactive zone. (3) For a fixed diameter an increase in vapor-liquid traffic (flow rates) implies an increase in liquid and vapor superficial velocities with corresponding improvement in mass-transfer efficiency. Since the stripping zone below the feed has high liquid and vapor rates, average HETP values in this section are 19% lower than those in the rectification section (Figures 8 and 9).

Figure 10. Fraction of maximum allowable vapor velocity (left ordinate) and TAME formation rate (right ordinate) profiles in a TAME reactive column with prereacted feed.

Figure 10 shows percent flood and reaction rate. Percent flood is highest on segment 9 because of two factors, the reduced void fraction of catalytic packing and the fact that the maximum vapor flow rate occurs on this segment. Vapor flow capacity also decreases with increasing liquid flow rate because liquid rivulets occupy a significant fraction of the packing void space and restrict area available for upward vapor flow.


Figure 11. Reactant and product mass-transfer rate profiles in a TAME reactive column with prereacted feed.

(4) The reaction rate (Figure 10) is maximum near the top of the reactive zone (segment 9-10) where there is minimal TAME present. It declines rapidly as the TAME content increases and is relatively constant throughout the reactive zone (segments 13-25). Increasing the reactive segments beyond segment 30 leads to very low reaction rates (equilibrium conditions) or negative rates (TAME decomposition) on the lower reactive segments. Positive mass-transfer rates indicate mass transport from vapor to liquid (Figure 11). Therefore, in the absence of azeotropes a component which has negative rates throughout the column is the light key, and a component with positive rates is the heavy key. For the TAME reaction system the methanol-2M2B azeotrope (pseudo-component) and TAME are the light and heavy keys, respectively. (5) The total mass-transfer rate (sum of mass-transfer rates) is negligible for much of the reactive and rectification sections. Thus, the constant molar overflow (CMO) assumption is valid for these sections. However, the stripping section does not operate at constant molar overflow because flow rates and flow ratio (liquid/vapor) change significantly in this section. (6) Isopentane removal from the other C5 components is the dominant separation process in the reactive zone (Figure 12). The reactive section is enriched with 2M2B removal from lighter C5 components (upper the stripping zone). Negligible interphase transport takes place in the reactive zone. This observation agrees with the flat composition and temperature profiles seen earlier. Implications for Other Systems Three important differences exist between TAME and MTBE reactive systems. First, equilibrium limits isoamylene conversion to 56% (stoichiometric reactant ratio and 60 °C) for the TAME system. Isobutene conversion for the MTBE system for the same temperature and reactant ratio (methanol/isobutene ) 1) is 91%. Second, at pressures suitable for reaction and distillation, the methanol content of the C5-methanol azeotropes greatly exceeds that of the corresponding C4-methanol azeotropes. Equilibrium conversion and azeotrope compositions for the MTBE system are shown in Tables 15 and 16. Third, reaction rates for the MTBE system are higher than those for the TAME system at

Figure 12. Mass-transfer rates of nonreactive C5 components in a TAME reactive column with prereacted feed. Table 15. Equilibrium Conversionas a Function of Temperature for the MTBE Reactiona

a

temperature (°C)

equilibrium conversion

60 70 80 90

0.945 0.905 0.88 0.85

Stoichiometric reactant ratio (methanol/isobutene ) 1).

Table 16. Azeotropes Present in MTBE Reaction Mixture P ) 6 bar component 1 methanol methanol methanol methanol methanol methanol

component 2 isobutene isobutane n-butane 1-butene trans-2-butene MTBE

x1

T (°C)

0.05 49.0 0.03 44.5 0.08 56.2 0.07 49.0 0.12 58.4 0.51 108.4

P ) 8 bar x1

T (°C)

0.067 60.3 0.05 56.0 0.11 67.5 0.09 60.0 0.14 69.2 0.54 119.5

P ) 10 bar x1

T (°C)

0.08 69.4 0.06 65.3 0.13 76.6 0.11 68.9 0.15 78.0 0.56 128.4

equivalent temperatures and conversions. Therefore, the MTBE system requires less methanol (typically 5-15% excess) and less catalyst. The high rates and favorable equilibrium may allow the elimination of the prereactor for the MTBE system, while this is rarely possible for the TAME system. Two hydration reactions that share common characteristics are isoamylene hydration to tert-amyl alcohol (TAA) and methyl acetate hydrolysis to acetic acid and methanol. In both cases the limiting reactants (isoamylene and methyl acetate) are the most volatile components and their volatility is substantially different from that of water. Therefore, maintaining high concentrations of both reactants in the reactive zone requires some ingenuity. Both systems also have individual characteristics that make reaction and separation difficult. The TAA reaction system requires a solvent (e.g., acetone) to prevent two liquid-phase formations while the methyl acetate system has several azeotropes (e.g., methanol-methyl acetate, methyl acetate-water, and acetic acid-water). We used the design guidelines to determine the reactive zone location and reactant ratio for these two systems. The reactive zone is located at the top of the column because the limiting reactant is the most volatile component. To minimize reactant loss, the column should be operated at or near total reflux (minimize distillate product). The lighter reactant (methyl acetate or isoamylene) should be fed at the

3708


bottom of the reactive zone and water should be fed at the top. Although these results are qualitative, they agree well with the experimental conditions provided by Gonzale´z et al.21 (isoamylene hydration) and Fuchigami37 (methyl acetate hydrolysis). Summary and Conclusions In this paper we discuss design guidelines for solidcatalyzed reactive distillation systems. Iterative procedures and methods to estimate reactive column catalyst mass, reactive zone height, theoretical stages, reflux ratio, packed height, and column diameter are presented. Superimposing reaction on separation leads to unique design trade-offs; thus, the column diameter depends both on the maximum vapor velocity and the packing catalyst density, and the reactant ratios depend on conversion requirements, azeotrope formation, and reactant and product volatilities. The operating pressure affects chemical equilibrium, reaction rate (reactive zone temperature), and relative volatility, and the reflux ratio impacts both separation and conversion. For equilibrium-limited systems, either excess reactant or increased reflux can be used to increase conversion. None of these trade-offs have to be considered in nonreactive column design; thus, reactive column design is inherently more complicated. The guidelines and procedures outlined in this paper provide good initial estimates and serve as a starting point for detailed design. Key conclusions from our modeling and experimental studies are summarized below: (1) For many systems the reaction rates decline abruptly only when compositions and temperatures approach equilibrium. We recommend the use of a prereactor upstream of a reactive column for such systems. (2) For most systems the operating pressure is driven by temperatures and reaction rates such that we maximize rates and minimize residence times and catalyst mass. However, there is an upper pressure limit beyond which rates actually decrease because of reactant depletion due to relative volatility and/or azeotropic effects, and a decrease in departure from chemical equilibrium. (3) Reactive columns require one reactant in large excess when large relative volatility differences exist between reactants. A subsidiary feed location is often required. (4) Stringent purity specifications on either the distillate or bottoms product can alter the reactive zone location when either the height of the stripping or rectification section is increased. (5) The feed location is adjusted to minimize product decomposition when a prereacted product-containing stream is fed to a reactive column. Such a stream should enter the column at some distance from the reactive zone, thereby ensuring that separation between products and reactants occurs on stages between the feed point and the reactive zone. (6) Selecting an appropriate amount of catalyst is an important aspect of reactive column design. Inadequate catalyst volume reduces residence time and gives poor conversion. For equilibrium-limited systems, excess catalyst, particularly when it is inappropriately located, may result in product decomposition. A good initial estimate for catalyst mass is obtained by simulating a sequence of ideal plug-flow reactors and separators in series.

(7) Short-cut methods (e.g., Fenske-Underwood method) give reasonable estimates for the number of nonreactive stripping and rectification stages. (8) Since the packing catalyst density limits the amount of catalyst per unit volume, when residence time limitations occur (inadequate catalyst mass), either the column diameter or reactive zone height has to be increased to accommodate the additional catalyst mass. (9) Since the reflux ratio impacts both reaction and separation efficiency, conventional short-cut methods for determining minimum reflux cannot be easily extended to reactive columns. (10) We recommend that the rate-based (nonequilibrium) approach be used for simulation and design for reactive columns, particularly when thermophysical and transport property data and reliable column hydraulic models are available. Acknowledgment We thank Aspen Technology Inc. for supplying the simulation tools used in this work and M. Sivasubramanium for his help with RATEFRAC. This work was funded by the Phillips Petroleum Foundation and the Separations Research Program. Nomenclature ai ) activity of species i dcol ) column diameter (m) hcat ) height of catalytic packing bed (m) HETPcat ) height equivalent of a theoretical plate for catalytic section (m) k2 ) reaction rate constant for etherification of 2M1B k4 ) reaction rate constant for etherification of 2M2B k5 ) reaction rate constant for isomerization of 2M1B KT/KM ) ratio of adsorption equilibrium constants mcat ) catalyst mass (kg) Ncat ) number of reactive or catalytic stages r ) reaction rate (gmol h-1 g-1) T ) temperature (K) Vcat ) volume of catalyst (m3) Greek Symbols η ) catalyst effectiveness factor Fpacking ) catalytic packing density (kg of catalyst/ m3 of column) Subscripts T ) TAME M ) Methanol 1B ) 2-methyl-1-butene 2B ) 2-methyl-2-butene

Literature Cited (1) Smith, L. A. Catalyst System for Separating Isobutene from C4 Streams. U.S. Patent No. 4,215, 011, 1980. (2) DeGarmo, J. L.; Parulekar, V. N.; Pinjala, V. Consider Reactive Distillation. Chem. Eng. Prog. 1992, 88 (3), 43. (3) Pyha¨lahti, A.; Klemola, K. Bench Scale Experiments With TAME for the Comparison of Two Catalytic Distillation Packing Arrangements. Acta Polytech. Scand. 1998, 257, 101. (4) Sundmacher, K.; Hoffmann, U. Activity Evaluation of a Catalytic Distillation Packing for MTBE Production. Chem. Eng. Technol. 1993, 16, 279. (5) Agreda, V. H.; Partin, L.; Heise, W. High-Purity Methyl Acetate Via Reactive Distillation. Chem. Eng. Prog. 1990, 86 (2), 40. (6) Gildert, G.; Groten, W.; Putman, H. Hydrogenation by Catalytic Distillation. Presented at the AIChE Spring National Meeting, Houston, TX, March 1995.

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Received for review January 5, 1999 Revised manuscript received June 28, 1999 Accepted July 5, 1999 IE990008L

Design Guidelines for Solid-Catalyzed Reactive Distillation Systems

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