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Ind. Eng. Chem. Res. 2004, 43, 5474-5482
Features of Fluidized Catalyst Beds for Proper Design and Operation of Catalytic Reactions Takami Kai and Toshio Tsutsui Department of Applied Chemistry and Chemical Engineering, Kagoshima University, Kagoshima 890-0065, Japan
Shintaro Furusaki* Department of Applied Life Science, Sojo University, Kumamoto 860-0082, Japan
Fluidized catalyst beds have been used in chemical and refinery processes and studied extensively for more than 50 years. However, the features of this reactor are still not understood perfectly. There is still much more to learn about the fluidized catalyst beds. It is important to elucidate the bubble behavior and the contact between catalyst particles and reactant components in the reactors. Understanding the features of the reactors contributes to the proper catalyst design together with the appropriate design and operation of the reactors. In this paper, the influence of gas properties on the bubble behavior is described, and pseudo-3-dimensional images of bubbles are shown. The important effects of direct and successive contact in the bed are discussed regarding conversion and selectivity. The reaction analysis showed that the selectivity could be improved by the effective use of the direct contact catalyst particles. The synergy effect of the three factors, i.e., catalyst activity, contacting feature between gas and particles, and transport/ mixing properties in the bed, plays a salient effect in raising selectivity of consecutive reactions. 1. Introduction Fluidized catalyst beds (FCBs) have been applied widely to reactions with a large thermal load and/or with regeneration of the catalyst, such as fluidized catalytic cracking (FCC), catalyst regeneration, partial oxidations to form acrylonitrile, ethylene dichloride, maleic anhydride, etc. Recently, some processes have chosen riser-type reactors for catalytic reactions such as the FCC process. Contractor1 summarized Dupont’s new circulating fluidized bed technology for maleic anhydride from its conception to commercialization. On the other hand, Sasol changed the Fischer-Tropsch synthesis reactor from the riser-type Synthol reactor to a dense-bed type fluidized bed called the Sasol Advanced Synthol.2 A dense-bed-type reactor is not an old technology and still has advantages for some reaction systems. Applications of FCBs have been put into practice by a number of researchers since early in the past century, and many generalized models have been proposed. However, problems are still unresolved, and the design of FCBs is considered to be somehow within the world of “art”. There is still much more to learn about fluidized bed technology,3 because phenomena that are not perfectly elucidated remain, and it is difficult to predict the problems intrinsic to the reaction system. Reactor models have been proposed by many researchers even recently. Although the information on bubbles is one of the important parameters to predict the performance of FCBs, there are some factors that should be considered in a model. Miyauchi et al.4 reviewed the salient character and behavior of FCBs extensively with * Corresponding author. Tel.: +81-96-326-3111 extension 5223. Fax: +81-96-323-1331. E-mail: furusaki@ life.sojo-u.ac.jp.
regard to transport phenomena and reactions. In the review article, they claimed the importance of the role of the catalyst itself on the extent of reaction in FCBs, i.e., the direct contact between gas and catalyst particles, particularly considering the role of the dilute phase (free board region). There were several significant developments after the review article. Some important topics are, for example, the behavior of FCBs accompanied by the gas volume decrease due to reaction and the role of the adsorption of reactants on the catalyst particles (the capacitance effect). On the basis of the above consequences, it can be said that we must think about the proper catalyst (content) design together with the appropriate design of the fluidized bed (vessel). Here we describe the above important features in designing and operating FCBs. 2. Bubble Behavior in FCBs 2.1. Features of Particles Used in FCBs. The physical properties of catalyst particles are important factors to establish good fluidization. Many researchers have mentioned that fluidized beds of fine particles classified in group A powders5 are used for catalytic reactions. Although Geldart’s classification is a useful guide when air is used as the fluidizing gas at ambient conditions, the borderlines in this map should be compensated according to the operating conditions. Newton et al.3 claimed the same point. The physical properties of the particles used for commercial FCBs are limited to a narrow region. Ikeda6 investigated the particle properties used in industrial processes and empirically classified these particles as A′ powders.4 This criterion has been a useful guide for many Japanese companies to design catalysts properly. Figure 1 indicates this criterion for good fluidization here as group AA (advantageous aeration) powders.
10.1021/ie030761l CCC: $27.50 © 2004 American Chemical Society Published on Web 04/14/2004
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Figure 1. Classification of group AA powders in Geldart map.
Figure 2. Effect of types of gas on bubble size (dp ) 55 µm, Fp ) 765 kg m-3).13
2.2. Bubble Behavior of FCBs. The mass transfer rate between bubble and emulsion phases influences the overall reaction rate in FCB reactors. It is therefore important to obtain an accurate value of the bubble surface area for the design of the reactors when a reactor model based on bubbles is employed. To predict the bubble size, many correlations have been proposed on the basis of the measurements with air as the fluidizing gas at ambient temperature. The fluidity of group A powders, especially group AA powders, is sensitive to gas properties, and the fluidity is affected significantly by temperature. The homogeneous bed expansion was affected by temperature.7 The effects of temperature involved the effect of hydrodynamic and interparticle forces. Xie and Geldart8 studied the effect of types of gases and temperature on the bed behavior in the particulate regime. The emulsion-phase expansion in the bubbling regime is also affected by gas properties. The emulsion phase voidage of bubbling beds increased with increasing gas viscosity, gas density, and temperature.9,10 Bubble size decreased with increasing temperature for group AA powders.11 The types of fluidizing gases also affected the bubble behavior.12,13 Figure 2 shows the change in bubble size with gas velocity when the bed was fluidized by air and hydrogen at ambient temperature. When the bed was fluidized by hydrogen, bubble size was larger than that in the bed fluidized by air. Under ambient pressure, the difference in bubble size is caused mainly by the expansion of the emulsion phase due to an increase in gas viscosity.14 When the voidage of the emulsion phase becomes large, the bubble size
becomes small. This phenomenon can be discussed on the basis of apparent bed viscosity.14,16 Kono et al.17 reported that a good correlation was obtained between the quasisolid viscosity and mean bubble size in a fluidized bed at elevated temperature (293.2-723.2 K). As described above, because the bubble behavior is affected by gas properties, it is difficult to estimate the mean bubble size for group AA powders using bubble models based on the measurements at ambient temperature. The reactor model based on bubbles relies on the premise that spherical bubbles accompanied by wakes ascend in the beds. When good fluidization is established in the fluidized bed of group AA powders, the bubble size is strongly influenced by the phenomenon of bubble splitting.4,18 In addition, bubbles ascend in the bed while deforming their shapes, and they are no longer spherical. It is impossible to observe directly the phenomena occurring inside a bed even if transparent materials are used for the apparatus. Therefore, the information on bubble behavior has been estimated on the basis of the data obtained using probes immersed in the bed such as a hot-wire probe, an optical probe, or a capacitance probe. In these cases, bubble behavior was influenced by the probe head geometry.19 2.3. Direct Measurement of Bubbles. On the other hand, to obtain information inside the bed, some researchers have used a different method that does not require immersing a probe in the bed. The method utilizes permeable electromagnetic waves such as X-rays and γ-rays, and electrical capacitance. Simons20 and Yates21 reviewed the noninvasive measurement of voidage in gas fluidized beds. Because the hydrodynamic behavior is not disturbed by the measurement, this method provides useful information about the voidage and bubbles. In the last 20 years, the technology of computerized tomography has been developed remarkably and has been applied to the measurement of fluidized beds in the past decade. Images were acquired at high frame rates using an electrical capacitance tomography (ECT) system. High frequency allows the real time imaging of the cross section of the measuring plane.22 Generally, the measurement frequency of ECT can be increased up to 100 Hz.23 Wang et al.24,25 measured the solid concentration distribution in a fluidized bed using an ECT system. In their case, however, the spatial resolution of the imaged cross-section was 32 pixels × 32 pixels, which corresponded to a pixel area of 22 mm2. It is difficult to enhance the spatial resolution of the ECT system. This system has been used for the measurement of timeaveraged particle concentrations over the cross-section. On the other hand, the spatial resolution of X-ray tomography is very high. Kantzas26-28 obtained images of the density and holdup using a modified medical scanner. Although the spatial resolution of the system was high (0.16 mm2), it took 2 or 3 s to scan one slice of the column; therefore, the image obtained was the timeaveraged density of a cross section. 2.4. Measurement Using a Fast X-ray CT Scanner. Measuring the shape of a moving object requires not only spatial resolution but also temporal resolution. Misawa et al.29,30 developed a fast X-ray CT scanner to measure the void fraction in a gas-liquid two-phase flow. They examined the limitations of the spatial resolution and the object velocity with spherical and
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Figure 4. Pseudo-vertical section image constructed from horizontal section images obtained from a CT scanner (dp ) 54 µm, Fp ) 1610 kg m-3).33
Figure 3. Pseudo-3-dimensional expression of an ascending bubble in an FCB (dp ) 36 µm, Fp ) 980 kg m-3).32
cylindrical acrylic phantoms of known dimensions.31 It was found that the responsibility and resolution of the system were sufficient to observe the shape of an object when the velocity was smaller than 1 m s-1. Kai et al.32 applied this system to the visualization of bubbles in a fluidized bed. Using an appropriate threshold value, the reconstructed images were binarized to quantify the bubble holdup distribution. Threedimensional images of a bubble could be obtained on the basis of these binarized images. Figure 3 shows pseudo-3-dimensional images of bubbles. The height of the measured section was 0.6 m from the distributor. The vertical axis in this figure is the elapsed time. If the shape and the ascending velocity of the bubbles do not change with time, the time axis is equivalent to the spatial coordinate. However, this assumption is not always true for bubble flow in a fluidized bed. Therefore, the bubble shape shown in Figure 3 is a pseudo-3dimensional image. This image shows that the bubble shape was not a spherical cap. The surface area was 3 times larger than that obtained from the calculation on the basis of the assumption that the bubble shape was a spherical cap.33 In addition, the image shown in Figure 3 implies that the main bubble was composed of some smaller bubbles. Figure 4 shows a pseudo-vertical-section image constructed from horizontal section images obtained using the CT scanner.33 This figure also shows that the main bubble was composed of some smaller bubbles. In a 2-dimensional fluidized bed, Kai et al.34 observed the bubble behavior and the particle movement around bubbles by the photocromic dye activation method. They described that the ascending bubble in FCBs suffered from sheer stress by the emulsion phase and that the bubble split into several smaller parts. The split parts then merged with the main bubble at the bottom of the main bubble. The bubbles in FCBs ascend in the bed while splitting and coalescing as shown in Figure 5.34 As shown in these figures, there were many particles in a bubble cluster. We should consider the role of these particles in a reactor model to predict the conversion
Figure 5. Bubble splitting and coalescence observed in a 2-dimensional FCB.34
and selectivity correctly. This is one of the causes of direct contact between bubble-gas and particles. 3. Successive Contact Mechanism 3.1. General Behavior of FCBs. Most FCBs are operated for oxidation and often in the range of detonative combustion. The stability in this case is affected by smooth fluidity, because a reactive gas must be in contact with the catalyst particles within very short intervals. According to this requirement, slugging should be avoided, and the particles used in FCBs are group AA powders (Figure 1). Particularly, the content of fines under 44 µm is said to be crucial.4,6 The existence of fines guarantees good fluidization and stable operation. It is well-known that gases flow as bubbles in fluidized beds, and this fact is one of the main reasons for the poor contact efficiency. However, the fines in FCBs provide the chance of direct contact. Lewis et al.35 first analyzed the good contact in their experiment on ethylene hydrogenation with a Ni-impregnated porous fine fluidized catalyst. The contact efficiency increased with increasing gas velocity when the catalytic activity was high, although it was not influenced by gas velocity when the activity was low. Similar reaction behavior was also found in the data of van Swaay and Zuider-
Ind. Eng. Chem. Res., Vol. 43, No. 18, 2004 5477
Figure 7. Axial distribution of local mass transfer term kobab and fraction of direct contact catalyst.4
toward the top of the bed. The reactant concentration is considered to be uniform laterally.
-UG dc/dz - efbkrc ) 0 Figure 6. General flow features of the FCB.4
weg.36 This characteristic of the FCBs is referred to as the direct contact mechanism that will be explained below. 3.2. Direct Contact Mechanism. It is well-known that gas flows through fluidized beds in the form of bubbles. However, there is a certain fraction of particles which contact the fluidizing gas directly as written above. In modeling this mechanism, we apply a two-phase model; i.e., the gas flows in the form of bubbles, and a small portion of the gas flows through the emulsion phase, which occupies the space of the dense phase outside of the bubbles (Figure 6). There is a circulation flow in the emulsion phase, but under certain conditions the flow can be ignored. The effect of mixing in the emulsion phase will be discussed in section 4. If bubbles are considered to flow in a plug-type flow, the simplified governing equations to calculate the extent of reaction in the steady state are given as follows.4
-UG dcb/dz - kobab(cb - ce) - νkrcb ) 0
(1)
kobab(cb - ce) ) (e - ν)krce
(2)
Here, e is the volume fraction of the emulsion phase, and ν is that of the catalyst particles directly contacting with bubbles. The third term in eq 1 expresses the contribution of the directly contacted catalyst particles. The particles suspended in the bubbles and partly those in the cloud region of the bubbles are thought to contribute to the direct contact mechanism. Thus, the right side of eq 2 denotes the reaction by the catalyst particles in the emulsion phase. 3.3. Successive Contact Mechanism. In addition to the above contribution of the direct contact particles, there is a contribution from particles in the free board (Figure 6) proposed by Miyauchi.37,38 In the free board, gas directly contacts these particles, and the contact efficiency is quite high. If we consider simply that the contact efficiency is 100% and the gas flow is plug flow, the following equation can be applied. Here, the volume fraction of particles is expressed by efb, where efb obviously changes with the vertical position, decreasing
(3)
Thus, the reactant flowing out of the dense phase into the free board further reacts due to the existence of floating catalyst particles. A similar phenomenon occurs in the vicinity of the gas distributor (jetting zone). Here, the conversion can be calculated by the above direct contact model or by using eq 3. Accordingly, the three zones, i.e., the jetting zone, the bubbling zone, and the free board, compose the FCB. The first and third zones are the good contact regions. 3.4. Experimental Analysis. The apparent reaction rate of the Deacon reaction in a fluidized bed increased with temperature,39 which could not be explained by conventional bubbling bed models. The paper was the first elucidation to present the effect of direct contact. Later, most of the direct contact particles are attributed to the particles floating in the free board region. The axial distribution of reactivity due to successive contact was investigated and affirmed by Furusaki et al.40,41 Grace and De Lasa42,43 discuss a similar profile of reactivity particularly in the effect of the jetting (grid) zone and the free board. The values of the mass transfer coefficient and the amount of directly contacting particles are derived from the experimental data from ethylene hydrogenation,4 which are shown in Figure 7. The experimental analysis of the mass transfer coefficient and the amount of direct contact catalyst were given by Furusaki et al.44 According to our studies,44-46 the value of kobab increased with UG and varied from 0.2 to 0.8 s-1 when UG was changed from 0.1 to 0.6 m/s. The fraction of direct contacting particles, ν, was below 0.05 and was not affected by gas veolocity.46 The fraction of dilute phase particles was obtained from the static pressure distribution, and their values were above 0.25 when UG ) 0.6 m/s.45 These parameters were affected by not only gas velocity but also particle properties. Further experimental studies were conducted using methanation of CO and CO2 to observe the effect of volume change caused by reaction.47,48 This effect has been studied by some researchers recently49-51 as well. 3.5. Stability Analysis by Successive Contact Mechanism. In the free board region, the contact efficiency is high, but the heat capacity is low due to the small particle density. Therefore, a temperature rise can be observed in the case of fast and exothermic
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reactions, e.g., ethylene hydrogenation.44 In highly exothermic reactions such as the partial oxidation of hydrocarbons, this may cause drastic instability. The simulation of such cases is given by Miyauchi et al.4 4. Assessment of Selectivity in FCB Reactions 4.1. Significance of Selectivity in FCB Reactions. Most industrial catalytic reactions with FCBs exhibit a consecutive reaction nature, and the objective products are usually intermediate products. For example, acrylonitrile and maleic anhydride are the intermediate products in the ammoxidation of propylene and oxidation of butane, respectively. In these processes, the formation of CO and CO2 as the end product should be suppressed even at high conversion. In the phenolmethylation process developed by Asahi Chemical Corp.,52 both o-cresol as an intermediate product and 2,6-xylenol as an end product are the objective products, where the production ratio should be controlled to meet the demand. Thus, selectivity is very important for industrial reactions, but the control of selectivity is not an easy task because many factors concerning the selectivity are interrelated. Although an appropriate way of designing and controlling the FCB reaction has been sought for improving selectivity, very few studies have focused on the theoretical treatment of consecutive reactions in FCBs. 4.2. Extension of the Successive Contact Mechanism Considering Direct Contact to Consecutive Reactions. Recently, Tsutsui53 extended the successive contact mechanism for applying to consecutive reactions in order to discuss the selectivity. In his analysis, direct contact is considered as in the original model,4 but axial mixing of the emulsion gas was not neglected in order to cover a wide range of catalyst activity and operation conditions. By converting the basic equations of the model into dimensionless forms, the basic parameters dominating the FCB reaction were clarified as follows
Nrr )
krA kobab UGLf ν , Nmr ) , PeB ) , e, , k aθ krM ekrA eEG e ob b f
where Nrr is the ratio of reaction rate constants in a consecutive reaction, Nmr is the ratio of the mass transfer capacitance coefficient to the reaction rate constant, e and ν/e are parameters indicating the direct contact ratio, and kobabθf is the dimensionless contact time. This means that the rate constant ratio between different reaction steps, Nrr, between different rate processes in a fluidized bed, Nmr, and the ratio of direct contact particles dominate the FCB reactions. When these parameters are equal even for different reaction systems, the reaction results are similar. This is the similarity rule in FCB reactions. The analytical solutions of the consecutive reaction model were derived for two cases; one is the VUME (vertically unmixed emulsion) case, neglecting the effect of axial dispersion, and the other the PME (perfectly mixed emulsion) case, as Lewis et al.35 used in the analysis of a single reaction in a fluidized bed. By applying these simplifications, simple solutions were obtained for more complex consecutive reactions. Con-
sequently, more detailed analysis and interpretation were made possible. According to the successive contact mechanism for a consecutive reaction, the yield of intermediate product is expressed by the following equations.53 At the outlet of a dense bed, for the VUME case
k/rM
k/rA + νkrA
(e - ν)krM
CbM1 k/rM / ) ‚ ‚{e-mkrMθf CbA0 (e - ν)krM k/ + νk - mk/ rA rA rM CbM0 -mk/rM / e-(krA + νkrA)θf} + ‚e θf (4) CbA0 where
1 1 1 1 1 1 ) + , ) + , k/rA kobab (e - ν)krA k/rM kobab (e - ν)krM krM 1 + (5) m)1+ν e - ν kobab
(
)
and for the PME case
CbM1 ) {φAφM(e - ν) + CbA0
(
νR}‚
1+
NrA AφANrA
(
)(
1+
MφMNrM
)
+
)
1 - e-φANb 1 - e-φMNb CbM0 -φMNb 1 - e-φMNb ‚e + (6) MφMNrM CbA0 1+ 1 - e-φMNb
where
Nb ) kobabθf, NrA ) krAθf, NrM ) krMθf, krA krM ν, φM ) 1 + ν, φA ) 1 + kobab kobab NrANrM (e-φMNb - e-φANb) , R ) R1 + R2 Nb(1 - e-φMNb) NrA NrM φAφM + M , R2 ) AM , R1 ) 1 + A Nb Nb φA - φM (e - ν)krA (e - ν)krM A ) e + ν, M ) e + ν (7) kobab kobab
(
)
At the outlet of the reactor, i.e., at the top of the freeboard, for both VUME and PME cases
krA CbM1 -eekrMθf CM2 ) (e-eekrMθf - e-eekrAθf) + e CbA1 krA - krM CbA1 (8) With this extended successive contact model, the selectivity of consecutive reactions and the reactor performance in FCBs were analyzed and discussed by Tsutsui53 as follows. 4.3. Analysis of Selectivity by the Consecutive Reaction Model Based on the Successive Contact Mechanism. To verify the consecutive reaction model based on the successive contact mechanism, hydrodealkylation of alkylnaphthalene was conducted, where triand di-methylnaphthalene were consecutively demethy-
Ind. Eng. Chem. Res., Vol. 43, No. 18, 2004 5479
Figure 8. Change in concentration of raw material, intermediate product, and end product with contact time in an FCB hydrodealkylation of alkylnaphthalene and comparison with the models (Nmr ) 1.78, Nrr ) 3.65, e ) 0, ν/e ) 0).
Figure 10. Effect of the ratio of direct contact particles in the direct contact zone to whole particles in the reactor, Ndcf, on the yield of the intermediate product.
Figure 11. Effect of the ratio of direct contact particles in the direct contact zone to whole particles in the reactor, Ndcf, on the selectivity of the intermediate product.
Figure 9. Effect of contact time on apparent reaction constant in an FCB hydrodealkylation and comparison with the models.
lated through methylnaphthalene to naphthalene.54 Figure 8 shows the observed change in concentration of each component with contact time and the simulated results by the model. The concentration of naphthalenes was not zero at θq ) 0, because the feed oil contained these components. Good agreement can be seen between the experimental results and the model, particularly the PME model, indicating that back-mixing should be considered in the reaction analysis. Figure 9 shows a relationship between kORq and contact time in the hydrodealkylation. The parameter kORq is the apparent reaction rate constant based on the settled bed height. Also here, the PME model could explain the decreasing tendency of kORq with increasing contact time. Ikeda55 reported such a decreasing tendency for propylene ammoxidation in an FCB reactor. Figure 9 also indicates that a mass-transfer controlling region and a back-mixing controlling region exist depending on the value of kobabθf. For kobabθf < 3, FCB reactions are in the mass transfer controlling region even if the PME model is applicable. This condition is often seen in the case of reactions with a very active catalyst and a contact time of several seconds. For kobabθf > 3, the reaction is in the back-mixing controlling region, where the apparent reaction rate constant decreases with the increase in the contact time. In the back-mixing controlling region, both the actual reaction rate constant krA and the mass transfer capacitance
coefficient kobab can be estimated by comparing the reaction data with the PME model. In the case of the results shown in Figure 9, kobab is estimated to be about 0.3 s-1, and this value is in good agreement with the data reported by Nozaki et al.,45 Furusaki et al.,44 and Kai et al.46 Figures 10 and 11, respectively, show the effect of the direct contact ratio Ndcf on the yield and selectivity of the intermediate product (M) for the case of Nmr ) 0.1, Nrr ) 40, and ν ) 0. The remarkable effect of the direct contact on the yield and selectivity of the intermediate product can be understood. To increase only the conversion, direct contact is not always necessary. It can be increased by another method, for example, by increasing the reaction temperature when Nmr is high enough. However, selectivity is hardly increased by raising the temperature, because usually both krA and krM are increased simultaneously. Therefore, utilization of direct contact is indispensable for improving selectivity. Recently, a new process of methanol ammoxidation was developed by the use of a very active catalyst.56 It is reported that the selectivity and yield of the intermediate product in this process were remarkably improved by enhancing the direct contact in the reactor.57 This fact is in good agreement with the reaction analysis by this model. 4.4. Design and Performance of FCB Reaction. Figure 12 shows the cooperative effect of the basic parameters on the maximum yield of the intermediate product comprehensively. When there is no contribution of direct contact, the ratio of the mass transfer capacitance coefficient to the reaction rate constant Nmr is important. When the reaction rate constant is increased
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son. The PeB number is changed from 0.001 to 1000, i.e., from nearly perfect mixing to almost a plug flow case. The PeB number in a riser reactor, which is thought to be a plug-flow-type reactor, is usually in the range 1-4. From Figure 12, it may be understood that FCB reactors with a dense bed can show nearly the same selectivity as riser reactors if the synergy effect of the basic parameters is fully utilized. Since mass transfer resistance exists also in riser reactors, higher selectivity is obtainable in FCB reactors. In addition, the dense bed in FCB reactors has another advantageous nature, the homogenizing effect of temperature. It can be said that sufficient analysis and reaction design including both catalyst design and fluidized state design for attaining the best combination of the basic parameters are indispensable for a high performance FCB reactor. Conclusion
Figure 12. Cooperative effect of the basic parameters on the maximum yield of the intermediate product in FCB reactions and comparison with that in an ideally dispersed reactor.
but the mass transfer capacitance coefficient is not increased as much, the selectivity is shown to decrease conversely as is sometimes seen in industrial processes. Introduction of an active catalyst is not successful in this case. On the contrary, when the direct contact ratio is large, the limitation of mass transfer does not arise, and high selectivity can be obtained. In the case of using a catalyst with a higher ratio of the consecutive reaction rate constant Nrr, selectivity becomes higher at any level of the direct contact ratio. Thus, the combination of the development of a selective catalyst with a higher Nrr and the attainment of a higher direct contact ratio Ndc or a higher mass transfer to reaction rate ratio Nmr is necessary for obtaining the highest selectivity in a fluidized bed reaction. In the case of using a highly active catalyst with a reaction rate constant of 1-10 s-1 such as an oxidation catalyst, the enhancement of the direct contact ratio should be essential because Nmr is small. On the other hand, in the case of using catalyst with a moderate activity, it is also effective to enhance the mass transfer capacitance coefficient to attain an Nmr larger than 1. In this case, suppression of back-mixing of the emulsion gas contributes to a further increase in selectivity. As shown in Figure 12, the maximum yield of the intermediate product in an FCB reactor varies very widely depending on the basic parameter values. The combination of higher values of Ndc, Nmr, and Nrr along with suppression of back-mixing brings about the maximum synergy effect on the selectivity. Thus, the synergy effect of the following three factors, i.e., the catalyst performance, the contact state between particles and gas, and the transport and mixing properties in the fluidized bed, is quite important in designing and establishing a selective process. Depending on this synergy effect, it is expected that the selectivity in FCB reaction processes will be increased drastically. 4.5. Evaluation of the FCB Reactor as a Fluidized Contacting Method. In Figure 12, the maximum yields of intermediate products, obtainable in an assumed reactor with ideally dispersed particles where no mass transfer resistance exists, are shown for compari-
Important features affecting the property of the FCB are described in this review. First, recent developments in basic researches and applications of the FCB are overviewed. Bubble behavior in the FCB is then presented with recent developments in sensing techniques. The importance of the AA flow regime in the FCB is reconsidered, and the fluidity is related to gas properties. In this discussion, the successive contact mechanism, i.e., the importance of free board, is introduced with consideration of direct contact between the gas and the particles. This direct contact is attributed to the contact in the free board and in the bubbles. Reaction data verified the model analysis. Finally, selectivity of the reaction in FCB is discussed extensively. It is shown that the direct contact plays an important role in raising the selectivity. A well-designed FCB is expected to have the same selectivity as that in riser reactors, and in order to realize high contact efficiency and high selectivity, we can say that appropriate designs of both the FCB and the catalyst particles are essential. Nomenclature C ) molar concentration in free board, mol m-3 Cb ) concentration in bubble phase, mol m-3 Ce ) concentration in emulsion phase, mol m-3 db ) bubble size, m dp) average particle size, m EG ) axial dispersion coefficient of gas in emulsion phase, m2 s-1 e ) amount of particles in direct contact zone/amount of particles in dense bed zone kobab ) overall mass transfer capacitance coefficient between bubble and emulsion phase per unit volume of bed, s-1 kORq ) apparent reaction rate constant in FCB reaction based on contact time θq, s-1 kr ) rate constant for first-order irreversible reaction based on the emulsion density, s-1 Lf ) height of the fluidized dense bed zone, m Lq ) height of a settled bed, m Ndc ) ratio of direct contact particles to whole particles in reactor Ndcf ) ratio of direct contact particles in the direct contact zone to whole particles in the reactor (e/(1+e)) Nmr ) nondimensional number of mass transfer and reaction rate ratio Nrr ) ratio of reaction rate constants in a consecutive reaction (krA/krM)
Ind. Eng. Chem. Res., Vol. 43, No. 18, 2004 5481 PeB ) Peclet-Bodenstein number (UGLf /eEG) S ) selectivity UG ) superficial gas velocity, m s-1 Y ) yield Y* ) maximum yield z ) axial coordinate, m Greek Symbols e ) volume fraction of emulsion including direct contact particles efb ) volume fraction occupied by particles in the free board θf ) contact time in the fluidized dense bed (Lf /UG), s θq ) contact time defined as Lq/UG, s θR ) contact time in FCB reactor defined as (1+e) Lf/UG, s ν ) volume fraction of direct contact particles expressed by the equivalent density of emulsion Fp ) particle density, kg m-3 Subscripts A ) feed component in a consecutive reaction M ) intermediate component in a consecutive reaction 0 ) inlet of the dense bed zone (inlet of the fluidized bed reactor) 1 ) outlet of the dense bed zone (inlet of the direct contact zone) 2 ) outlet of the direct contact zone (outlet of the fluidized bed reactor)
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Received for review October 10, 2003 Revised manuscript received January 20, 2004 Accepted January 22, 2004 IE030761L
