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Process intensification using multifunctional reactors
Chemical Engineering Science 56 (2001) 251}267
Process intensi"cation using multifunctional reactors F. M. Dautzenberg*, M. Mukherjee ABB Lummus Global Inc., Technology Development Center, 1515 Broad Street, Bloomxeld, NJ 07003-3096, USA
Abstract Process Intensi"cation is the strategy of making signi"cant reductions in the size of a chemical plant in order to achieve a given production objective. Innovations in catalytic reactors, which constitute the heart of such process technologies, are often the preferred starting point. Over the last two decades, innovative multifunctional reactor systems have been developed to intensify chemical processes by synergistically combining chemical reaction with momentum, heat and mass transport in a single vessel. This paper aims to emphasize an interdisciplinary approach to multifunctional reactor design. Examples from commercially practiced technologies are used to illustrate how process intensi"cation can be achieved through multifunctional reactors. Furthermore, new developments in this area will be highlighted. 2001 Elsevier Science Ltd. All rights reserved.
1. Introduction Process intensi"cation (PI) is a term used to describe the strategy of making dramatic reductions in the physical size of a chemical plant while achieving a given production objective. The concept was pioneered by ICI during the late 1970s when the primary goal was to reduce the capital cost of a production system. Currently, increased pressures to remain cost competitive in an environmentally responsible manner have led to a renewed interest in PI. PI is being applied throughout the chemical process industry (CPI) to reduce investment and operating costs of chemical plants to increase pro"tability and mitigate greenhouse gas emissions. Multifunctional reactive systems (catalysts, reactors, etc.) allow a unique way of achieving PI in the chemical process and re"ning industry. By its very de"nition, several `functionsa or processes are designed to occur simultaneously in multifunctional reactors. One aims at on optimum integration of mass, heat and momentum transfer within a single reactor vessel. Multifunctional reactors are not new to the chemical industry. They have been in use in the industry for a very long time. Traditional FCC units are a good example,
* Corresponding author. Tel.: #1-973-893-3319; fax: #1-973-8932745. E-mail address: [email protected] (F. M. Dautzenberg).
incorporating several `functionsa in one catalytic reactor such as heat management, catalyst transport and two chemical reactions (cracking and coke removal by oxidation). Only recently have these reactors been formally classi"ed as being multifunctional and the large bene"ts derived from PI have been acknowledged. Consequences of using multifunctional reactors are reduced investment costs and signi"cant energy recovery or savings. Furthermore, improved product selectivity leads to a reduction in raw material consumption and, hence, operating costs. To achieve optimal performance with multifunctional reactors, it is important to understand where the integration of functionalities occurs. Referring to Fig. 1 (Dautzenberg, 1999), the following classi"cations can be made. E Type A multifunctionality operates at the catalyst level: may be introduced by combining catalytic properties with an engineered catalyst structure. E Type B multifunctionality at the reaction inter-phase: by integrating chemical reaction with enhanced interphase transport properties. E Type C multifunctionality at the intra-reactor level: by marrying chemical reaction with intra-reactor unit processes such as heat transfer or separations. E Type D multifunctionality at inter-reactor level: by combining two reactor operations using solids recirculation. This paper gives an overview of multifunctional reactors. It stresses the strategy employed in achieving
0009-2509/01/$ - see front matter 2001 Elsevier Science Ltd. All rights reserved. PII: S 0 0 0 9 - 2 5 0 9 ( 0 0 ) 0 0 2 2 8 - 1
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Fig. 1. Catalyst particle in a reacting medium.
the multifunctional character of these reactors, explaining the type of multifunctionality achieved and the most important parameters involved in this task. Recommendations for future designs will also be indicated.
the catalyst to perform functions that could otherwise only be done by using very expensive process equipment (see Table 1). 2.1. Bifunctional catalysts
2. Type A multifunctionality: chemical reaction and catalyst particle design The performance of a majority of chemical reactors (and hence the process) is signi"cantly in#uenced by the performance of the catalyst. A sizeable amount of research dollars is thus spent in upgrading performance of industrial catalysts. Catalyst research has been devoted to increase the catalyst activity and selectivity to improve process economics and reduce greenhouse gas emissions through better feedstock utilization. Researchers have discovered the impressive bene"ts that can be derived from the intensifying processes by using multifunctional catalysts. This can be achieved by combining a catalyst function with: E E E E E
Secondary catalytic function. Enhanced transport characteristics. Novel catalyst morphologies. Energy transfer properties. Reactant transfer properties.
The synergy through combining two processes in a single catalyst particle can lead to signi"cant process advantages. In some cases, this multifunctional behavior allows
Catalytic reforming is an important process in the re"ning industry in which the octane number of a hydrocarbon fraction is increased without signi"cantly changing its molecular weight. It is the primary source of aromatic chemicals (BTX). The main reactions that occur are dehydrogenation followed by isomerization and dehydrocyclization. For these reactions to occur, two types of sites are needed: metallic sites for dehydrogenation or cyclization, and acid sites for isomerization. Bifunctional catalysts (Antos, Aitani & Parera, 1995) consisting of metallic sites, e.g. Pt, and acid sites (e.g. alumina) are typically used for these applications. The synergistic bene"ts of using such dual functional catalysts become clear when one compares the performance of a physical mixture of two monofunctional catalysts. The physical mixture of Pt on carbon catalyst and acidic alumina has a much lower activity and deactivates rapidly. 2.2. Attrition resistant catalysts In circulating #uid-bed reactors (CFB), rapid transport of catalyst particles between vessels subjects the particles to strong attrition forces. Attrition resistance then becomes an important criterion. DuPont's recently
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Table 1 Examples of type A multifunctionality Application
Function 1
Function 2
Bene"ts/Impact
Catalytic reforming
Catalyse dehydrogenation
Catalyse molecular structure rearrangements
High space time yields
Butane oxidation
Catalyse oxidation
Enhance attrition resistance
Temperature control and selectivity enhancement
Hydroformulation of ole"ns
Catalyse reaction
Encapsulate homogeneous catalyst
Easy product separation from catalyst
Re"nery alkylation
Catalyse ole"n alkylation of para"ns
Enhance di!usion and stability
Better product quality and longer catalyst life
Fluid catalytic cracking
Catalyse cracking reaction
Provide energy transfer
Energy savings
Fig. 2. Attrition resistant VPO catalyst for butane.
commercialized tetrahydrofuran (THF) process shows a unique way of overcoming this challenge by endowing the catalyst with multifunctional behavior. The vanadium phosphate (VPO) catalyst traditionally used for this process is too weak to withstand forces associated with cycling between the #uidized bed regenerator and the dense phase riser reactor, where the gas velocities are in the &1 m/s range. Attrition resistance is imparted to the catalyst by spray drying this material with a silica hydrogel (Haggin, 1995) under conditions that allow the silica to migrate to the outer regions, thus encapsulating the active VPO catalyst into a porous silica shell (Fig. 2)
whose pore openings allow product and reactant molecules to di!use in and out of the active catalyst zone. 2.3. Encapsulated homogeneous catalysts A technique of `heterogenizinga homogeneous catalyst involves encapsulating the catalyst by a porous shell. This technique has been applied to a hydroformylation reaction where a Rhodium-based homogeneous catalyst is synthesized inside a hollow silica micro-sphere with pores (Davis, 1994) of the order of 1 nm. The metal complex is too bulky to di!use through the pores of the
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Fig. 3. Encapsulated homogeneous catalysts.
micro-sphere (Fig. 3). Deliberate introduction of di!usional limitations for CO to the liquid has potential bene"cial e!ects for product selectivity. Recently, another approach has been proposed (Mortensen, 1998). In this case (Fig. 3, right-hand side), surfactants are used to create micellar structures of nanodimensions that can be designed to contain homogeneous catalysts. Among others, the oxidation of propylene to propylene oxide with hydrogen peroxide is being developed at the Technical University of Eindhoven, Netherlands.
Fig. 4. Thin "lm catalyst.
2.4. Thin xlm catalyst A thin layer of active catalyst is deposited precisely on a solid support such that the catalytic material does not penetrate the support. The small thickness of the catalyst layer leads to a low di!usion barrier (Fig. 4). This concept has been applied to the lummus solid acid gasoline alkylation technology (Murrell, Overbeek, & Khonsari, 1999). By depressing the formation of heavy by-products through a highly e!ective catalyst layer, the catalyst life and product selectivity is greatly enhanced. 2.5. Energy transfer FCC catalysts are a prime example of multifunctional reactive systems. The active 1 lm zeolite crystals are embedded in a dense matrix of 50}100 lm silica}alumina particles. This allows the zeolite to survive the severe mechanical forces while being recirculated from the riser-tube reactor to the regenerator and back. Catalytic cracking is an endothermic reaction that varies from 5503C at the inlet of the reactor to 4753C at the outlet. Heat for the cracking reaction is supplied partly by the catalyst particle itself * the catalyst particle picks up
heat in the regenerator through burning of the coke on its surface &0.8 wt% at the inlet and less than 0.1 wt% at the outlet of the regenerator * and releases it in the riser reactor. 2.6. Dry zeolite catalyst Mass transfer limitation of reactants and products in zeolite catalyst extrudates was observed (Ercan, Dautzenberg, Yeh & Barner, 1998) for the alkylation of benzene to ethylbenzene and cumene. In laboratory tests, signi"cant performance enhancement was observed with decreasing catalyst particle sizes. Due to process limitations (pressure drop increase) and catalyst manufacturing limitations, catalyst particle size reduction cannot be implemented easily in industry. ABB Lummus Global Inc. developed a zeolite synthesis procedure (Murrell, Overbeek, Chang, van der Puil & Yeh, 1999) that provides ready-to-use zeolite catalyst particles with ultra small crystals and a high meso- and macropore volume. These catalyst features reduced mass transfer limitations. Furthermore, the preparation method is faster, uses raw materials more e$ciently than the conventional zeolite
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Fig. 5. `Dry zeolitea catalyst preparation.
synthesis technique methods and eliminates the "ltration step to collect the product. The new method employs mesoporous silica}alumina precursor particles that are converted to catalyst particles with the same morphology as the precursor particles (Fig. 5). This is achieved by e!ecting zeolite crystallization in the absence of an external liquid phase. The total pore volume that is formed by the micro-, meso- and macropores of the starting particle is maintained because the particle shape and size of the starting material are maintained upon conversion. Compared to conventional zeolite catalysts, the volume of meso- and macropores in catalyst particles from the novel synthesis procedure is enhanced. Moreover, the zeolite crystals obtained by this method are a factor of 3}10 smaller than those from commercial synthesis methods. The zeolite catalyst particles from the novel procedure showed 2}3 times higher alkylation activity
(per unit weight) than commercial zeolite catalyst particles.
3. Type B multifunctionality: chemical reaction and inter-phase transport A general classi"cation of Type B multifunctionality is provided in Fig. 6. 3.1. Inter-phase mass transfer Inter-phase mass transfer is the rate-limiting step for many hydrocarbon oxidations in the gas and liquid phase. In these instances, the overall reaction rate is not a function of catalyst activity but depends on the #uid Reynolds number. For commercial scale reactors,
Fig. 6. Classi"cation of Type B multifunctionality.
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Fig. 8. Buss' jet loop reactor.
Fig. 7. Liquid oxidation reactor.
inter-phase resistance can become extremely large requiring long residence times to achieve economic conversions. In these cases, the incorporation of a mass-transfer device in the reactor can lead to signi"cant process advantages. 3.1.1. Liquid oxidation reactor As stated earlier, many hydrocarbon oxidations are run in the mass-transfer-limited regime. The concentration of oxygen in the dispersed bubbles has a direct e!ect on the overall rate. To overcome this limitation, Praxair Inc. has developed a unique liquid oxidation reactor (LOR) (Fig. 7), a gas}liquid reactor that uses a downward pumping impeller contained within a draft tube to disperse and circulate the reactive gas. By e!ectively mixing the hydrocarbon with oxygen through recirculation of oxygen bubbles, the LOR (Roby, 1996) achieves better gas utilization, higher reaction rates and increased energy e$ciency compared with traditional CSTRs. The impeller draws the bubbles downward and redisperses them. The #ow is further directed by two sets of vertical ba%es, one above the inlet of the draft tube and one at the discharge end. 3.1.2. Jet loop (buss) reactors A loop reactor consists of a pipe through which process #uids are recirculated. Often it contains a multitubular heat exchanger and in-line disperser (e.g. jet or ori"ce). By incorporating a jet, high-energy dissipation rates can be conveniently supplied which are 10}100 times higher than those available in stirred tanks. Such dispersion intensities are useful to minimize the size and cost of the reactor, and to maximize yields when the key reaction rates are intrinsically rapid but controlled by
inter-phase mass-transfer rates. The Buss reactor (Fig. 8) is based on this principle and if often used for liquid hydrogenations (Moeller & O'Connor, 1996). It has the advantage of rapid gas-liquid mass transfer in the initial zone, combined with high heat removal capability in the tubular heat exchanger. 3.2. Inter-phase energy transfer A large percentage of all useful intermediate petrochemicals are products of selective oxidations of hydrocarbons. These reactions, however, are highly exothermic, generating large amounts of heat. Multitubular reactors have been used as a way of controlling the reaction exotherm by conducting the heat away from the packed bed to the shell-side heat-transfer #uid. The radial temperature pro"le for an exothermic reaction in a packed tubular-bed reactor is shown in Fig. 9. There is a steep gradient near the inside wall and a nearly parabolic temperature pro"le over the rest of the catalyst bed. This phenomenon is caused as a result of a high voidage near the tube wall. 3.2.1. Sintered metal reactor The sintered metal reactor overcomes the steep temperature gradient at the wall in tubular reactors by sintering metal particles to the tube wall. By doing so, the wall heat transfer resistance (by inter-phase forced convection) is largely eliminated. In addition, the heat transfer in the packed bed is mainly by conduction through a continuous metal matrix (Mulder, 1996). Since thermal conductivity of metals is an order-of-magnitude greater than conduction through loosely packed catalyst particles, the temperature control in a sintered metal reactor is optimal. This interesting new reactor concept still awaits commercial demonstration. 3.3. Inter-phase momentum transfer Design and performance of commercial-scale chemical reactors is often dictated by the inter-phase transport
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Fig. 9. Porosity, velocity and temperature in tubular reactor.
rates between the reacting surface and the bulk #uid. In several processes, these rates can be much lower than the chemical reaction rate leading to low reactor e$ciencies. For instance, inter-phase momentum transport (characterized by the parameter `f a friction factor) in packed bed reactors gives rise to pressure drop. High system pressure drops leads to higher compression costs and, in equilibrium limited reactions, a lower overall conversion. PI can then be achieved by incorporating a catalytic function integrated into a system with a low `f a using novel catalyst/morphologies and reactor con"gurations. 3.3.1. Monolithic reactors Flow through the interstices in packed beds is tortuous and disordered. The frictional resistance comes from two sources: E drag friction as a result of viscous e!ects ( fakl and surface-to-volume ratio) and E from friction as a result of inertial forces ( fakl and geometry of the surface). For smooth tubes, however, the only contribution of the friction factor is the drag friction. The inter-phase momentum resistance, as measured by the pressure drop for equivalent Reynolds numbers, is thus much higher in packed beds than in smooth tubes. This bene"t is exploited in monolithic reactors. Monolith reactors consist of straight channels in parallel (Fig. 12) with small diameters. The resulting laminar #ow (to Reynolds numbers of 500) does not show the kinetic energy losses that occur
in "xed-beds due to inertial forces at comparable super"cial velocities. The ability to o!er high surface-tovolume ratios at low-pressure drop makes it the reactor of choice for catalytic combustion processes. Almost all automobile catalytic converters today use monolithic reactors to control exhaust emissions (Bayegan, 1999).
4. Type C multifunctionality: chemical reaction and intra-reactor processes The majority of multifunctional reactors operating in the process and re"ning industry fall under this category, a combination of chemical reaction coupled with an intra-reactor process operating at the macro-scale. The examples stretch from well-established multi-tubular reactors for reaction temperature control to the more recent work investigating reactions using in situ supercritical #uids as extractants. In most cases, the integration of the unit process with the chemical reaction in the same vessel allows bene"ts in process operability that are impossible to achieve merely by externally coupling the two operations. Typical unit operations that are used in multifunctional reactors can be subdivided into three categories: E Chemical reaction with intra-reactor energy transfer. E Chemical reaction with intra-reactor momentum transfer. E Chemical reaction with intra-reactor mass transfer. Table 2 provides examples of Type C multifunctionality.
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Table 2 Examples of type C multifunctionality Intra-reactor transfer
Process
Application
Impact/bene"t
Energy transfer
Heat transfer reactor Energy storage
Multi-tubular control; higher yields Cato"n, VOC
Temperature
Filtration
Catalytic "lter for No removal V Radial #ow rector
Reduced investment
Distillation
Catalytic distillation
Stripping
Catalytic stripping reactor
Reduced investment Higher yields Better temperature control Product purity Smaller reactor
Momentum transfer
Low pressure drop Mass transfer
4.1. Chemical reaction with intra-reactor energy transfer Chemical reactions tend to either generate heat (exothermic reaction) or absorb heat (endothermic reaction). Oxidation, hydrogenation and alkylation are typical examples of exothermic reactions while catalytic dehydrogenation and catalytic cracking reactions are endothermic reactions Temperature control of reactors is important for several reasons: 1. For endothermic processes, the reaction absorbs heat and cools the reactants to e!ectively quench the product reaction. 2. Exothermic reactions may run away at high conversions if the heat is not dissipated. 3. Equilibrium-limited reactions are temperature-dependent * higher outlet temperatures typically increase equilibrium conversions. 4. Temperature control reduces loss of selectivity by reducing thermal degradation of products and feedstock. 5. Better catalyst utilization is achieved through temperature pro"ling in the reactor. 4.1.1. Reaction in a heat exchange device Coupling of a reactive system with heat transfer equipment has been practiced in the industry for a long time. Tubular reactors have been used extensively in the process industry to control temperatures for both exothermic and endothermic reactions. The catalyst pellets reside in tubes in a shell and tube heat exchanger. A heat transfer #uid is then circulated on the shell side; highpressure steam is often used as a heat transfer medium. Catalyst tubes may number from 500 to 10,000 per reactor. An interesting new development has been introduced by Linde: the tube bundle is composed of counter-wound
Reduced energy costs
Higher yields
spirals in which upwardly #owing water is evaporated (Boelt, Walser, Merz, Stabel, & Wunsch, 1994). The tubes run into a vertical vapor drum located at the reactor head. Advantages in construction and in heat transfer from the reaction gas to the tubes of the bundle are claimed. 4.1.2. Coupling of exothermic and endothermic reactions Dehydrogenation reactions are endothermic in nature but release a mole of hydrogen per mole of ole"n formed. A unique technology developed by Lummus/UOP called SMART couples the endothermic dehydrogenation reaction with an exothermic hydrogen combustion step to balance the energy requirements in the styrene reactor. Not only does this create a better temperature pro"le in the catalyst bed by removing hydrogen in the reactor, it reduces the reverse reaction leading to high EB conversions. 4.1.3. Energy storage The catalyst bed itself is sometimes used to store heat as in the case of C and C dehydrogenation reactions practiced in Lummus' Cato"n technology (Feldman, Dufallo, Tucci, & Balogh, 1992) (Fig. 10). The Cato"n approach utilizes "xed-bed reactors operated under vacuum conditions to achieve favorable equilibrium. The heat required for the endothermic reaction is supplied partly by the sensible heat of the feed and by the thermal energy contained in the catalyst bed itself. The "xed beds are cycled between reaction and reheat, which maintains the energy balance in the reactor. 4.1.4. Transient reactor operation Reactors can be operated advantageously with moving thermal fronts that are created by periodic #ow reversal (Matros, 1985). Low-level contaminants or waste
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Fig. 10. Cato"n reactor system.
Fig. 11. Reverse #ow reactor concept.
products (VOCs) can be e$ciently removed in adiabatic "xed beds with periodic reversal by taking advantage of higher outlet temperatures generated in earlier cycles to accelerate exothermic reactions (Fig. 11). Energy and cost savings are e!ected by this substitution of internal heat transfer for external exchange. 4.2. Chemical reaction with intra-reactor momentum transfer Some chemical processes require high gas #ow rates through reactors, as in the case of ammonia synthesis or production of styrene from ethylbenzene. In the catalytic treatment of #ue gases to reduce atmospheric pollutants, e.g., sulfur and nitrogen oxides, very large volumes of air at relatively low pressures must be handled. In all these situations, pressure drop through traditional "xed-bed reactors will become excessively large, which calls for reactor design with very low-pressure-drop.
4.2.1. Parallel passage reactor Particles of a catalyst or a regenerable adsorbent are enclosed in wire screen envelopes, which are mounted in a parallel fashion. Gas #ows in the empty passages between envelopes. The straight, unobstructed gas channels give rise to a low-pressure-drop and good tolerance (Fig. 12). Pollutant molecules are transported to the catalyst or adsorbent by lateral di!usion. This type of reactor has been applied in the shell #ue gas desulfurization process (Dautzenberg, Naber, & van Ginneken, 1971), which is based on the use of a regenerable copper adsorbent. 4.2.2. Radial yow reactor High temperatures and low operating pressures favor dehydrogenation of ethylbenzene to styrene. The Lummus-UOP Classic Styrene process (Dautzenberg & Mukherjee, 1998) uses two radial #ow reactors with an inter-stage heater to get around the constraints of the
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Fig. 12. Low-pressure-drop reactor systems.
dehydrogenation reaction. By increasing the #ow surface area, the mass-#ux of reactants through the bed diminishes signi"cantly resulting in low bed pressure drop. An optimally designed de#ector housed in the annular space of the adiabatic reactor ensures even #ow distribution from the very bottom of the bed to the very top. 4.2.3. Composite structured packing reactor The composite structured packing (CSP) reactor (Strangio, Dautzenberg, Calis, & Gupta, 1999) achieves low-pressure-drop through the regular vertical stacking of catalyst particles inside parallel channels. The regular stacking, shown schematically in Fig. 13, results in a minimization of the momentum loss which normally occurs in random packed beds due to #ow directional changes. The CSP can be considered a commercially viable extension of the `bead-stringa reactor concept proposed by co-workers at the Technical University of Delft in Delft, Netherlands. The low-pressure-drop/length of the CSP o!ers considerable advantage to the radial #ow reactors in that it allows the use of several catalyst beds with inter-stage reheat within one reactor shell, resulting in a higher overall average operating temperature and a higher styrene yield. 4.3. Chemical reaction with intra-reactor mass transfer Chemical reaction with mass transfer at the intrareactor level is mainly concerned with in situ separation of products from the reaction zone. The main reasons for considering in situ product removal are
E To enhance conversion in equilibrium limited reactions by shifting the equilibrium to the right. E To prevent further, undesirable, reaction of products and consequently improve selectivity. E To increase reaction rates for product inhibited reactions. The various techniques that can be considered are given below. 4.3.1. Catalytic distillation Catalytic distillation (Smith, 1984) combines reaction and distillation in one vessel using structured catalysts as the enabling element (Fig. 14). The combination results in a constant-pressure boiling system, ensuring precise temperature control in the catalyst zone. The heat of reaction directly vaporizes the reaction products for e$cient energy utilization. By distilling the products from the reactants in the reactor, catalytic distillation breaks the reaction equilibrium barrier; it eliminates the need for additional fractionation and reaction stages, while increasing conversion and improving product quality. Both investment and operating costs are far lower than with conventional reaction followed by distillation. CDMtbe威, CD¹ame威 and CDEtbe威 technologies use catalytic etheri"cation of alkenes with an alcohol to produce high-quality ethers. Currently, there are more than 50 CDTECH威 units operational worldwide. The catalytic distillation technology is currently also applied commercially for aromatics alkylation, benzene removal from reformate fractions, selective desulfurization of FCC gasoline fractions, as well as for various selective hydrogenations (Sy, Smith, Chen, & Dautzenberg, 1993;
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Fig. 13. Composite structured packing reactor.
Fig. 14. Catalytic distillation reactor.
Rock, Gildert, & McGuirk, 1997; Hearn & Putman, 1998). 4.3.2. Catalytic stripping There are instances, where the reaction temperature does not match the boiling points of the reactants or products. However, there is still a need to remove the products from the reaction zone to improve reactor performance. The Sinopec/Lummus bisphenol-A process for the condensation of phenol with acetone uses a CSR, catalytic stripping reactor to solve this problem. This new reaction system promotes the BPA condensation
reaction under highly favorable conditions while simultaneously removing the water of reaction by stripping it from the reaction zone using an inert gas. The CSR (Yu, Zhou, & Tan, 1997) provides a very high phenol to acetone ratio that, together with water removal, leads to very high phenol conversions while producing superior quality BPA. The high phenol conversion also permits a simple, one-stage crystallization system. 4.3.3. Selective extraction Reaction rate and product selectivity in catalytic liquidphase hydrogenation can pro"t from the introduction
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EO using air as O -containing gas. The majority of reactions studied in catalytic membrane reactors have taken advantage of the ability of membranes to selectively remove a product, usually hydrogen, through Pd alloys or by Knudsen di!usion through microporous glass or ceramic membranes. Signi"cant improvements over conventional packed bed reactors (Dittmeyer, 1999) have been demonstrated for EB dehydrogenation reactions. Studies have now started looking at other molecules such as ethane and cyclohexene.
5. Type D multifunctionality: chemical reaction with recirculating solids Fig. 15. Asahi technology for the production of cyclohexene.
of an extractant phase. This strategy has been applied to the selective hydrogenation of benzene to cyclohexene using Pt or Ru catalyst. Introducing an aqueous layer surrounding the catalyst particle (Fig. 15) is found to increase the selectivity towards cyclohexene (the desired product). The intermediate hydrogenation product cyclohexene distributes preferentially into the organic phase (Struijk, 1992) and is thus prevented from further hydrogenation to cyclohexane. Asahi Chemicals has utilized this concept in their commercial process. 4.3.4. Membrane reactors Membrane reactors may well be the future for many equilibrium-limited reactions requiring removal of small molecules such as hydrogen or water. The basic concepts involved are illustrated in Fig. 16 for the production of
Design of a reactor is often dictated by catalyst deactivation kinetics and the time interval between successive regenerations. If this time interval is on the order of 1 year, "xed-bed operation is preferred. If the time interval is on the order of 1 week, swing-type operation using two beds is usually selected, as in naphtha reforming. If the time interval between successive regenerations is on the order of a day to a few hours, then moving-bed operations are considered. If the time is less than an hour, circulating #uid-bed operation is called for. 5.1. Circulating yuid-bed reactor At the heart of the re"nery is the #uid catalytic cracker (FCC). The unit is divided into two main sections: the riser and the regenerator (Fig. 17). The riser is the reactor where the feed oil is injected with #owing catalyst and cracked. At the end of the riser is a stripper and disengager which separates catalyst and hydrocarbons, the latter going on to the fractionator. In FCC operations, coke formation on the catalyst is extremely fast. The catalyst is usually onstream for a few seconds before it deactivates. Catalytic cracking is an endothermic reaction. The need to regenerate the catalyst rapidly, coupled with the transport of heat via solid particles, led to the development of the FCC circulating #uid-bed reactor (Fig. 17). The catalytic process takes place in the riser and the products and the catalyst are immediately separated: the catalyst particles being sent to the regenerator, where the coke is combusted and the exothermic heat of reaction carried back to the riser reactor. 5.2. Moving-bed reactor
Fig. 16. Conceptual membrane reactor for ethylene oxide.
Continuous reforming (Trambouze, van Landeghem, & Wanquier, 1984) permits higher operating severity by maintaining the high catalyst activity of near-fresh catalyst through regeneration cycles of a few days. A moving-bed reactor system (Fig. 18) has the advantage of maintaining production while the catalyst is removed or
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Fig. 17. Modern FCC unit.
Fig. 18. Industrial moving-bed reactors.
replaced. Catalyst particles pass by gravity through one or more reactors in a moving bed and are conveyed to a continuous regeneration zone. Continuous catalyst regeneration generally is e!ected by passing catalyst particles downwardly by gravity in a moving-bed mode (Fig. 18) through various treatment zones in a regeneration vessel. Although movement of catalyst through the zones is often designated as continuous, in practice it is semi-continuous in the sense that relatively small amounts of catalyst particles are transferred at closely
spaced points in time. For example, one batch per minute may be withdrawn from the bottom of a reaction zone and withdrawal may take one-half minute; e.g., catalyst particles #ow for one-half minute in the 1-min period. Since the inventory in the reaction and regeneration zones generally is large in relation to the batch size, the catalyst bed may be envisaged as moving continuously. Moving-bed reactors have also been introduced in the upgrading of residual oil facilities (Pegels & Wij!els, 1974) where catalyst replacement is required because
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of deactivation by metal deposition and coking of the catalyst.
A new approach (Jacoby, 1999) is the hydrocarbon trap which absorbs the HCs emitted during the cold start period, releasing them to the catalyst only when it is hot (Fig. 19).
6. Recent developments and opportunities 6.3. Ultra-low sulfur fuels technology The following examples illustrate that the area of process intensi"cation using novel catalysts integrated with advanced reactors is getting increased attention and is opening new opportunities. 6.1. Programmed release catalyst Wax deposition in sub-sea pipelines is a common problem in transporting crude oil. The primary challenge in clearing pipeline blockages due to wax deposits is in supplying heat to regions farther down the pipeline, which are more prone to wax deposition. A novel catalytic system (Fogler, 1999) has been developed where the catalyst is introduced with the feed mixture but the catalytic reaction is deliberately delayed by incorporating a physical barrier over the active sites. The catalyst is encapsulated by a polymeric wall, which slowly dissolves in the crude oil mix over a period of 2 h, thus successfully delaying the exothermic chemical reaction to downstream regions of the transportation pipelines. 6.2. Fast light-ow catalytic converters To meet demands for stricter pollutant standards, researchers are investigating ways of boosting catalytic converter e$ciencies. In this respect, cold startup of the car poses a signi"cant challenge * especially in achieving California's 0.04 gm/mile ULEV hydrocarbon levels standards. In order to completely oxidize hydrocarbon emissions during the "rst 50 s of the cold start, a catalytic converter in the typical under-#oor position needs to be active at temperatures less than 1003C. One approach is to heat the catalyst chemically or electrically to attain temperature required for hydrocarbon oxidation.
On May 1, 1999, the US EPA proposed regulations limiting the gasoline sulfur level to be an average of 30 ppm to become e!ective in the 2004}2006 time frame. Most US re"ners marketing gasoline will need to add new desulfurization schemes to meet this stringent requirements. Conventional hydrodesulfurization is carried out in co-current down #ow trickle-bed reactors. The desulfurization reaction is product inhibited. As the reaction progresses down the reactor, the reaction slows down considerably for two reasons: (1) the hydrogen partial pressure reduces as hydrogen is consumed and (2) the H S formed, inhibits the reaction rate. Deep desulfur ization in conventional trickle-bed reactors is thus extremely di$cult (Sutikno, 1999). However, if the trickle-bed reactor were to be run in counter-current #ow mode, the major part of the catalyst bed would operate in an H S lean mode thus allowing ultra-low sulfur levels to be achieved economically in the hydrocarbon liquid phase. One could conceive a multifunctional catalytic system that would allow counter #ow trickle-bed operation without a high *P penalty. Ring-shaped catalysts have been used in some desulfurization processes such as SynSat, but the bed depths are small as ring-shaped catalysts cannot take much weight. Structured catalysts such as Sulzer's Katapak-S or M-Series or the composite structured packing described above may be worthwhile investigation targets. HDS performs well in removing 80}90% of sulfurbearing species in distillates, but increasingly poor at removing the remaining substantial polyaromatic sulfur-bearing compounds. The latter species are most susceptible to catalytic biodesulfurization (Pacheco et al., 1999). This clearly suggests the application of both hydrodesulfurization and biodesulfurization hybrids to achieve ultra-low sulfur levels in fuels. 6.4. Microtechnology
Fig. 19. Zeolite trap for cold start emission reduction.
Microreactors' narrow channels and thin walls (Fig. 20) make them especially e$cient for processes (Shanley, 1997) that are mass- and heat-transfer limited. In addition, residence times become strikingly short in microreactors. The laminar #ow in microreactor channels is yet another bene"t, leading to lower pressure drops than traditional reactors. While CPI applications of microreactors may be a decade away, there are niche areas where microreactors easily outperform traditional reactive systems. In this respect, recent studies have
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Fig. 20. Microreactor technology.
Fig. 21. Microengineered catalyst technology.
stressed the role and potential of fuel cells as an electric power source to run the motors of hybrid vehicles. However, the current infrastructure for transportation fuels does not have the capability to deliver the hydrogen used in fuel cells directly to such hybrid vehicles. For the near-term viability of such systems, technologies are necessary to convert gasoline, natural gas, and other hydrocarbon fuels, which can be handled by the current distribution system, to hydrogen onboard the vehicle. (Because of its low density, hydrogen requires extensive storage space in its pure form.) Microfabricated reformers could be developed to meet this need: compact size would make them ideal for cars. With heat transfer coe$cients that are an order-of-magnitude greater
than current heat exchangers, these multifunctional compact miniature reactors could become the enabling element for successful employment of hydrogen-based automobiles. 6.5. Microengineered catalyst systems A team of ABB chemical engineers working with university scientists has developed a novel catalytic material (Bayegan, 1999) that improves reactor and catalyst performance by e!ectively utilizing extremely small catalyst particles in "xed bed reactors. The new system, called micro-engineered catalyst (MEC), is based on mechanically robust, highly porous structures made of
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micro"bers in which small (nm) catalyst particles are entrapped (Fig. 21). The new material with the microengineered catalyst can be formed into a variety of shapes and structures, hydrodynamically tailored for various applications. In this way, excellent mass transfer, improved heat transfer and optimal pressure drop are achieved. Using computational #uid dynamics (CFD) and experimental modeling, various novel designs for reactor internals have been developed. Furthermore, the development of a new electrophoretic coating procedure enables the coating of each "ber within three-dimensional reactor structures. So far, signi"cant progress has been achieved toward a variety of industrial applications. Commercialization e!orts are currently underway to bring MEC to market.
7. Conclusions and recommendations Through the development of highly active and selective catalysts, feedstock utilization can be optimized which leads to a signi"cant reduction in energy consumption per unit weight of product and thus to a reduction in the build-up of greenhouse gas emissions. By employing multifunctional reactors, energy reduction is achieved by integrating several process steps in one vessel, hence reducing the overall utility requirements of the process. By in situ removal of products from the reaction zone, multifunctional reactors can reduce by-product generation from a process by lowering the rates of multiple or consecutive reactions. Waste generation in the chemical industry and the environmental burden caused as a result is mitigated. PI through use of multifunctional reactors permits signi"cant reductions in both investment and plant operating costs by optimizing the process. In an era of emaciated pro"t margins, it allows chemical producers more leverage in competing in the global marketplace. Focussing on the integration of new catalyst developments and advances in reaction engineering may lead to the development of additional or improved multifunctional reactors. These can assist in overcoming some of the challenges to the process industries in the future. By systematic thinking about the type of multifunctionality that is required, it will be easier to propose viable solutions. As illustrated in this paper, each type of multifunctionality requires `out-of-the-boxa thinking and thus R&D teams with a larger plurality in basic disciplines.
Acknowledgements The author wishes to thank Mr. Mitrajit Mukherjee for his assistance in creating the basic outline of this paper with its many illustrations. Dr. Nelleke van der Puil, Dr. Rudolf Overbeek and Mr. Robert Trubac are
acknowledged for their contributions regarding recent developments at ABB Lummus Global Inc. References Antos, G. J., Aitani, A. M., & Parera, J. M. (1995). Catalytic Naphtha Rexning. New York: Marcel Dekker, Inc. Bayegan, M. (1999). R&D at ABB * achievements 1998 and vision of the future. ABB Technology Press conference, February, DaK ttwil, Switzerland. Boelt, H., Walser, F., Merz, G., Stabel, U., & Wunsch, G. (1994). Heated reactors containing catalyst-xlled tubes for performing endothermic reactions. German Patent DE 4,306,648. Dautzenberg, F. M. (1999). Quality principles for catalyst testing during process development. NATO Advanced Study Institute Conference, Villamoura, Portugal, July 19}24. Dautzenberg, F. M., & Mukherjee, M. (1998). Current challenges in styrene technology. Large chemical plants * 10th symposium, Antwerp, Belgium, September. Dautzenberg, F. M., Naber, J. E., & van Ginneken, A. J. (1971). Shell's #ue gas desulfurization process. Chemical Engineering Progress, 67, 86}91. Davis, M. E. (1994). Reaction chemistry and reaction engineering principles in catalyst design. Chemical Engineering Science, 49, 3971}3980. Dittmeyer, R. (1999). Dehydrogenation of ethylbenzene in palladium composite membrane reactors. Chemical Engineering Science, 54, 1431}1439. Ercan, C., Dautzenberg, F. M., Yeh, C. Y., & Barner, H. E. (1998). Masstransfer e!ects on liquid-phase alkylation of benzene with zeolite catalysts. Industrial and Engineering Chemistry, 37(5), 1724}1728. Feldman, R. J., Dufallo, J., Tucci, E. L., & Balogh, P. (1992). Commercial performance of the Houdry Catoxn process. Paper presented at 1992 National Petroleum Re"ners Association Meeting, New Orleans, March 22}24. Fogler, H. S. (1999). Fused chemical reactions, Paper presented at the A.I.Ch.E. Annual Meeting. Haggin, J. (1995). Innovations in catalysis create environmentally friendly THF process. C&EN, April 3. Hearn, D., & Putman, H. M. (1998). Hydrodesulfurization process utilizing a distillation column reactor. US Patent 5,779,883. Jacoby, M. (1999). Getting auto exhausts to pristine. C&EN, January 25. Matros, Y. S. (1985). Unsteady processes in catalytic reactors. Amsterdam, The Netherlands: Elsevier. Moeller, K. P., & O'Connor, C. T. (1996). Gas}solid mass transfer in a jet-loop reactor. A.I.Ch.E. Journal, 42, 1187}1190. Mortensen, K. (1998). Current Opinion in colloid and Interface Science, 3, 12. Mulder, A. (1996). Catalytic combustion in a sintered metal reactor with integrated heat exchanger. Applied Thermal Engineering. Murrell, L. L., Overbeek, R. A., Chang, Y-f., van der Puil, N., & Yeh, C. Y. (1999). Method of making molecular sieves and novel molecular sieve composites. US Patent 6,004,527. Murrell, L. L., Overbeek, R. A., & Khonsari, A. M. (1999). Catalyst and method of preparation. US Patent 5,935,889. Pacheco, M. A. et al. (1999). Recent advances in biodesulfurization of diesel fuel. 1999 NPRA Annual Meeting, March 21}23, San Antonio, TX. Pegels, A. A., & Wij!els, J-B. (1974). Process for the catalytic treatment of hydrocarbon oils. US Patent 3,826,737. Roby, A. K. (1996). Oxidize safely with pure oxygen. Chemtech, 26. Rock, K., Gildert, G. R., & McGuirk, T. (1997). Catalytic distillation extends its reach. Chemical Engineering, 78, 78}84. Shanley, A. (1997). Microreactors "nd new niches. Chemical Engineering.
F. M. Dautzenberg, M. Mukherjee / Chemical Engineering Science 56 (2001) 251}267 Smith, L. A. (1984). Catalytic distillation structure. US Patent 4,443,559. Strangio, V. A., Dautzenberg, F. M., Calis, H. P. A., & Gupta, A. (1999). Fixed catalytic bed reactor, PCT Application WO99/48604. Struijk, J. (1992). Catalytic routes to cyclohexene. Ph.D. thesis in Chemical Engineering, Delft University. Sutikno, T. (1999). Optimal HDS for lower sulfur gasoline depends on several factors. Oil & Gas Journal, June, 55}59.
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Sy, A., Smith, L., Chen, J., & Dautzenberg, F. M. (1993). Catalytic distillation route for cumene. DeWitt Petrochemical Review, Houston, March 23}25. Trambouze, P., van Landeghem, H., & Wanquier, J. P. (1984). Les Reacteurs Chimiques, IFP Publication No. 26, Paris: Editions Technip. Yu, S., Zhou, A., & Tan, Q. (1997). Simulation of multistage catalytic stripping with a non-equilibrium stage model. Computers in Chemical Engineering, 21, 409}415.
Process intensi"cation using multifunctional reactors F. M. Dautzenberg*, M. Mukherjee ABB Lummus Global Inc., Technology Development Center, 1515 Broad Street, Bloomxeld, NJ 07003-3096, USA
Abstract Process Intensi"cation is the strategy of making signi"cant reductions in the size of a chemical plant in order to achieve a given production objective. Innovations in catalytic reactors, which constitute the heart of such process technologies, are often the preferred starting point. Over the last two decades, innovative multifunctional reactor systems have been developed to intensify chemical processes by synergistically combining chemical reaction with momentum, heat and mass transport in a single vessel. This paper aims to emphasize an interdisciplinary approach to multifunctional reactor design. Examples from commercially practiced technologies are used to illustrate how process intensi"cation can be achieved through multifunctional reactors. Furthermore, new developments in this area will be highlighted. 2001 Elsevier Science Ltd. All rights reserved.
1. Introduction Process intensi"cation (PI) is a term used to describe the strategy of making dramatic reductions in the physical size of a chemical plant while achieving a given production objective. The concept was pioneered by ICI during the late 1970s when the primary goal was to reduce the capital cost of a production system. Currently, increased pressures to remain cost competitive in an environmentally responsible manner have led to a renewed interest in PI. PI is being applied throughout the chemical process industry (CPI) to reduce investment and operating costs of chemical plants to increase pro"tability and mitigate greenhouse gas emissions. Multifunctional reactive systems (catalysts, reactors, etc.) allow a unique way of achieving PI in the chemical process and re"ning industry. By its very de"nition, several `functionsa or processes are designed to occur simultaneously in multifunctional reactors. One aims at on optimum integration of mass, heat and momentum transfer within a single reactor vessel. Multifunctional reactors are not new to the chemical industry. They have been in use in the industry for a very long time. Traditional FCC units are a good example,
* Corresponding author. Tel.: #1-973-893-3319; fax: #1-973-8932745. E-mail address: [email protected] (F. M. Dautzenberg).
incorporating several `functionsa in one catalytic reactor such as heat management, catalyst transport and two chemical reactions (cracking and coke removal by oxidation). Only recently have these reactors been formally classi"ed as being multifunctional and the large bene"ts derived from PI have been acknowledged. Consequences of using multifunctional reactors are reduced investment costs and signi"cant energy recovery or savings. Furthermore, improved product selectivity leads to a reduction in raw material consumption and, hence, operating costs. To achieve optimal performance with multifunctional reactors, it is important to understand where the integration of functionalities occurs. Referring to Fig. 1 (Dautzenberg, 1999), the following classi"cations can be made. E Type A multifunctionality operates at the catalyst level: may be introduced by combining catalytic properties with an engineered catalyst structure. E Type B multifunctionality at the reaction inter-phase: by integrating chemical reaction with enhanced interphase transport properties. E Type C multifunctionality at the intra-reactor level: by marrying chemical reaction with intra-reactor unit processes such as heat transfer or separations. E Type D multifunctionality at inter-reactor level: by combining two reactor operations using solids recirculation. This paper gives an overview of multifunctional reactors. It stresses the strategy employed in achieving
0009-2509/01/$ - see front matter 2001 Elsevier Science Ltd. All rights reserved. PII: S 0 0 0 9 - 2 5 0 9 ( 0 0 ) 0 0 2 2 8 - 1
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Fig. 1. Catalyst particle in a reacting medium.
the multifunctional character of these reactors, explaining the type of multifunctionality achieved and the most important parameters involved in this task. Recommendations for future designs will also be indicated.
the catalyst to perform functions that could otherwise only be done by using very expensive process equipment (see Table 1). 2.1. Bifunctional catalysts
2. Type A multifunctionality: chemical reaction and catalyst particle design The performance of a majority of chemical reactors (and hence the process) is signi"cantly in#uenced by the performance of the catalyst. A sizeable amount of research dollars is thus spent in upgrading performance of industrial catalysts. Catalyst research has been devoted to increase the catalyst activity and selectivity to improve process economics and reduce greenhouse gas emissions through better feedstock utilization. Researchers have discovered the impressive bene"ts that can be derived from the intensifying processes by using multifunctional catalysts. This can be achieved by combining a catalyst function with: E E E E E
Secondary catalytic function. Enhanced transport characteristics. Novel catalyst morphologies. Energy transfer properties. Reactant transfer properties.
The synergy through combining two processes in a single catalyst particle can lead to signi"cant process advantages. In some cases, this multifunctional behavior allows
Catalytic reforming is an important process in the re"ning industry in which the octane number of a hydrocarbon fraction is increased without signi"cantly changing its molecular weight. It is the primary source of aromatic chemicals (BTX). The main reactions that occur are dehydrogenation followed by isomerization and dehydrocyclization. For these reactions to occur, two types of sites are needed: metallic sites for dehydrogenation or cyclization, and acid sites for isomerization. Bifunctional catalysts (Antos, Aitani & Parera, 1995) consisting of metallic sites, e.g. Pt, and acid sites (e.g. alumina) are typically used for these applications. The synergistic bene"ts of using such dual functional catalysts become clear when one compares the performance of a physical mixture of two monofunctional catalysts. The physical mixture of Pt on carbon catalyst and acidic alumina has a much lower activity and deactivates rapidly. 2.2. Attrition resistant catalysts In circulating #uid-bed reactors (CFB), rapid transport of catalyst particles between vessels subjects the particles to strong attrition forces. Attrition resistance then becomes an important criterion. DuPont's recently
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Table 1 Examples of type A multifunctionality Application
Function 1
Function 2
Bene"ts/Impact
Catalytic reforming
Catalyse dehydrogenation
Catalyse molecular structure rearrangements
High space time yields
Butane oxidation
Catalyse oxidation
Enhance attrition resistance
Temperature control and selectivity enhancement
Hydroformulation of ole"ns
Catalyse reaction
Encapsulate homogeneous catalyst
Easy product separation from catalyst
Re"nery alkylation
Catalyse ole"n alkylation of para"ns
Enhance di!usion and stability
Better product quality and longer catalyst life
Fluid catalytic cracking
Catalyse cracking reaction
Provide energy transfer
Energy savings
Fig. 2. Attrition resistant VPO catalyst for butane.
commercialized tetrahydrofuran (THF) process shows a unique way of overcoming this challenge by endowing the catalyst with multifunctional behavior. The vanadium phosphate (VPO) catalyst traditionally used for this process is too weak to withstand forces associated with cycling between the #uidized bed regenerator and the dense phase riser reactor, where the gas velocities are in the &1 m/s range. Attrition resistance is imparted to the catalyst by spray drying this material with a silica hydrogel (Haggin, 1995) under conditions that allow the silica to migrate to the outer regions, thus encapsulating the active VPO catalyst into a porous silica shell (Fig. 2)
whose pore openings allow product and reactant molecules to di!use in and out of the active catalyst zone. 2.3. Encapsulated homogeneous catalysts A technique of `heterogenizinga homogeneous catalyst involves encapsulating the catalyst by a porous shell. This technique has been applied to a hydroformylation reaction where a Rhodium-based homogeneous catalyst is synthesized inside a hollow silica micro-sphere with pores (Davis, 1994) of the order of 1 nm. The metal complex is too bulky to di!use through the pores of the
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Fig. 3. Encapsulated homogeneous catalysts.
micro-sphere (Fig. 3). Deliberate introduction of di!usional limitations for CO to the liquid has potential bene"cial e!ects for product selectivity. Recently, another approach has been proposed (Mortensen, 1998). In this case (Fig. 3, right-hand side), surfactants are used to create micellar structures of nanodimensions that can be designed to contain homogeneous catalysts. Among others, the oxidation of propylene to propylene oxide with hydrogen peroxide is being developed at the Technical University of Eindhoven, Netherlands.
Fig. 4. Thin "lm catalyst.
2.4. Thin xlm catalyst A thin layer of active catalyst is deposited precisely on a solid support such that the catalytic material does not penetrate the support. The small thickness of the catalyst layer leads to a low di!usion barrier (Fig. 4). This concept has been applied to the lummus solid acid gasoline alkylation technology (Murrell, Overbeek, & Khonsari, 1999). By depressing the formation of heavy by-products through a highly e!ective catalyst layer, the catalyst life and product selectivity is greatly enhanced. 2.5. Energy transfer FCC catalysts are a prime example of multifunctional reactive systems. The active 1 lm zeolite crystals are embedded in a dense matrix of 50}100 lm silica}alumina particles. This allows the zeolite to survive the severe mechanical forces while being recirculated from the riser-tube reactor to the regenerator and back. Catalytic cracking is an endothermic reaction that varies from 5503C at the inlet of the reactor to 4753C at the outlet. Heat for the cracking reaction is supplied partly by the catalyst particle itself * the catalyst particle picks up
heat in the regenerator through burning of the coke on its surface &0.8 wt% at the inlet and less than 0.1 wt% at the outlet of the regenerator * and releases it in the riser reactor. 2.6. Dry zeolite catalyst Mass transfer limitation of reactants and products in zeolite catalyst extrudates was observed (Ercan, Dautzenberg, Yeh & Barner, 1998) for the alkylation of benzene to ethylbenzene and cumene. In laboratory tests, signi"cant performance enhancement was observed with decreasing catalyst particle sizes. Due to process limitations (pressure drop increase) and catalyst manufacturing limitations, catalyst particle size reduction cannot be implemented easily in industry. ABB Lummus Global Inc. developed a zeolite synthesis procedure (Murrell, Overbeek, Chang, van der Puil & Yeh, 1999) that provides ready-to-use zeolite catalyst particles with ultra small crystals and a high meso- and macropore volume. These catalyst features reduced mass transfer limitations. Furthermore, the preparation method is faster, uses raw materials more e$ciently than the conventional zeolite
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Fig. 5. `Dry zeolitea catalyst preparation.
synthesis technique methods and eliminates the "ltration step to collect the product. The new method employs mesoporous silica}alumina precursor particles that are converted to catalyst particles with the same morphology as the precursor particles (Fig. 5). This is achieved by e!ecting zeolite crystallization in the absence of an external liquid phase. The total pore volume that is formed by the micro-, meso- and macropores of the starting particle is maintained because the particle shape and size of the starting material are maintained upon conversion. Compared to conventional zeolite catalysts, the volume of meso- and macropores in catalyst particles from the novel synthesis procedure is enhanced. Moreover, the zeolite crystals obtained by this method are a factor of 3}10 smaller than those from commercial synthesis methods. The zeolite catalyst particles from the novel procedure showed 2}3 times higher alkylation activity
(per unit weight) than commercial zeolite catalyst particles.
3. Type B multifunctionality: chemical reaction and inter-phase transport A general classi"cation of Type B multifunctionality is provided in Fig. 6. 3.1. Inter-phase mass transfer Inter-phase mass transfer is the rate-limiting step for many hydrocarbon oxidations in the gas and liquid phase. In these instances, the overall reaction rate is not a function of catalyst activity but depends on the #uid Reynolds number. For commercial scale reactors,
Fig. 6. Classi"cation of Type B multifunctionality.
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Fig. 8. Buss' jet loop reactor.
Fig. 7. Liquid oxidation reactor.
inter-phase resistance can become extremely large requiring long residence times to achieve economic conversions. In these cases, the incorporation of a mass-transfer device in the reactor can lead to signi"cant process advantages. 3.1.1. Liquid oxidation reactor As stated earlier, many hydrocarbon oxidations are run in the mass-transfer-limited regime. The concentration of oxygen in the dispersed bubbles has a direct e!ect on the overall rate. To overcome this limitation, Praxair Inc. has developed a unique liquid oxidation reactor (LOR) (Fig. 7), a gas}liquid reactor that uses a downward pumping impeller contained within a draft tube to disperse and circulate the reactive gas. By e!ectively mixing the hydrocarbon with oxygen through recirculation of oxygen bubbles, the LOR (Roby, 1996) achieves better gas utilization, higher reaction rates and increased energy e$ciency compared with traditional CSTRs. The impeller draws the bubbles downward and redisperses them. The #ow is further directed by two sets of vertical ba%es, one above the inlet of the draft tube and one at the discharge end. 3.1.2. Jet loop (buss) reactors A loop reactor consists of a pipe through which process #uids are recirculated. Often it contains a multitubular heat exchanger and in-line disperser (e.g. jet or ori"ce). By incorporating a jet, high-energy dissipation rates can be conveniently supplied which are 10}100 times higher than those available in stirred tanks. Such dispersion intensities are useful to minimize the size and cost of the reactor, and to maximize yields when the key reaction rates are intrinsically rapid but controlled by
inter-phase mass-transfer rates. The Buss reactor (Fig. 8) is based on this principle and if often used for liquid hydrogenations (Moeller & O'Connor, 1996). It has the advantage of rapid gas-liquid mass transfer in the initial zone, combined with high heat removal capability in the tubular heat exchanger. 3.2. Inter-phase energy transfer A large percentage of all useful intermediate petrochemicals are products of selective oxidations of hydrocarbons. These reactions, however, are highly exothermic, generating large amounts of heat. Multitubular reactors have been used as a way of controlling the reaction exotherm by conducting the heat away from the packed bed to the shell-side heat-transfer #uid. The radial temperature pro"le for an exothermic reaction in a packed tubular-bed reactor is shown in Fig. 9. There is a steep gradient near the inside wall and a nearly parabolic temperature pro"le over the rest of the catalyst bed. This phenomenon is caused as a result of a high voidage near the tube wall. 3.2.1. Sintered metal reactor The sintered metal reactor overcomes the steep temperature gradient at the wall in tubular reactors by sintering metal particles to the tube wall. By doing so, the wall heat transfer resistance (by inter-phase forced convection) is largely eliminated. In addition, the heat transfer in the packed bed is mainly by conduction through a continuous metal matrix (Mulder, 1996). Since thermal conductivity of metals is an order-of-magnitude greater than conduction through loosely packed catalyst particles, the temperature control in a sintered metal reactor is optimal. This interesting new reactor concept still awaits commercial demonstration. 3.3. Inter-phase momentum transfer Design and performance of commercial-scale chemical reactors is often dictated by the inter-phase transport
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Fig. 9. Porosity, velocity and temperature in tubular reactor.
rates between the reacting surface and the bulk #uid. In several processes, these rates can be much lower than the chemical reaction rate leading to low reactor e$ciencies. For instance, inter-phase momentum transport (characterized by the parameter `f a friction factor) in packed bed reactors gives rise to pressure drop. High system pressure drops leads to higher compression costs and, in equilibrium limited reactions, a lower overall conversion. PI can then be achieved by incorporating a catalytic function integrated into a system with a low `f a using novel catalyst/morphologies and reactor con"gurations. 3.3.1. Monolithic reactors Flow through the interstices in packed beds is tortuous and disordered. The frictional resistance comes from two sources: E drag friction as a result of viscous e!ects ( fakl and surface-to-volume ratio) and E from friction as a result of inertial forces ( fakl and geometry of the surface). For smooth tubes, however, the only contribution of the friction factor is the drag friction. The inter-phase momentum resistance, as measured by the pressure drop for equivalent Reynolds numbers, is thus much higher in packed beds than in smooth tubes. This bene"t is exploited in monolithic reactors. Monolith reactors consist of straight channels in parallel (Fig. 12) with small diameters. The resulting laminar #ow (to Reynolds numbers of 500) does not show the kinetic energy losses that occur
in "xed-beds due to inertial forces at comparable super"cial velocities. The ability to o!er high surface-tovolume ratios at low-pressure drop makes it the reactor of choice for catalytic combustion processes. Almost all automobile catalytic converters today use monolithic reactors to control exhaust emissions (Bayegan, 1999).
4. Type C multifunctionality: chemical reaction and intra-reactor processes The majority of multifunctional reactors operating in the process and re"ning industry fall under this category, a combination of chemical reaction coupled with an intra-reactor process operating at the macro-scale. The examples stretch from well-established multi-tubular reactors for reaction temperature control to the more recent work investigating reactions using in situ supercritical #uids as extractants. In most cases, the integration of the unit process with the chemical reaction in the same vessel allows bene"ts in process operability that are impossible to achieve merely by externally coupling the two operations. Typical unit operations that are used in multifunctional reactors can be subdivided into three categories: E Chemical reaction with intra-reactor energy transfer. E Chemical reaction with intra-reactor momentum transfer. E Chemical reaction with intra-reactor mass transfer. Table 2 provides examples of Type C multifunctionality.
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Table 2 Examples of type C multifunctionality Intra-reactor transfer
Process
Application
Impact/bene"t
Energy transfer
Heat transfer reactor Energy storage
Multi-tubular control; higher yields Cato"n, VOC
Temperature
Filtration
Catalytic "lter for No removal V Radial #ow rector
Reduced investment
Distillation
Catalytic distillation
Stripping
Catalytic stripping reactor
Reduced investment Higher yields Better temperature control Product purity Smaller reactor
Momentum transfer
Low pressure drop Mass transfer
4.1. Chemical reaction with intra-reactor energy transfer Chemical reactions tend to either generate heat (exothermic reaction) or absorb heat (endothermic reaction). Oxidation, hydrogenation and alkylation are typical examples of exothermic reactions while catalytic dehydrogenation and catalytic cracking reactions are endothermic reactions Temperature control of reactors is important for several reasons: 1. For endothermic processes, the reaction absorbs heat and cools the reactants to e!ectively quench the product reaction. 2. Exothermic reactions may run away at high conversions if the heat is not dissipated. 3. Equilibrium-limited reactions are temperature-dependent * higher outlet temperatures typically increase equilibrium conversions. 4. Temperature control reduces loss of selectivity by reducing thermal degradation of products and feedstock. 5. Better catalyst utilization is achieved through temperature pro"ling in the reactor. 4.1.1. Reaction in a heat exchange device Coupling of a reactive system with heat transfer equipment has been practiced in the industry for a long time. Tubular reactors have been used extensively in the process industry to control temperatures for both exothermic and endothermic reactions. The catalyst pellets reside in tubes in a shell and tube heat exchanger. A heat transfer #uid is then circulated on the shell side; highpressure steam is often used as a heat transfer medium. Catalyst tubes may number from 500 to 10,000 per reactor. An interesting new development has been introduced by Linde: the tube bundle is composed of counter-wound
Reduced energy costs
Higher yields
spirals in which upwardly #owing water is evaporated (Boelt, Walser, Merz, Stabel, & Wunsch, 1994). The tubes run into a vertical vapor drum located at the reactor head. Advantages in construction and in heat transfer from the reaction gas to the tubes of the bundle are claimed. 4.1.2. Coupling of exothermic and endothermic reactions Dehydrogenation reactions are endothermic in nature but release a mole of hydrogen per mole of ole"n formed. A unique technology developed by Lummus/UOP called SMART couples the endothermic dehydrogenation reaction with an exothermic hydrogen combustion step to balance the energy requirements in the styrene reactor. Not only does this create a better temperature pro"le in the catalyst bed by removing hydrogen in the reactor, it reduces the reverse reaction leading to high EB conversions. 4.1.3. Energy storage The catalyst bed itself is sometimes used to store heat as in the case of C and C dehydrogenation reactions practiced in Lummus' Cato"n technology (Feldman, Dufallo, Tucci, & Balogh, 1992) (Fig. 10). The Cato"n approach utilizes "xed-bed reactors operated under vacuum conditions to achieve favorable equilibrium. The heat required for the endothermic reaction is supplied partly by the sensible heat of the feed and by the thermal energy contained in the catalyst bed itself. The "xed beds are cycled between reaction and reheat, which maintains the energy balance in the reactor. 4.1.4. Transient reactor operation Reactors can be operated advantageously with moving thermal fronts that are created by periodic #ow reversal (Matros, 1985). Low-level contaminants or waste
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Fig. 10. Cato"n reactor system.
Fig. 11. Reverse #ow reactor concept.
products (VOCs) can be e$ciently removed in adiabatic "xed beds with periodic reversal by taking advantage of higher outlet temperatures generated in earlier cycles to accelerate exothermic reactions (Fig. 11). Energy and cost savings are e!ected by this substitution of internal heat transfer for external exchange. 4.2. Chemical reaction with intra-reactor momentum transfer Some chemical processes require high gas #ow rates through reactors, as in the case of ammonia synthesis or production of styrene from ethylbenzene. In the catalytic treatment of #ue gases to reduce atmospheric pollutants, e.g., sulfur and nitrogen oxides, very large volumes of air at relatively low pressures must be handled. In all these situations, pressure drop through traditional "xed-bed reactors will become excessively large, which calls for reactor design with very low-pressure-drop.
4.2.1. Parallel passage reactor Particles of a catalyst or a regenerable adsorbent are enclosed in wire screen envelopes, which are mounted in a parallel fashion. Gas #ows in the empty passages between envelopes. The straight, unobstructed gas channels give rise to a low-pressure-drop and good tolerance (Fig. 12). Pollutant molecules are transported to the catalyst or adsorbent by lateral di!usion. This type of reactor has been applied in the shell #ue gas desulfurization process (Dautzenberg, Naber, & van Ginneken, 1971), which is based on the use of a regenerable copper adsorbent. 4.2.2. Radial yow reactor High temperatures and low operating pressures favor dehydrogenation of ethylbenzene to styrene. The Lummus-UOP Classic Styrene process (Dautzenberg & Mukherjee, 1998) uses two radial #ow reactors with an inter-stage heater to get around the constraints of the
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Fig. 12. Low-pressure-drop reactor systems.
dehydrogenation reaction. By increasing the #ow surface area, the mass-#ux of reactants through the bed diminishes signi"cantly resulting in low bed pressure drop. An optimally designed de#ector housed in the annular space of the adiabatic reactor ensures even #ow distribution from the very bottom of the bed to the very top. 4.2.3. Composite structured packing reactor The composite structured packing (CSP) reactor (Strangio, Dautzenberg, Calis, & Gupta, 1999) achieves low-pressure-drop through the regular vertical stacking of catalyst particles inside parallel channels. The regular stacking, shown schematically in Fig. 13, results in a minimization of the momentum loss which normally occurs in random packed beds due to #ow directional changes. The CSP can be considered a commercially viable extension of the `bead-stringa reactor concept proposed by co-workers at the Technical University of Delft in Delft, Netherlands. The low-pressure-drop/length of the CSP o!ers considerable advantage to the radial #ow reactors in that it allows the use of several catalyst beds with inter-stage reheat within one reactor shell, resulting in a higher overall average operating temperature and a higher styrene yield. 4.3. Chemical reaction with intra-reactor mass transfer Chemical reaction with mass transfer at the intrareactor level is mainly concerned with in situ separation of products from the reaction zone. The main reasons for considering in situ product removal are
E To enhance conversion in equilibrium limited reactions by shifting the equilibrium to the right. E To prevent further, undesirable, reaction of products and consequently improve selectivity. E To increase reaction rates for product inhibited reactions. The various techniques that can be considered are given below. 4.3.1. Catalytic distillation Catalytic distillation (Smith, 1984) combines reaction and distillation in one vessel using structured catalysts as the enabling element (Fig. 14). The combination results in a constant-pressure boiling system, ensuring precise temperature control in the catalyst zone. The heat of reaction directly vaporizes the reaction products for e$cient energy utilization. By distilling the products from the reactants in the reactor, catalytic distillation breaks the reaction equilibrium barrier; it eliminates the need for additional fractionation and reaction stages, while increasing conversion and improving product quality. Both investment and operating costs are far lower than with conventional reaction followed by distillation. CDMtbe威, CD¹ame威 and CDEtbe威 technologies use catalytic etheri"cation of alkenes with an alcohol to produce high-quality ethers. Currently, there are more than 50 CDTECH威 units operational worldwide. The catalytic distillation technology is currently also applied commercially for aromatics alkylation, benzene removal from reformate fractions, selective desulfurization of FCC gasoline fractions, as well as for various selective hydrogenations (Sy, Smith, Chen, & Dautzenberg, 1993;
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Fig. 13. Composite structured packing reactor.
Fig. 14. Catalytic distillation reactor.
Rock, Gildert, & McGuirk, 1997; Hearn & Putman, 1998). 4.3.2. Catalytic stripping There are instances, where the reaction temperature does not match the boiling points of the reactants or products. However, there is still a need to remove the products from the reaction zone to improve reactor performance. The Sinopec/Lummus bisphenol-A process for the condensation of phenol with acetone uses a CSR, catalytic stripping reactor to solve this problem. This new reaction system promotes the BPA condensation
reaction under highly favorable conditions while simultaneously removing the water of reaction by stripping it from the reaction zone using an inert gas. The CSR (Yu, Zhou, & Tan, 1997) provides a very high phenol to acetone ratio that, together with water removal, leads to very high phenol conversions while producing superior quality BPA. The high phenol conversion also permits a simple, one-stage crystallization system. 4.3.3. Selective extraction Reaction rate and product selectivity in catalytic liquidphase hydrogenation can pro"t from the introduction
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EO using air as O -containing gas. The majority of reactions studied in catalytic membrane reactors have taken advantage of the ability of membranes to selectively remove a product, usually hydrogen, through Pd alloys or by Knudsen di!usion through microporous glass or ceramic membranes. Signi"cant improvements over conventional packed bed reactors (Dittmeyer, 1999) have been demonstrated for EB dehydrogenation reactions. Studies have now started looking at other molecules such as ethane and cyclohexene.
5. Type D multifunctionality: chemical reaction with recirculating solids Fig. 15. Asahi technology for the production of cyclohexene.
of an extractant phase. This strategy has been applied to the selective hydrogenation of benzene to cyclohexene using Pt or Ru catalyst. Introducing an aqueous layer surrounding the catalyst particle (Fig. 15) is found to increase the selectivity towards cyclohexene (the desired product). The intermediate hydrogenation product cyclohexene distributes preferentially into the organic phase (Struijk, 1992) and is thus prevented from further hydrogenation to cyclohexane. Asahi Chemicals has utilized this concept in their commercial process. 4.3.4. Membrane reactors Membrane reactors may well be the future for many equilibrium-limited reactions requiring removal of small molecules such as hydrogen or water. The basic concepts involved are illustrated in Fig. 16 for the production of
Design of a reactor is often dictated by catalyst deactivation kinetics and the time interval between successive regenerations. If this time interval is on the order of 1 year, "xed-bed operation is preferred. If the time interval is on the order of 1 week, swing-type operation using two beds is usually selected, as in naphtha reforming. If the time interval between successive regenerations is on the order of a day to a few hours, then moving-bed operations are considered. If the time is less than an hour, circulating #uid-bed operation is called for. 5.1. Circulating yuid-bed reactor At the heart of the re"nery is the #uid catalytic cracker (FCC). The unit is divided into two main sections: the riser and the regenerator (Fig. 17). The riser is the reactor where the feed oil is injected with #owing catalyst and cracked. At the end of the riser is a stripper and disengager which separates catalyst and hydrocarbons, the latter going on to the fractionator. In FCC operations, coke formation on the catalyst is extremely fast. The catalyst is usually onstream for a few seconds before it deactivates. Catalytic cracking is an endothermic reaction. The need to regenerate the catalyst rapidly, coupled with the transport of heat via solid particles, led to the development of the FCC circulating #uid-bed reactor (Fig. 17). The catalytic process takes place in the riser and the products and the catalyst are immediately separated: the catalyst particles being sent to the regenerator, where the coke is combusted and the exothermic heat of reaction carried back to the riser reactor. 5.2. Moving-bed reactor
Fig. 16. Conceptual membrane reactor for ethylene oxide.
Continuous reforming (Trambouze, van Landeghem, & Wanquier, 1984) permits higher operating severity by maintaining the high catalyst activity of near-fresh catalyst through regeneration cycles of a few days. A moving-bed reactor system (Fig. 18) has the advantage of maintaining production while the catalyst is removed or
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Fig. 17. Modern FCC unit.
Fig. 18. Industrial moving-bed reactors.
replaced. Catalyst particles pass by gravity through one or more reactors in a moving bed and are conveyed to a continuous regeneration zone. Continuous catalyst regeneration generally is e!ected by passing catalyst particles downwardly by gravity in a moving-bed mode (Fig. 18) through various treatment zones in a regeneration vessel. Although movement of catalyst through the zones is often designated as continuous, in practice it is semi-continuous in the sense that relatively small amounts of catalyst particles are transferred at closely
spaced points in time. For example, one batch per minute may be withdrawn from the bottom of a reaction zone and withdrawal may take one-half minute; e.g., catalyst particles #ow for one-half minute in the 1-min period. Since the inventory in the reaction and regeneration zones generally is large in relation to the batch size, the catalyst bed may be envisaged as moving continuously. Moving-bed reactors have also been introduced in the upgrading of residual oil facilities (Pegels & Wij!els, 1974) where catalyst replacement is required because
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of deactivation by metal deposition and coking of the catalyst.
A new approach (Jacoby, 1999) is the hydrocarbon trap which absorbs the HCs emitted during the cold start period, releasing them to the catalyst only when it is hot (Fig. 19).
6. Recent developments and opportunities 6.3. Ultra-low sulfur fuels technology The following examples illustrate that the area of process intensi"cation using novel catalysts integrated with advanced reactors is getting increased attention and is opening new opportunities. 6.1. Programmed release catalyst Wax deposition in sub-sea pipelines is a common problem in transporting crude oil. The primary challenge in clearing pipeline blockages due to wax deposits is in supplying heat to regions farther down the pipeline, which are more prone to wax deposition. A novel catalytic system (Fogler, 1999) has been developed where the catalyst is introduced with the feed mixture but the catalytic reaction is deliberately delayed by incorporating a physical barrier over the active sites. The catalyst is encapsulated by a polymeric wall, which slowly dissolves in the crude oil mix over a period of 2 h, thus successfully delaying the exothermic chemical reaction to downstream regions of the transportation pipelines. 6.2. Fast light-ow catalytic converters To meet demands for stricter pollutant standards, researchers are investigating ways of boosting catalytic converter e$ciencies. In this respect, cold startup of the car poses a signi"cant challenge * especially in achieving California's 0.04 gm/mile ULEV hydrocarbon levels standards. In order to completely oxidize hydrocarbon emissions during the "rst 50 s of the cold start, a catalytic converter in the typical under-#oor position needs to be active at temperatures less than 1003C. One approach is to heat the catalyst chemically or electrically to attain temperature required for hydrocarbon oxidation.
On May 1, 1999, the US EPA proposed regulations limiting the gasoline sulfur level to be an average of 30 ppm to become e!ective in the 2004}2006 time frame. Most US re"ners marketing gasoline will need to add new desulfurization schemes to meet this stringent requirements. Conventional hydrodesulfurization is carried out in co-current down #ow trickle-bed reactors. The desulfurization reaction is product inhibited. As the reaction progresses down the reactor, the reaction slows down considerably for two reasons: (1) the hydrogen partial pressure reduces as hydrogen is consumed and (2) the H S formed, inhibits the reaction rate. Deep desulfur ization in conventional trickle-bed reactors is thus extremely di$cult (Sutikno, 1999). However, if the trickle-bed reactor were to be run in counter-current #ow mode, the major part of the catalyst bed would operate in an H S lean mode thus allowing ultra-low sulfur levels to be achieved economically in the hydrocarbon liquid phase. One could conceive a multifunctional catalytic system that would allow counter #ow trickle-bed operation without a high *P penalty. Ring-shaped catalysts have been used in some desulfurization processes such as SynSat, but the bed depths are small as ring-shaped catalysts cannot take much weight. Structured catalysts such as Sulzer's Katapak-S or M-Series or the composite structured packing described above may be worthwhile investigation targets. HDS performs well in removing 80}90% of sulfurbearing species in distillates, but increasingly poor at removing the remaining substantial polyaromatic sulfur-bearing compounds. The latter species are most susceptible to catalytic biodesulfurization (Pacheco et al., 1999). This clearly suggests the application of both hydrodesulfurization and biodesulfurization hybrids to achieve ultra-low sulfur levels in fuels. 6.4. Microtechnology
Fig. 19. Zeolite trap for cold start emission reduction.
Microreactors' narrow channels and thin walls (Fig. 20) make them especially e$cient for processes (Shanley, 1997) that are mass- and heat-transfer limited. In addition, residence times become strikingly short in microreactors. The laminar #ow in microreactor channels is yet another bene"t, leading to lower pressure drops than traditional reactors. While CPI applications of microreactors may be a decade away, there are niche areas where microreactors easily outperform traditional reactive systems. In this respect, recent studies have
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Fig. 20. Microreactor technology.
Fig. 21. Microengineered catalyst technology.
stressed the role and potential of fuel cells as an electric power source to run the motors of hybrid vehicles. However, the current infrastructure for transportation fuels does not have the capability to deliver the hydrogen used in fuel cells directly to such hybrid vehicles. For the near-term viability of such systems, technologies are necessary to convert gasoline, natural gas, and other hydrocarbon fuels, which can be handled by the current distribution system, to hydrogen onboard the vehicle. (Because of its low density, hydrogen requires extensive storage space in its pure form.) Microfabricated reformers could be developed to meet this need: compact size would make them ideal for cars. With heat transfer coe$cients that are an order-of-magnitude greater
than current heat exchangers, these multifunctional compact miniature reactors could become the enabling element for successful employment of hydrogen-based automobiles. 6.5. Microengineered catalyst systems A team of ABB chemical engineers working with university scientists has developed a novel catalytic material (Bayegan, 1999) that improves reactor and catalyst performance by e!ectively utilizing extremely small catalyst particles in "xed bed reactors. The new system, called micro-engineered catalyst (MEC), is based on mechanically robust, highly porous structures made of
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micro"bers in which small (nm) catalyst particles are entrapped (Fig. 21). The new material with the microengineered catalyst can be formed into a variety of shapes and structures, hydrodynamically tailored for various applications. In this way, excellent mass transfer, improved heat transfer and optimal pressure drop are achieved. Using computational #uid dynamics (CFD) and experimental modeling, various novel designs for reactor internals have been developed. Furthermore, the development of a new electrophoretic coating procedure enables the coating of each "ber within three-dimensional reactor structures. So far, signi"cant progress has been achieved toward a variety of industrial applications. Commercialization e!orts are currently underway to bring MEC to market.
7. Conclusions and recommendations Through the development of highly active and selective catalysts, feedstock utilization can be optimized which leads to a signi"cant reduction in energy consumption per unit weight of product and thus to a reduction in the build-up of greenhouse gas emissions. By employing multifunctional reactors, energy reduction is achieved by integrating several process steps in one vessel, hence reducing the overall utility requirements of the process. By in situ removal of products from the reaction zone, multifunctional reactors can reduce by-product generation from a process by lowering the rates of multiple or consecutive reactions. Waste generation in the chemical industry and the environmental burden caused as a result is mitigated. PI through use of multifunctional reactors permits signi"cant reductions in both investment and plant operating costs by optimizing the process. In an era of emaciated pro"t margins, it allows chemical producers more leverage in competing in the global marketplace. Focussing on the integration of new catalyst developments and advances in reaction engineering may lead to the development of additional or improved multifunctional reactors. These can assist in overcoming some of the challenges to the process industries in the future. By systematic thinking about the type of multifunctionality that is required, it will be easier to propose viable solutions. As illustrated in this paper, each type of multifunctionality requires `out-of-the-boxa thinking and thus R&D teams with a larger plurality in basic disciplines.
Acknowledgements The author wishes to thank Mr. Mitrajit Mukherjee for his assistance in creating the basic outline of this paper with its many illustrations. Dr. Nelleke van der Puil, Dr. Rudolf Overbeek and Mr. Robert Trubac are
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