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Experimental study on bubble behavior in gas–liquid–solid three-phase circulating fluidized beds
Powder Technology 137 (2003) 83 – 90 www.elsevier.com/locate/powtec
Experimental study on bubble behavior in gas–liquid–solid three-phase circulating fluidized beds Tiefeng Wang, Jinfu Wang*, Weiguo Yang, Yong Jin Department of Chemical Engineering, Tsinghua University, Beijing 100084, PR China
Abstract The gas – liquid – solid three-phase circulating fluidized bed (TPCFB) is a novel multiphase reactor with outstanding performance in hydrodynamics, heat transfer, mass transfer and catalyst treatment. The object of this work is to study the bubble behavior in TPCFBs for better understanding of the hydrodynamics of the reactor. The local gas holdup, bubble size distribution, and bubble rise velocity in different radial positions are measured using a fiber optic probe, and the effects of operating conditions on bubble rise velocity are also investigated. Moreover, the bubble rise velocity and its distribution are compared with the data from gas – liquid two-phase external loop airlift reactors and gas – liquid – solid three phase fluidized beds without external circulation of liquid and particles. D 2003 Elsevier B.V. All rights reserved. Keywords: Bubble rise velocity; Three-phase fluidized bed; Particle circulation; Fiber optic probe
1. Introduction Gas –liquid – solid three-phase fluidized beds have been widely used in petrochemical, metallurgical, environmental and coal liquefaction processes [1]. Conventional threephase fluidized beds without external particle circulation are problematic when applied to systems with a catalyst liable to be deactivated or with small/light particles, moreover, the heat removal is a troublesome problem for a strong exothermic reaction. Based on the flow regime map of the gas – liquid – solid co-current three-phase fluidization proposed by Muroyama and Fan [2], Liang et al. [3,4] presented a new flow regime map taking into account the circulating fluidization regime operated in a high liquid velocity range. The particle entrainment leads to external particle circulation and the particle circulating rate can be controlled to regulate the solid holdup in the reactor. This feature is of great advantage for catalyst regeneration, reactor design and operation. Previous studies on hydrodynamics and mass transfer show that three-phase circulating fluidized beds (TPCFBs) have some better characteristics than the conventional three-phase fluidized beds [3– 8]. Liang et al. [4] investigated the axial phase holdup variation from a conventional fluidization regime to trans* Corresponding author. Tel.: +86-10-6278-5464; fax: +86-10-62772051. E-mail address: [email protected] (J. Wang). 0032-5910/$ - see front matter D 2003 Elsevier B.V. All rights reserved. doi:10.1016/j.powtec.2003.08.032
port regime. The axial profile of phase holdups in TPCFBs are rather flat, which is different from the convention threephase fluidized bed where a clear distinction exists between the bulk fluidized bed region and the freeboard region, and the solid holdup in the freeboard region is nearly zero. Han et al. [9] investigated the effects of the operating parameters on the axial and radial dispersion coefficients of the liquid phase and concluded that TPCFBs approach the ideal plug-flow reactor more closely compared with conventional threephase fluidized beds. Yang et al. [6] studied the radial distribution of liquid velocity using the electrolyte tracer method and found that liquid velocity distribution is nonuniform in the radial direction but more even than that in conventional three-phase fluidized beds. These results show that the hydrodynamic behavior in TPCFBs is different from that in convention three-phase fluidized beds; therefore further investigations are still needed. The bubble behavior, which is related to the profile of the phase holdup, the interaction between phases, and the mass transfer, have been important aspects for multiphase reactors. Remarkable efforts have been devoted to bubble behavior, such as the gas holdup, bubble size and bubble rise velocity in three-phase fluidized beds without external particle circulation [10 – 12], while very limited work on bubble behavior in TPCFBs can be found in the literature. Wang et al. [5] studied the bubble size distribution in TPCFB, and the data of the bubble size distribution was used to obtain the mass transfer coefficient and gas – liquid interfacial area from the
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T. Wang et al. / Powder Technology 137 (2003) 83–90
beads of 0.4 mm diameter and density of 2460 kg/m3 are the solid phase. A detailed description was given in our previous study [5]. A fiber optic probe system was developed for measuring the bubble behavior in multiphase flow as shown in Fig. 2. The probe consists of two parallel communication optic fibers of 62.5 Am diameter with a distance 1.2 mm. The use of tenuous optic fiber improves the ability of detecting small bubbles (in the literature, an optic fiber of diameter 1 mm was used by several investigators [13,14]). The technique of emitting and receiving light in the same fiber makes the probe structure compact and ensures a quick signal response and slight disturbance to the flow field when measuring the bubble parameters. More than 1000 bubbles are sampled each measurement insuring that the effect of the sample size on the measured bubble chord length distribution is eliminated. A specially written program has been developed to control data sampling and to analyze the measured data [5].
3. Signal processing Fig. 1. Scheme of the experimental apparatus.
volumetric mass transfer coefficient. However, the bubble dynamic behavior has not been reported. This paper addresses the bubble rise velocity and its distribution in TPCFBs, which is important in determining the RTD of the gas phase and the diffusion between the gas and liquid phase.
2. Experimental A schematic diagram of the experimental apparatus is shown in Fig. 1. The riser is a vertical Plexiglas column, 140 mm in inner diameter and 3.0 m in height. Tap water and air are used as the liquid and gas phase, respectively. Glass
The bubble rise velocity can be calculated by cross correlating the two probe fiber signals. Typical signals are shown in Fig. 3. Due to the distance between the two probe fibers, the signal of the down-stream fiber is slightly delayed compared to that of the up-stream counterpart, which is clearer in the locally enlarged signal shown in Fig. 4. When a bubble rises through the probe, the up-stream fiber first contacts and pierces the upper surface of the bubble at time t1 and outputs a high level signal, and then the same course occurs to the down-stream fiber at t1V. The difference between t1Vand t1 is the time that a rising bubble takes to move from the up-stream fiber to the down-stream one. The bubble rise velocity can be calculated by dividing the distance between the two fibers, which is a given structure parameter, by t1V t1. Considering that the probe signals are not ideally rectangular, t1V t1 can be calculated by signal
Fig. 2. Hardware structure of the fiber optic probe for measuring bubbles.
T. Wang et al. / Powder Technology 137 (2003) 83–90
85
be calculated and the bubble size distribution can be further derived based on the algorithm proposed first by Liu and Clark [15] and improved by Wang et al. [5]. Thus the bubble size and rise velocity are obtained to characterize the bubble behavior.
4. Results and discussions 4.1. Distribution of ub
Fig. 3. Typical signals from the optical fiber probe.
cross correlation. The cross correlation function of the two tip signals V1(t) and V2(t) is as follows cðsÞ ¼
Z
tiþ1
V1 ðtÞV2 ðt sÞdt
ð1Þ
t¼ti
The value of s corresponding to the maximum of c(s) is the time delay t1V t1. Fig. 3 shows that two signals for most bubbles present high quality cross correlation, which makes it possible to obtain ub of each individual bubble. The distribution and average value of ub can be further determined by statistical analysis. With the determined bubble rise velocity, combined with the bubble signal duration time, the bubble chord length can
Fig. 4. Locally enlarged signal marked by the ellipse in Fig. 3.
Due to the bubble size distribution and flow turbulence, the bubble rise velocity is distributed, as shown in Fig. 5. Matsuura and Fan [10] studied the bubble behavior in a gas – liquid – solid fluidized bed with glass beads of 3 and 6 mm and a binary mixture of these particles. They found that in the dispersed bubble flow regime, the bubble rise velocity is narrowly distributed. In the coalesced bubble flow regime the bubble rise velocity distribution is similar to that in the dispersed bubble flow regime, except that the tail of the distribution in the coalesced bubble flow extends to much higher bubble rise velocities resulting in a larger bubble rise velocity. The experiments in this work have obtained similar results, as shown in Fig. 5. The TPCFB operates in the dispersed bubble flow regime and the bubble rise velocity distribution is narrow when the superficial gas velocity is low (Ug = 0.018 m/s), and operates in the coalesced bubble flow regime with a more spread distribution of the bubble rise velocity when the superficial gas velocity is higher (Ug = 0.054 m/s). 4.2. Radial profile of ub The radial profiles of the liquid velocity, gas holdup and bubble size all affect the radial profile of bubble rise
Fig. 5. Bubble rise velocity distribution at radial position r/R = 0 under different Ug (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s).
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Fig. 8. Effect of the superficial gas velocity on radial profile of the average bubble rise velocity (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137). Fig. 6. Bubble rise velocity distribution at different radial position (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s, Ug = 0.036 m/s, Gs = 10.8 kg/m2/s).
velocity. The distributions of ub in different radial positions shown in Fig. 6 indicate that ub in the central region has a larger average value and smaller standard deviation than that in the wall region. Bubble rise velocity is dependent on the liquid velocity, physical properties of the liquid – solid suspension, gas holdup and bubble size. Absolute bubble velocities (i.e. relative to the column wall) are a combination of the relative bubble velocity and the continuous phase velocity [16,17]. In TPCFBs, the bubble size has relatively narrow distribution and varies just slightly with the operating conditions within a certain range. The variation of the bubble size distribution with the superficial velocity is shown in Fig. 7. Under the experimental conditions (Ug = 0.018 – 0.063 m/s; Ul = 0.063 – 0.011 m/s), the bubble Sauter diameter varies in 4– 6 mm range [5]. In this range,
the rise velocity of a single bubble almost remains constant [18], so the influence of the bubble size distribution on the bubble rise velocity is negligible. The liquid velocity [6] and gas holdup [4] are higher in the central region, both favorable to increase the bubble rise velocity. The scatter distribution of ub in the near wall region implies that the distribution of ub is mainly caused by the turbulence of liquid phase and gas phase. Large vortexes in dynamic equilibrium of formation and disintegration were observed in the near wall region, causing more violent turbulence and more scattered distribution of ub. Fig. 8 shows the influence of the superficial gas velocity on the radial profile of average bubble rise velocity. The bubble rise velocity increases with the superficial gas velocity, and the maximum value is not present in the center line but at the position of r/R = 0.2, probably due to the spiral motion of bubble swarm. The radial profile of the bubble rise
Fig. 7. Effect of superficial gas velocity on bubble size distribution at radial position r/R = 0 (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s).
Fig. 9. Radial profile of the dimensionless bubble rise velocity ub/uba at different superficial gas velocity in TPCFBs under operating conditions shown in Table 1.
T. Wang et al. / Powder Technology 137 (2003) 83–90 Table 1 Operating conditions for Fig. 9 n
o
E
87
Table 2 Bubble slip velocity q
x
E
a
B
*
Ug (m/s) 0.018 0.027 0.036 0.045 0.054 0.036 0.036 0.036 0.036 Ul1 (m/s) 0.054 0.054 0.054 0.054 0.054 0.054 0.063 0.072 0.081 Ul2 (m/s) 0.018 0.018 0.018 0.018 0.018 0.018 0.018 0.018 0.018 Gs (kg/ 10.5 10.5 10.8 11.0 11.5 10.8 10.5 10.3 11.2 m2/s)
Ug 0.018
0.018
0.018
0.027
0.027
0.027
Ul1 Ul2 r/R ubs
0.054 0.018 0.029 0.403a
0.054 0.018 0.058 0.379a
0.054 0.018 0 0.373a
0.054 0.018 0.029 0.467a
0.054 – – – 0.018 – – – 0.058 – – – 0.437a 0.28b 0.31b 0.37b
0.054 0.018 0 0.355a a b
velocity in conventional three-phase fluidized bed by Lee and De Lasa [19] is also shown in Fig. 8, showing that the radial profile of the bubble rise velocity is more even in TPCFBs. Radial profiles of dimensionless bubble rise velocity ub/uba are plotted in Fig. 9, showing that the variation of liquid velocity, gas velocity and particle circulating rate have little effect on ub/uba (Table 1). The radial profile of ub/ uba can be described by the following equation:
0.029 0.047 0.068
Data from our experiment. Data from Marianne et al. [20].
understanding of the multiphase flow and it is especially useful for testing the validation of CFD models. 4.3. Influence of operating conditions
ð2Þ
In TPCFBs, the particle circulating rate Gs and solid holdup can be adjusted by regulating the flow ratio between the primary and secondary stream of the liquid, thus the superficial gas velocity Ug, superficial liquid velocity Ul and particle circulating rate Gs can be independently controlled.
Fig. 10 illustrates the radial profile of average bubble rise velocity in gas – liquid two-phase external loop reactors [20]. It can be seen that radial profile of ub in TPCFBs is slightly flatter compared to that in the external loop reactors. This result is reasonable considering the uniform radial profile of hydrodynamic parameters such as the gas holdup and liquid velocity [4,6,8] reported by previous investigators. The bubble slip velocity ubs can be computed from the bubble and liquid velocity data. The results shown in Table 2 indicate that the slip velocities may differ considerably from the value of 0.25 m/s often used in gas –liquid two-phase system and gas – liquid –solid three-phase system [20]. Local information about the bubble rise velocity allows a better
4.3.1. Influence of Ug on ub The gas holdup eg increases with increasing Ug, leading to increase in bubble rise velocity. The distributions of ub under different Ug are shown in Fig. 5. Both the standard deviation and the average value of the bubble rise velocity distribution increase with increasing Ug. The influence of Ug on average bubble rise velocity is shown in Fig. 11, indicating that with increase of Ug, ub in the near wall region increases more obviously than that in the central region, leading to more uniform radial profile of ub. The influence of Ug on the bubble rise velocity can be explained by the bubble wake effect. When the gas holdup increases with Ug, the bubble number within a bubble swarm also increases, resulting in a decrease of the bubble moving friction and increase of the
Fig. 10. Radial profile of the dimensionless bubble rise velocity ub/uba at different superficial gas velocity in gas – liquid external loop reactors (data from Marianne et al. [20]).
Fig. 11. Effect of superficial gas velocity on average bubble rise velocity at different radial positions (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/ m2/s, es = 0.137).
r r 2 r 3 ub ¼ 1:089 þ 0:416 þ0:5237 1:22 R R R uba
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bubble rise velocity. On the other hand, increasing Ug enhances the flow turbulence and bubble breakup and coalescence, causing more spread distribution of ub. The bubble rise velocity is influenced by the local gas holdup, bubble size and liquid velocity. The variation of the local gas holdup with Ug is also measured, and the results are shown in Fig. 12. The gas holdup in the near wall region does not increase as significantly as that in the central region. Considering that the bubble size has no significant influence on the bubble rise velocity, the liquid velocity in the near wall region should increase with an increase of Ug. This analysis is supported by the result of Yang et al. [6]. They reported that with an increase of Ug, the liquid velocity decreases in the central region and increases in the near wall region, resulting in a more even radial profile. 4.3.2. Influence of Ul on the bubble rise velocity The influence of Ul on ub is also experimentally investigated, and the results are shown in Figs. 13 and 14. It can be seen that Ul has no significant influence on ub, except a slight decrease of the proportion of bubbles with low velocity, which is reasonable since the superficial liquid velocity only varies in a relatively small range [3] ensuring that the reactor operates in the circulation regime. The bubble rise velocity is influenced by the gas holdup, local liquid velocity and solid holdup. Both the local liquid velocity and gas holdup increase with increasing U l , while the solid holdup decreases. The increase of the liquid velocity and gas holdup tends to increase the bubble rise velocity, while the decrease of solid holdup has the opposite influence. As a result, the bubble rise velocity has no remarkable variance with Ul, as shown in Fig. 14.
Fig. 13. Effect of superficial liquid velocity on bubble rise velocity distribution at radial position r/R = 0 (Ug = 0.036 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137).
4.3.3. Influence of Gs on the bubble rise velocity TPCFBs have the advantage to adjust Gs easily, which in turn influences es, by regulating the flow ratio between the
primary and secondary stream. With increase of es, the apparent viscosity and density of the suspension increase, causing larger buoyancy and bubble size, hence increase in the bubble rise velocity. The increase of es also increases the drag coefficient, tending to decrease ub, but this influence is more moderate than that mentioned above. As a result, the bubble rise velocity increases with Gs, and this variance becomes negligible when Gs exceeds a certain value, as shown in Fig. 15. Li and Prakash [21] studied the influence of slurry concentrations on bubble population and their rise velocities in a three-phase slurry bubble column (Ug = 0.10– 0.30 m/s; es = 0– 40% v/v), measuring the small and large bubble rise velocity using fast response pressure transducers and the dynamic gas disengagement technique. The rise velocities of both large and small bubbles increase with
Fig. 12. Radial profile of eg under different Ug. (Ul1 = 0.054 m/s Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137).
Fig. 14. Effect of superficial liquid velocity on average bubble rise velocity (Ug = 0.036 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137).
T. Wang et al. / Powder Technology 137 (2003) 83–90
R t Ug Ul1 Ul2 ub uba ubs u¯b V
89
radius of the fluidized bed (m) time (s) superficial gas velocity (m/s) superficial liquid velocity of the primary stream (m/s) superficial liquid velocity of the secondary stream (m/s) absolute bubble rise velocity (m/s) cross-sectional average bubble rise velocity (m/s) bubble slip velocity (m/s) local mean bubble rise velocity (m/s) output signal of the probe (V)
Greek symbols s time variable of c(s) (s)
Fig. 15. Effect of particle circulation rate on average bubble rise velocity (Ug = 0.036 m/s, Ul = 0.072 m/s).
increasing slurry concentrations up to slurry concentration of about 20% (v/v), consistent with our results.
5. Conclusions Experimental study on the bubble rise velocity with a fiber optic probe has resulted in the following conclusions. The bubble rise velocity has a radial profile with larger average value and smaller standard deviation in the central region than that in the near wall region. The radial profile of the bubble rise velocity is more even than that in conventional three-phase fluidized beds and gas – liquid external loop reactors in the literature. The average value and standard derivation of ub distribution increase with Ug. The superficial liquid velocity Ul has no significant influence on ub under the experimental conditions. The bubble rise velocity ub increases with es in low es range and this effect becomes negligible when es exceeds a certain value. The bubble slip velocity calculated from the local bubble rise velocity and liquid velocity is quite different from the value of 0.25 m/s often used in gas – liquid two-phase system and gas – liquid – solid three-phase system.
Notations c cross correlation function d distance between probe tips (m) db bubble diameter defined by the maximum vertical chord length (mm) Gs circulation rate of the solid particle (kg/m2/s) Ps(db) probability density function of the bubble size (1/mm) PDF probability density function of the bubble rise velocity (s/m) r radial position of the fluidized bed (m)
Subscripts g gas phase l liquid phase b bubble References [1] L.S. Fan, Gas – Liquid – Solid Fluidization Engineering, Butterworths Publishers, New York, 1989. [2] K. Muroyama, L.S. Fan, Fundamentals of gas – liquid – solid fluidization, AIChE J. 31 (1985) 1 – 34. [3] W.G. Liang, Q.W. Wu, Z.Q. Yu, Y. Jin, H.T. Bi, Flow regimes of the three-phase circulating fluidized bed, AIChE J. 41 (2) (1995) 267 – 271. [4] W.G. Liang, Z.Q. Yu, Y. Jin, Hydrodynamics of a gas – liquid – solid three phase circulating fluidized bed, Can. J. Chem. Eng. 73 (1995) 656 – 661. [5] T.F. Wang, W.G. Yang, Y. Jin, Experimental study on bubble size distribution in three-phase circulating fluidized beds, J. Chem. Ind. Eng., China 52 (3) (2001) 197 – 203. [6] W.G. Yang, J.F. Wang, J. Zhou, Y. Jin, Liquid phase flow structure and backmixing characteristics of gas – liquid – solid three-phase circulating fluidized bed, Chem. Eng. Sci. 54 (1999) 5293 – 5298. [7] W.G. Yang, J.F. Wang, L.M. Zhou, Y. Jin, Gas – liquid mass transfer behavior in three-phase CFB reactors, Chem. Eng. Sci. 54 (1999) 5523 – 5528. [8] T.F. Wang, J.F. Wang, W.G. Yang, Y. Jin, Experimental study on gasholdup and gas – liquid interfacial area in TPCFBs, Chem. Eng. Commun. 187 (2001) 251 – 263. [9] S.J. Han, J. Zhou, Y. Jin, K.C. Loh, Z.W. Wang, Liquid dispersion in gas – liquid – solid circulating fluidized beds, Chem. Eng. J. 70 (1) (1998) 9 – 14. [10] A. Matsuura, L.S. Fan, Distribution of bubble properties in a gas – liquid – solid fluidized bed, AIChE J. 30 (6) (1984) 894 – 903. [11] Z. Chen, C. Zheng, Y. Feng, H. Hofman, Local bubble behavior in three-phase fluidized beds, Can. J. Chem. Eng. 76 (1998) 315 – 318. [12] T. Miyahara, L.S. Fan, Properties of a large bubble in a bubble swarm in a three-phase fluidized bed, J. Chem. Eng. Jpn. 25 (1992) 278 – 382. [13] S. Saberi, K. Shakourzadeh, D. Bastoul, J. Militzer, Bubble size and velocity measurement in gas – liquid systems: application of fiber optic technique to pilot plant scale, Can. J. Chem. Eng. 70 (1995) 253 – 257. [14] A. Cartellier, Optical probes for local void fraction measurements: characterization of performance, Rev. Sci. Instrum. 61 (2) (1990) 874 – 886. [15] W. Liu, N.N. Clark, Relationships between distributions of chord lengths and distributions of bubble sizes including their statistical parameters, Int. J. Multiph. Flow 21 (1995) 1073 – 1089.
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[16] L. Sylvain, G. Christophe, Characterization of bubble column hydrodynamics with local measurements, Chem. Eng. Sci. 54 (1999) 4895 – 4902. [17] L.H. Caroline, L. Faical, G. Christophe, Understanding gas-phase hydrodynamics in bubble columns: a convective model based on kinetic theory, Chem. Eng. Sci. 52 (1997) 63 – 77. [18] K. Tsuchiya, A. Furumoto, L.S. Fan, J.P. Zhang, Suspension viscosity and bubble rise velocity in liquid – solid fluidized beds, Chem. Eng. Sci. 52 (1997) 3053 – 3066.
[19] S.L.P. Lee, H.I. De Lasa, Radial dispersion model for bubble phenomena in three-phase fluidized beds, Chem. Eng. Sci. 43 (9) (1988) 2445 – 2449. [20] U. Marianne, S. Frank, D. Anne-marie, D. Henri, G. Christiophe, Local measurements for the study of external loop airlift hydrodynamics, Can. J. Chem. Eng. 77 (1999) 375 – 382. [21] H. Li, A. Prakash, Influence of slurry concentrations on bubble population and their rise velocities in a three-phase slurry bubble column, Powder Technol. 113 (2000) 158 – 167.
Experimental study on bubble behavior in gas–liquid–solid three-phase circulating fluidized beds Tiefeng Wang, Jinfu Wang*, Weiguo Yang, Yong Jin Department of Chemical Engineering, Tsinghua University, Beijing 100084, PR China
Abstract The gas – liquid – solid three-phase circulating fluidized bed (TPCFB) is a novel multiphase reactor with outstanding performance in hydrodynamics, heat transfer, mass transfer and catalyst treatment. The object of this work is to study the bubble behavior in TPCFBs for better understanding of the hydrodynamics of the reactor. The local gas holdup, bubble size distribution, and bubble rise velocity in different radial positions are measured using a fiber optic probe, and the effects of operating conditions on bubble rise velocity are also investigated. Moreover, the bubble rise velocity and its distribution are compared with the data from gas – liquid two-phase external loop airlift reactors and gas – liquid – solid three phase fluidized beds without external circulation of liquid and particles. D 2003 Elsevier B.V. All rights reserved. Keywords: Bubble rise velocity; Three-phase fluidized bed; Particle circulation; Fiber optic probe
1. Introduction Gas –liquid – solid three-phase fluidized beds have been widely used in petrochemical, metallurgical, environmental and coal liquefaction processes [1]. Conventional threephase fluidized beds without external particle circulation are problematic when applied to systems with a catalyst liable to be deactivated or with small/light particles, moreover, the heat removal is a troublesome problem for a strong exothermic reaction. Based on the flow regime map of the gas – liquid – solid co-current three-phase fluidization proposed by Muroyama and Fan [2], Liang et al. [3,4] presented a new flow regime map taking into account the circulating fluidization regime operated in a high liquid velocity range. The particle entrainment leads to external particle circulation and the particle circulating rate can be controlled to regulate the solid holdup in the reactor. This feature is of great advantage for catalyst regeneration, reactor design and operation. Previous studies on hydrodynamics and mass transfer show that three-phase circulating fluidized beds (TPCFBs) have some better characteristics than the conventional three-phase fluidized beds [3– 8]. Liang et al. [4] investigated the axial phase holdup variation from a conventional fluidization regime to trans* Corresponding author. Tel.: +86-10-6278-5464; fax: +86-10-62772051. E-mail address: [email protected] (J. Wang). 0032-5910/$ - see front matter D 2003 Elsevier B.V. All rights reserved. doi:10.1016/j.powtec.2003.08.032
port regime. The axial profile of phase holdups in TPCFBs are rather flat, which is different from the convention threephase fluidized bed where a clear distinction exists between the bulk fluidized bed region and the freeboard region, and the solid holdup in the freeboard region is nearly zero. Han et al. [9] investigated the effects of the operating parameters on the axial and radial dispersion coefficients of the liquid phase and concluded that TPCFBs approach the ideal plug-flow reactor more closely compared with conventional threephase fluidized beds. Yang et al. [6] studied the radial distribution of liquid velocity using the electrolyte tracer method and found that liquid velocity distribution is nonuniform in the radial direction but more even than that in conventional three-phase fluidized beds. These results show that the hydrodynamic behavior in TPCFBs is different from that in convention three-phase fluidized beds; therefore further investigations are still needed. The bubble behavior, which is related to the profile of the phase holdup, the interaction between phases, and the mass transfer, have been important aspects for multiphase reactors. Remarkable efforts have been devoted to bubble behavior, such as the gas holdup, bubble size and bubble rise velocity in three-phase fluidized beds without external particle circulation [10 – 12], while very limited work on bubble behavior in TPCFBs can be found in the literature. Wang et al. [5] studied the bubble size distribution in TPCFB, and the data of the bubble size distribution was used to obtain the mass transfer coefficient and gas – liquid interfacial area from the
84
T. Wang et al. / Powder Technology 137 (2003) 83–90
beads of 0.4 mm diameter and density of 2460 kg/m3 are the solid phase. A detailed description was given in our previous study [5]. A fiber optic probe system was developed for measuring the bubble behavior in multiphase flow as shown in Fig. 2. The probe consists of two parallel communication optic fibers of 62.5 Am diameter with a distance 1.2 mm. The use of tenuous optic fiber improves the ability of detecting small bubbles (in the literature, an optic fiber of diameter 1 mm was used by several investigators [13,14]). The technique of emitting and receiving light in the same fiber makes the probe structure compact and ensures a quick signal response and slight disturbance to the flow field when measuring the bubble parameters. More than 1000 bubbles are sampled each measurement insuring that the effect of the sample size on the measured bubble chord length distribution is eliminated. A specially written program has been developed to control data sampling and to analyze the measured data [5].
3. Signal processing Fig. 1. Scheme of the experimental apparatus.
volumetric mass transfer coefficient. However, the bubble dynamic behavior has not been reported. This paper addresses the bubble rise velocity and its distribution in TPCFBs, which is important in determining the RTD of the gas phase and the diffusion between the gas and liquid phase.
2. Experimental A schematic diagram of the experimental apparatus is shown in Fig. 1. The riser is a vertical Plexiglas column, 140 mm in inner diameter and 3.0 m in height. Tap water and air are used as the liquid and gas phase, respectively. Glass
The bubble rise velocity can be calculated by cross correlating the two probe fiber signals. Typical signals are shown in Fig. 3. Due to the distance between the two probe fibers, the signal of the down-stream fiber is slightly delayed compared to that of the up-stream counterpart, which is clearer in the locally enlarged signal shown in Fig. 4. When a bubble rises through the probe, the up-stream fiber first contacts and pierces the upper surface of the bubble at time t1 and outputs a high level signal, and then the same course occurs to the down-stream fiber at t1V. The difference between t1Vand t1 is the time that a rising bubble takes to move from the up-stream fiber to the down-stream one. The bubble rise velocity can be calculated by dividing the distance between the two fibers, which is a given structure parameter, by t1V t1. Considering that the probe signals are not ideally rectangular, t1V t1 can be calculated by signal
Fig. 2. Hardware structure of the fiber optic probe for measuring bubbles.
T. Wang et al. / Powder Technology 137 (2003) 83–90
85
be calculated and the bubble size distribution can be further derived based on the algorithm proposed first by Liu and Clark [15] and improved by Wang et al. [5]. Thus the bubble size and rise velocity are obtained to characterize the bubble behavior.
4. Results and discussions 4.1. Distribution of ub
Fig. 3. Typical signals from the optical fiber probe.
cross correlation. The cross correlation function of the two tip signals V1(t) and V2(t) is as follows cðsÞ ¼
Z
tiþ1
V1 ðtÞV2 ðt sÞdt
ð1Þ
t¼ti
The value of s corresponding to the maximum of c(s) is the time delay t1V t1. Fig. 3 shows that two signals for most bubbles present high quality cross correlation, which makes it possible to obtain ub of each individual bubble. The distribution and average value of ub can be further determined by statistical analysis. With the determined bubble rise velocity, combined with the bubble signal duration time, the bubble chord length can
Fig. 4. Locally enlarged signal marked by the ellipse in Fig. 3.
Due to the bubble size distribution and flow turbulence, the bubble rise velocity is distributed, as shown in Fig. 5. Matsuura and Fan [10] studied the bubble behavior in a gas – liquid – solid fluidized bed with glass beads of 3 and 6 mm and a binary mixture of these particles. They found that in the dispersed bubble flow regime, the bubble rise velocity is narrowly distributed. In the coalesced bubble flow regime the bubble rise velocity distribution is similar to that in the dispersed bubble flow regime, except that the tail of the distribution in the coalesced bubble flow extends to much higher bubble rise velocities resulting in a larger bubble rise velocity. The experiments in this work have obtained similar results, as shown in Fig. 5. The TPCFB operates in the dispersed bubble flow regime and the bubble rise velocity distribution is narrow when the superficial gas velocity is low (Ug = 0.018 m/s), and operates in the coalesced bubble flow regime with a more spread distribution of the bubble rise velocity when the superficial gas velocity is higher (Ug = 0.054 m/s). 4.2. Radial profile of ub The radial profiles of the liquid velocity, gas holdup and bubble size all affect the radial profile of bubble rise
Fig. 5. Bubble rise velocity distribution at radial position r/R = 0 under different Ug (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s).
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Fig. 8. Effect of the superficial gas velocity on radial profile of the average bubble rise velocity (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137). Fig. 6. Bubble rise velocity distribution at different radial position (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s, Ug = 0.036 m/s, Gs = 10.8 kg/m2/s).
velocity. The distributions of ub in different radial positions shown in Fig. 6 indicate that ub in the central region has a larger average value and smaller standard deviation than that in the wall region. Bubble rise velocity is dependent on the liquid velocity, physical properties of the liquid – solid suspension, gas holdup and bubble size. Absolute bubble velocities (i.e. relative to the column wall) are a combination of the relative bubble velocity and the continuous phase velocity [16,17]. In TPCFBs, the bubble size has relatively narrow distribution and varies just slightly with the operating conditions within a certain range. The variation of the bubble size distribution with the superficial velocity is shown in Fig. 7. Under the experimental conditions (Ug = 0.018 – 0.063 m/s; Ul = 0.063 – 0.011 m/s), the bubble Sauter diameter varies in 4– 6 mm range [5]. In this range,
the rise velocity of a single bubble almost remains constant [18], so the influence of the bubble size distribution on the bubble rise velocity is negligible. The liquid velocity [6] and gas holdup [4] are higher in the central region, both favorable to increase the bubble rise velocity. The scatter distribution of ub in the near wall region implies that the distribution of ub is mainly caused by the turbulence of liquid phase and gas phase. Large vortexes in dynamic equilibrium of formation and disintegration were observed in the near wall region, causing more violent turbulence and more scattered distribution of ub. Fig. 8 shows the influence of the superficial gas velocity on the radial profile of average bubble rise velocity. The bubble rise velocity increases with the superficial gas velocity, and the maximum value is not present in the center line but at the position of r/R = 0.2, probably due to the spiral motion of bubble swarm. The radial profile of the bubble rise
Fig. 7. Effect of superficial gas velocity on bubble size distribution at radial position r/R = 0 (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s).
Fig. 9. Radial profile of the dimensionless bubble rise velocity ub/uba at different superficial gas velocity in TPCFBs under operating conditions shown in Table 1.
T. Wang et al. / Powder Technology 137 (2003) 83–90 Table 1 Operating conditions for Fig. 9 n
o
E
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Table 2 Bubble slip velocity q
x
E
a
B
*
Ug (m/s) 0.018 0.027 0.036 0.045 0.054 0.036 0.036 0.036 0.036 Ul1 (m/s) 0.054 0.054 0.054 0.054 0.054 0.054 0.063 0.072 0.081 Ul2 (m/s) 0.018 0.018 0.018 0.018 0.018 0.018 0.018 0.018 0.018 Gs (kg/ 10.5 10.5 10.8 11.0 11.5 10.8 10.5 10.3 11.2 m2/s)
Ug 0.018
0.018
0.018
0.027
0.027
0.027
Ul1 Ul2 r/R ubs
0.054 0.018 0.029 0.403a
0.054 0.018 0.058 0.379a
0.054 0.018 0 0.373a
0.054 0.018 0.029 0.467a
0.054 – – – 0.018 – – – 0.058 – – – 0.437a 0.28b 0.31b 0.37b
0.054 0.018 0 0.355a a b
velocity in conventional three-phase fluidized bed by Lee and De Lasa [19] is also shown in Fig. 8, showing that the radial profile of the bubble rise velocity is more even in TPCFBs. Radial profiles of dimensionless bubble rise velocity ub/uba are plotted in Fig. 9, showing that the variation of liquid velocity, gas velocity and particle circulating rate have little effect on ub/uba (Table 1). The radial profile of ub/ uba can be described by the following equation:
0.029 0.047 0.068
Data from our experiment. Data from Marianne et al. [20].
understanding of the multiphase flow and it is especially useful for testing the validation of CFD models. 4.3. Influence of operating conditions
ð2Þ
In TPCFBs, the particle circulating rate Gs and solid holdup can be adjusted by regulating the flow ratio between the primary and secondary stream of the liquid, thus the superficial gas velocity Ug, superficial liquid velocity Ul and particle circulating rate Gs can be independently controlled.
Fig. 10 illustrates the radial profile of average bubble rise velocity in gas – liquid two-phase external loop reactors [20]. It can be seen that radial profile of ub in TPCFBs is slightly flatter compared to that in the external loop reactors. This result is reasonable considering the uniform radial profile of hydrodynamic parameters such as the gas holdup and liquid velocity [4,6,8] reported by previous investigators. The bubble slip velocity ubs can be computed from the bubble and liquid velocity data. The results shown in Table 2 indicate that the slip velocities may differ considerably from the value of 0.25 m/s often used in gas –liquid two-phase system and gas – liquid –solid three-phase system [20]. Local information about the bubble rise velocity allows a better
4.3.1. Influence of Ug on ub The gas holdup eg increases with increasing Ug, leading to increase in bubble rise velocity. The distributions of ub under different Ug are shown in Fig. 5. Both the standard deviation and the average value of the bubble rise velocity distribution increase with increasing Ug. The influence of Ug on average bubble rise velocity is shown in Fig. 11, indicating that with increase of Ug, ub in the near wall region increases more obviously than that in the central region, leading to more uniform radial profile of ub. The influence of Ug on the bubble rise velocity can be explained by the bubble wake effect. When the gas holdup increases with Ug, the bubble number within a bubble swarm also increases, resulting in a decrease of the bubble moving friction and increase of the
Fig. 10. Radial profile of the dimensionless bubble rise velocity ub/uba at different superficial gas velocity in gas – liquid external loop reactors (data from Marianne et al. [20]).
Fig. 11. Effect of superficial gas velocity on average bubble rise velocity at different radial positions (Ul1 = 0.054 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/ m2/s, es = 0.137).
r r 2 r 3 ub ¼ 1:089 þ 0:416 þ0:5237 1:22 R R R uba
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bubble rise velocity. On the other hand, increasing Ug enhances the flow turbulence and bubble breakup and coalescence, causing more spread distribution of ub. The bubble rise velocity is influenced by the local gas holdup, bubble size and liquid velocity. The variation of the local gas holdup with Ug is also measured, and the results are shown in Fig. 12. The gas holdup in the near wall region does not increase as significantly as that in the central region. Considering that the bubble size has no significant influence on the bubble rise velocity, the liquid velocity in the near wall region should increase with an increase of Ug. This analysis is supported by the result of Yang et al. [6]. They reported that with an increase of Ug, the liquid velocity decreases in the central region and increases in the near wall region, resulting in a more even radial profile. 4.3.2. Influence of Ul on the bubble rise velocity The influence of Ul on ub is also experimentally investigated, and the results are shown in Figs. 13 and 14. It can be seen that Ul has no significant influence on ub, except a slight decrease of the proportion of bubbles with low velocity, which is reasonable since the superficial liquid velocity only varies in a relatively small range [3] ensuring that the reactor operates in the circulation regime. The bubble rise velocity is influenced by the gas holdup, local liquid velocity and solid holdup. Both the local liquid velocity and gas holdup increase with increasing U l , while the solid holdup decreases. The increase of the liquid velocity and gas holdup tends to increase the bubble rise velocity, while the decrease of solid holdup has the opposite influence. As a result, the bubble rise velocity has no remarkable variance with Ul, as shown in Fig. 14.
Fig. 13. Effect of superficial liquid velocity on bubble rise velocity distribution at radial position r/R = 0 (Ug = 0.036 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137).
4.3.3. Influence of Gs on the bubble rise velocity TPCFBs have the advantage to adjust Gs easily, which in turn influences es, by regulating the flow ratio between the
primary and secondary stream. With increase of es, the apparent viscosity and density of the suspension increase, causing larger buoyancy and bubble size, hence increase in the bubble rise velocity. The increase of es also increases the drag coefficient, tending to decrease ub, but this influence is more moderate than that mentioned above. As a result, the bubble rise velocity increases with Gs, and this variance becomes negligible when Gs exceeds a certain value, as shown in Fig. 15. Li and Prakash [21] studied the influence of slurry concentrations on bubble population and their rise velocities in a three-phase slurry bubble column (Ug = 0.10– 0.30 m/s; es = 0– 40% v/v), measuring the small and large bubble rise velocity using fast response pressure transducers and the dynamic gas disengagement technique. The rise velocities of both large and small bubbles increase with
Fig. 12. Radial profile of eg under different Ug. (Ul1 = 0.054 m/s Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137).
Fig. 14. Effect of superficial liquid velocity on average bubble rise velocity (Ug = 0.036 m/s, Ul2 = 0.018 m/s, Gs = 0.072 kg/m2/s, es = 0.137).
T. Wang et al. / Powder Technology 137 (2003) 83–90
R t Ug Ul1 Ul2 ub uba ubs u¯b V
89
radius of the fluidized bed (m) time (s) superficial gas velocity (m/s) superficial liquid velocity of the primary stream (m/s) superficial liquid velocity of the secondary stream (m/s) absolute bubble rise velocity (m/s) cross-sectional average bubble rise velocity (m/s) bubble slip velocity (m/s) local mean bubble rise velocity (m/s) output signal of the probe (V)
Greek symbols s time variable of c(s) (s)
Fig. 15. Effect of particle circulation rate on average bubble rise velocity (Ug = 0.036 m/s, Ul = 0.072 m/s).
increasing slurry concentrations up to slurry concentration of about 20% (v/v), consistent with our results.
5. Conclusions Experimental study on the bubble rise velocity with a fiber optic probe has resulted in the following conclusions. The bubble rise velocity has a radial profile with larger average value and smaller standard deviation in the central region than that in the near wall region. The radial profile of the bubble rise velocity is more even than that in conventional three-phase fluidized beds and gas – liquid external loop reactors in the literature. The average value and standard derivation of ub distribution increase with Ug. The superficial liquid velocity Ul has no significant influence on ub under the experimental conditions. The bubble rise velocity ub increases with es in low es range and this effect becomes negligible when es exceeds a certain value. The bubble slip velocity calculated from the local bubble rise velocity and liquid velocity is quite different from the value of 0.25 m/s often used in gas – liquid two-phase system and gas – liquid – solid three-phase system.
Notations c cross correlation function d distance between probe tips (m) db bubble diameter defined by the maximum vertical chord length (mm) Gs circulation rate of the solid particle (kg/m2/s) Ps(db) probability density function of the bubble size (1/mm) PDF probability density function of the bubble rise velocity (s/m) r radial position of the fluidized bed (m)
Subscripts g gas phase l liquid phase b bubble References [1] L.S. Fan, Gas – Liquid – Solid Fluidization Engineering, Butterworths Publishers, New York, 1989. [2] K. Muroyama, L.S. Fan, Fundamentals of gas – liquid – solid fluidization, AIChE J. 31 (1985) 1 – 34. [3] W.G. Liang, Q.W. Wu, Z.Q. Yu, Y. Jin, H.T. Bi, Flow regimes of the three-phase circulating fluidized bed, AIChE J. 41 (2) (1995) 267 – 271. [4] W.G. Liang, Z.Q. Yu, Y. Jin, Hydrodynamics of a gas – liquid – solid three phase circulating fluidized bed, Can. J. Chem. Eng. 73 (1995) 656 – 661. [5] T.F. Wang, W.G. Yang, Y. Jin, Experimental study on bubble size distribution in three-phase circulating fluidized beds, J. Chem. Ind. Eng., China 52 (3) (2001) 197 – 203. [6] W.G. Yang, J.F. Wang, J. Zhou, Y. Jin, Liquid phase flow structure and backmixing characteristics of gas – liquid – solid three-phase circulating fluidized bed, Chem. Eng. Sci. 54 (1999) 5293 – 5298. [7] W.G. Yang, J.F. Wang, L.M. Zhou, Y. Jin, Gas – liquid mass transfer behavior in three-phase CFB reactors, Chem. Eng. Sci. 54 (1999) 5523 – 5528. [8] T.F. Wang, J.F. Wang, W.G. Yang, Y. Jin, Experimental study on gasholdup and gas – liquid interfacial area in TPCFBs, Chem. Eng. Commun. 187 (2001) 251 – 263. [9] S.J. Han, J. Zhou, Y. Jin, K.C. Loh, Z.W. Wang, Liquid dispersion in gas – liquid – solid circulating fluidized beds, Chem. Eng. J. 70 (1) (1998) 9 – 14. [10] A. Matsuura, L.S. Fan, Distribution of bubble properties in a gas – liquid – solid fluidized bed, AIChE J. 30 (6) (1984) 894 – 903. [11] Z. Chen, C. Zheng, Y. Feng, H. Hofman, Local bubble behavior in three-phase fluidized beds, Can. J. Chem. Eng. 76 (1998) 315 – 318. [12] T. Miyahara, L.S. Fan, Properties of a large bubble in a bubble swarm in a three-phase fluidized bed, J. Chem. Eng. Jpn. 25 (1992) 278 – 382. [13] S. Saberi, K. Shakourzadeh, D. Bastoul, J. Militzer, Bubble size and velocity measurement in gas – liquid systems: application of fiber optic technique to pilot plant scale, Can. J. Chem. Eng. 70 (1995) 253 – 257. [14] A. Cartellier, Optical probes for local void fraction measurements: characterization of performance, Rev. Sci. Instrum. 61 (2) (1990) 874 – 886. [15] W. Liu, N.N. Clark, Relationships between distributions of chord lengths and distributions of bubble sizes including their statistical parameters, Int. J. Multiph. Flow 21 (1995) 1073 – 1089.
90
T. Wang et al. / Powder Technology 137 (2003) 83–90
[16] L. Sylvain, G. Christophe, Characterization of bubble column hydrodynamics with local measurements, Chem. Eng. Sci. 54 (1999) 4895 – 4902. [17] L.H. Caroline, L. Faical, G. Christophe, Understanding gas-phase hydrodynamics in bubble columns: a convective model based on kinetic theory, Chem. Eng. Sci. 52 (1997) 63 – 77. [18] K. Tsuchiya, A. Furumoto, L.S. Fan, J.P. Zhang, Suspension viscosity and bubble rise velocity in liquid – solid fluidized beds, Chem. Eng. Sci. 52 (1997) 3053 – 3066.
[19] S.L.P. Lee, H.I. De Lasa, Radial dispersion model for bubble phenomena in three-phase fluidized beds, Chem. Eng. Sci. 43 (9) (1988) 2445 – 2449. [20] U. Marianne, S. Frank, D. Anne-marie, D. Henri, G. Christiophe, Local measurements for the study of external loop airlift hydrodynamics, Can. J. Chem. Eng. 77 (1999) 375 – 382. [21] H. Li, A. Prakash, Influence of slurry concentrations on bubble population and their rise velocities in a three-phase slurry bubble column, Powder Technol. 113 (2000) 158 – 167.