(Gas)âLiquidâSolid Circulating Fluidized Bed Reactors

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Ind. Eng. Chem. Res. 2009, 48, 7876–7892

(Gas)-Liquid-Solid Circulating Fluidized Bed Reactors: Characteristics and Applications Arnab Atta,† S. A. Razzak,‡ K. D. P. Nigam,*,† and J-X. Zhu‡ Department of Chemical Engineering, Indian Institute of Technology Delhi, New Delhi 110 016, India, and Department of Biochemical and Chemical Engineering, UniVersity of Western Ontario, London, ON, Canada N6A 5B9

Accepting considerable advantages of circulating fluidized beds (CFBs) over the conventional fluidized beds, there has been numerous studies on CFBs concentrating primarily on the development of gas-solid circulating fluidized beds (GSCFBs). However a substantial amount of research has also been devoted to other two types of CFBs, namely liquid-solid and gas-liquid-solid circulating fluidized beds (LS and GLSCFBs). In this effort, an attempt has been made to summarize and review the research and progresses made on the last two types of CFBs since the highlighting of their hydrodynamics and potential applications in various industries by Zhu et al. [Can. J. Chem. Eng. 2000, 78, 82-94]. The issues associated with its hydrodynamics, scale-up, and design have been discussed with a re-emphasis on its potential application for various cost-effective processes. 1. Introduction Liquid-solid and gas-liquid-solid circulating fluidized beds (LS and GLSCFBs) are gaining very extensive recognition in a diverse field of industrial processes, e.g., many new processes in biochemical technology,1-3 wastewater treatment,4 petroleum, and metallurgical industries.5 A typical circulating fluidized bed (CFB) has an upstream of solid particles which are entrained upward in a column (called a riser), then collected/separated at the riser top (separator), and finally recirculated through a particle storage vessel or conventional fluidized bed (called a downcomer or downer) back to the bottom of the riser. It has been reported that the flow characteristics in the riser were uniform6 (Figure 1) as well as the fact that it has the ability to accommodate a varied set of particulate materials with high liquid throughputs, which give CFBs an upper hand over the conventional fluidized beds. In conventional fluidized beds, there are limitations on liquid and gas velocities and solid particles size and density. However in a CFB, solid particles are circulated between the riser and the downer at higher velocities compared to conventional fluidized beds, which leads to better contacting efficiency between phases, and higher mass transfer can be achieved with CFBs, which makes this type of reactor more preferable over the conventional fluidized beds. LSCFB and GLSCFB are likely to find an immense number of applications in the biochemical and biological processes, where the significantly enhanced interfacial contact efficiency can lead to much more effective processing means. In addition, the nature of the CFB with two units under one system makes it possible to make many bioprocesses continuous, leading to a further increase of efficiency and reduced processor size. For example, the University of Western Ontario4,7,8 have developed a new wastewater treatment process using an LSCFB bioreactor, where the riser is used for the anoxic process and the downer is used for the aerobic process, thus accommodating simultaneously both the critical processes for wastewater treatment. Using this newly developed liquid-solid circulating fluidized bed biological reactor (LSCFBBR * To whom correspondence should be addressed. E-mail: drkdpn@ gmail.com. Tel.: (011) 26591020. Fax: (011) 26591020. † Indian Institute of Technology Delhi. ‡ University of Western Ontario.

or CFBBR for short) for biological nutrient removal from municipal waste waters, they found a 10-fold increase in treatment efficiency for a 5000 L/d pilot plant, owing to the advantages of higher biomass concentration and lower hydraulic retention time (HRT) in LSCFB over conventional treatment systems. LSCFB can also be used for chemical processes.9,10 LSCFB (and also GLSCFB when gas is present) can be especially advantageous for industrial chemical and biochemical processes where the biosolid particles or catalysts need to be continuously regenerated to ensure uninterrupted mode of operation. In those cases, the deactivated catalysts, biomedia, ion exchange resins, or adsorbents can be regenerated continuously by circulating solid particles between the main reactor or contactor where the principle reactions or adsorption processes are accomplished and the regenerator where the regeneration or desorption of absorbents is carried out, in a closed and continuous loop.1 For example, LSCFB had been found to be a very useful and effective rector for the continuous recovery of protein from unclarified whole broth where the adsorption and desorption (regeneration) of proteins were carried out separately in the downer and the riser in a continuous mode with the ion exchange particles circulated between the two columns.2,3

Figure 1. Comparison of the radial solids holdup profiles in LSCFBs and GSCFBs under the same cross-sectional average solids holdup (εs ) 0.095). Reprinted with permission from ref 6. Copyright 2002 Elsevier.

10.1021/ie900163t CCC: $40.75  2009 American Chemical Society Published on Web 06/12/2009

Ind. Eng. Chem. Res., Vol. 48, No. 17, 2009

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Table 1. Summary of Experimental Studies with (G)-LSCFBs by Different Researchers after the Year 2000 bed properties researcher

LSCFB/GLSCFB; system studied

riser

downer

Cho et al.,11 Kang et al.,12 Kang et al.13

GLSCFB; air-water

ID ) 0.102 m, H ) 3.5 m

Zheng et al.,6 Zheng and Zhu14

LSCFB; water

ID ) 0.0762 m, H)3m

Vatanakul et al.15

GLSCFB; air-water

ID ) 0.076 m, H)2m

Zheng16

LSCFB; water, glycerol solutions LSCFB; synthetic wastewater

ID ) 0.076 m, H)2m ID ) 0.038 m ID ) 0.076 m, until H ) 1.03 m H ) 1.62 m then narrowed to ID ) 0.0254 m for another H ) 1.52 m ID ) 0.15 m, H)3m

7

Cui et al., Patel et al.,6 Chowdhury et al.17

ID ) 0.2 m

solid particles type

size (mm)

density (kg/m3)

glass beads

2.1

2500

polycarbonate (plastic) beads glass beads glass beads

0.526

1100

0.508 0.433, 1.30

2490 2500

glass beads

0.500

2500

lava rock

0.67

2560

glass beads

2.5

2540

glass beads

1.0, 1.7, 2.1, 3.0

2500

Roy and Dudukovic,18 Roy et al.19

LSCFB; water

Shin et al.,20 Cho et al.21

LSCFB; water, carboxymethylcellulose (CMC) GLSCFB; AirWater

ID ) 0.102 m, H ) 3.5 m ID ) 0.0762m H) 2.7m

ID ) 0.2m

Glass beads

0.5

2490

Razzak et al.,23 Razzak et al.24

GLSCFB; air-water

ID ) 0.0762 m, H ) 5.97 m

ID ) 0.2 m, H ) 5.05 m

glass beads

0.5

2490

Son et al.25

GLSCFB; compressed air-synthesized wastewater with aqueous solutions of CMC GLSCFB; compressed airsynthesized wastewater

ID ) 0.102 m, H ) 3.5 m

glass beads

1.0

2500

ID ) 0.102 m, H)1m

anion polymer resin

0.4

1130

styrene resin

1.45

1264

lava rock

0.67

2560

Zheng et al.,22

Son et al.26

Cao et al.27

GLSCFB; ID ) 0.15 m, air-carboxymethyl H ) 4.35 m cellulose sodium

Chowdhury et al.8

LSCFB; municipal wastewater

ID ) 0.02 m, H)3m

ID ) 0.076 m, H ) 1.62 m

Observing the widespread use and popularity of (LS and GLS) CFBs (Table 1), this effort is to summarize the characteristic studies contributed in this area. There is a review article available on (gas)-liquid-solid CFBs by Zhu et al.,1 and therefore, this study has been focused on the efforts made in this research area for the last ten years.

parameter studied characteristics of heat transfer coefficient and temperature fluctuations in riser for different gas, liquid, and solids circulation rates: liquid dispersion in the riser. The bubble distributions in the radial direction have also been measured to analyze the relation between the liquid dispersion and the bubble distribution in the riser. radial distribution of solids holdup and the effect of particle density on the flow structure An ultrasonic technique has been applied to simultaneously measure phase holdups through analyzing the fluctuation of amplitude ratio and transmission time. The effects of particle size, gas flow rate, superficial liquid velocity, and solids circulating rate on the axial and radial phase holdup distribution have been investigated. radial particle profiles with varying liquid viscosity simultaneous elimination of organic carbon, nitrogen, and phosphorus from municipal wastewater dispersion of solids and use of noninvasive radiation-based techniques, computer-automated radioactive particle tracking (CARPT), to explore the hydrodynamics of liquid-solid flows in vertical risers Overall heat-transfer coefficient and Liquid dispersion in the radial direction with viscous liquid medium onset liquid velocity for lower limit of the three-phase circulating regime use of electrical resistance tomography (ERT), pressure transducers, and fiber optic probes for qualitative and quantitative radial profiles of the phase holdups and propagation velocities characteristics of pressure fluctuations and bubble size in the riser for different gas and liquid velocities characteristics of gas holdup and gas-liquid mass transfer where the nitrogenous component was removed from the synthetic wastewater local phase holdups profile and liquid flow velocity. Liquid flow velocity measurements were performed using the electrolyte tracer measurement (ETM) technique. removal of organics, nitrogen, and phosphorus from municipal wastewater without particle recirculation and comparison of the nutrient removal efficiencies especially phosphorus removal with particle recirculation using a lab-scale LSCFB

2. Typical (G)-LSCFB Setup A schematic diagram of the experimental setup24 of (G)-LSCFB is presented in Figure 2. The setup mainly consists of two cylindrical sections, the riser and the downer, both made of Plexiglas. The dimension of the riser section

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Figure 3. Flow regime map for LSCFB [d*p ) dimensionless solid particle diameter, dp(Fg∆F/µ2)1/3; U*l ) dimensionless liquid superficial velocity, Ul(Fl2/µg∆F)1/3]. Reprinted with permission from ref 1. Copyright 2000 John Wiley & Sons, Inc.

Figure 2. Schematic diagram of a typical GLSCFB system. Reprinted with permission from ref 24. Copyright 2009 Elsevier.

is 5.97 m high with a 0.0762 m diameter, and the downer is 0.2 m in diameter with a 5.05 m height. A gas-liquid-solid separator is placed at the top of the riser to separate out the solids from the gas and liquid flow. To measure the solids circulation rate, a solids circulation rate measurement device is positioned near the top of the downer. The details of this reactor and its distributor system have been elaborated in the work of Razzak et al.24 To maintain the continuous particle circulation in the riser (as the solid particles are being transported to the top), an auxiliary liquid stream is utilized to facilitate the flow of solid particles from the downer to the riser thus acting as a nonmechanical valve. A gas distributor is placed above the liquid distributor which forms dispersed bubble flow in the riser assisting the solids to rise upward when combined with the liquid stream. 3. Characteristics of Liquid-Solid Circulating Fluidized Beds 3.1. LSCFB Flow Regime. The fluidization of liquid-solid systems is mainly controlled by the liquid flow rate. When liquid flow rate is lower than the minimum fluidization velocity, Umf, the particles inside the bed are static. With an increase in liquid velocity above Umf, some particles start to move, and with a further increment in liquid velocity, it begins to entrain the particles out of the bed. This is the onset of transition from conventional fluidization to circulating fluidization1 (Figure 3). According to Liang et al.,28 the transition to the liquid-solid circulating fluidization regime occurs at the point where the particle circulation rate becomes zero with a decreasing liquid velocity. This velocity was termed as the critical transition velocity, Ucr. Just above the critical transition velocity, particles are carried out of the bed, and this point indicates the beginning of circulating fluidization regime. Therefore, it facilitates easy solids feed into and discharge from the beds.7

The critical transition velocity, Ucr, defined by Liang et al.,28 was found to be dependent on system design, its operating conditions, the total solids inventory, and the solids feeding system. Since this experimental method involved the operation of the solids circulation which could be affected by the pressure balance of the system, therefore, the critical transition velocity intrinsically becomes system dependent. In order to achieve the lowest value of Ucr and to provide the convenient demarcation velocity which will be independent of system geometry, Zheng and Zhu29 proposed an onset velocity for circulating fluidization regime, Ucf. They employed a new method, the method of measuring the bed emptying time in a batch-operated fluidized bed (due to its reliability and simplicity), to determine the transition velocity from conventional fluidization to circulating fluidization which is independent of the operating conditions and equipment configurations of the liquid-solid systems. The liquid velocity thus determined was termed as the onset velocity of circulating fluidization, Ucf, which can be envisaged as the minimum transport velocity or the developing of significant entrainment velocity. Their investigation reveals that Ucr is the actual transition velocity for a given liquid-solids system while Ucf gives the minimum transition velocity for systems of various geometries, operating conditions, and variable particles size distributions (Table 2). Conventional fluidized beds generally operate below the onset velocity. On the contrary, CFBs need to be run beyond Ucf in order to achieve a continuous solids circulation.29 It has also been found that the onset velocity is a more intrinsic parameter compared to the critical transition velocity. Similar to the particle terminal velocity, it is also a function of the particle properties which increases with the particle mean size and density. 3.2. Hydrodynamics. In order to predict the stable operating conditions and to explain the origin of the unstable operation phenomena of the liquid-solid circulating fluidized bed, Zheng and Zhu30 carried out a pressure balance analysis. They prescribed a semiempirical equation for the pressure drop across the nonmechanical control valve, as well as a dimensionless empirical correlation for the solids holdup in the riser. On the basis of the experimentations carried out by this group31 with glass and plastic beads, they proposed the following dimensionless empirical correlation for the solids holdup varying with solids circulation rate and superficial liquid velocity:


js G

0.8

1-ε)

(1)

jl 0.25U

1.9

j s is the dimensionless solids circulation rate and is where G defined by js ) G

Gs

(2)

(µg∆FFl)1/2

j l (the dimensionless superficial liquid velocity) is and U j l ) Ul U

( ) Fl2 µg∆F

1/2

(3)

From these equations, it is apparent that the bed voidage in the riser is influenced by the solids circulation rate, liquid flow rate, and the physical properties of particles as well as fluidization media. For the pressure drop across the valve, Zheng and Zhu30 suggested that it increases rapidly with the solids circulation rate and decreases with auxiliary liquid velocity, which was presented as ∇Pv ) K

Gs,v2.51 2[Fs(1 - εmf) + Flεmf]

(4)

where K)

[Fs(1 - εmf) + Flεmf]gDv 0.125UaGs,v

(5)

Ua ) superficial velocity of the auxiliary liquid flow and

(

Gs,v ) 1 +

)

Flεmf Ws Fs(1 - εmf) Av

(6)

Therefore in compact form, it can be written as ∇Pv )

Gs,v1.51gDv 0.25Ua

(7)

This study mentioned that by overall pressure balance analysis, one can find a maximum solids circulation rate for a given auxiliary liquid velocity, beyond which a stable operation of the LSCFB system is not possible. Zheng and Zhu30 have also discussed the other important parameter for the stable operation range of LSCFBs, the auxiliary liquid flow rate, which determines the possibility of the unstable operation. It has been observed that at low auxiliary liquid velocity, the reactor used to be in a stable state owing to the fact of limited solids feed rate which allows the maximum stable liquid flow rate to be controlled only by pump capacity. The increment in auxiliary liquid velocity increases the chance of unstable operation. If the condition demands high auxiliary liquid velocity, then the maximum solids circulation rate can

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be achieved with the increased total liquid velocity in the riser column. For example, according to Zheng and Zhu,30 the available operation range of the liquid velocity increased from 0-0.25 to 0-0.32 m/s when the auxiliary liquid velocity reduced from 0.069 to 0.055 m/s for their experiments. Moreover the interesting thing to be noted from that study is that the solids circulation rate becomes insensitive to the variation of the liquid velocity when the liquid velocity is already set high. At that point, it will be wise to increase the pressure head available for the solids feeding system by adding more particles to the storage vessel and/or reducing the pressure loss across the valve to achieve a higher solids circulation rate.30 This study concludes that the system geometry (in particular, the storage-vessel-to-riser diameter ratio, the returning-pipe-toriser diameter ratio, the feeding-pipe-to-riser diameter ratio, the static bed height in the storage vessel, and the riser height) can be another important factor influencing the stable operating conditions in an LSCFB. 3.3. Flow Characteristics. For the last several years, a number of studies have been carried out to reveal the fundamental characteristics, such as the axial and radial particle distribution profiles, in liquid-solid circulating fluidized beds. For the axial profiles, it is reported that the distribution of particles is generally uniform along the riser, except for heavy particle systems in which a dense-bottom and a dilute-top structure occurs under relatively low liquid velocity.31 However, in the radial direction, nonuniformity of solids holdup distribution has been identified, with particle concentration increasing toward the wall of the riser.6,32 Zheng et al.6 showed the radial holdup distributions for two different types of particles (glass beads and plastic beads) under the same solids flow rate and different liquid velocities (Figure 4). The nonuniform distribution of radial solids holdup was established after the fluidized bed enters the liquid-solid circulating fluidization regime. The radial nonuniformity of solids holdup, dilute in the center and dense near the wall, can clearly be observed from the results. Comparatively better homogeneous distribution of the solids holdup was noted in the middle of the reactor and the magnitude was lower than the cross-sectional average solids holdup. Near the wall, solids holdup reached the maxima. With increasing liquid velocity, this nonuniformity was found to increase again to some extent. However, further increase in liquid velocity showed significant decrease in the radial nonuniformity of the solids holdup, which in turn indicated the transition from the circulating fluidization regime to the dilute transport regime. This study also established that under a specific operating condition, there exists a similar flow structure along the height of the circulating fluidized bed, indicating a rather uniform axial flow structure in the LSCFB. At the same time, even though parabolic profiles were observed for the radial distribution of solids holdup for both types of particles (glass beads and plastic beads), the local radial particle distribution for the light particles (plastic beads) appeared to feature a somewhat more uniform contour under the same crosssectional average solids holdup (Figure 5). In all cases, however,

Table 2. Onset and Critical Transition Velocities for Different Particles29

plastic beads glass beads I glass beads II glass beads III

Fs (kg/m3)

dp (mm)

Ucra (cm/s)

Ucf (cm/s)

Ut (cm/s)

Ucr/Ut

Ucf/Ut

1100 2490 2541 7000

0.526 0.508 1.000 0.580

1.17 6.47 24.84

1.15 6.45 16.30 23.70

1.0 5.9 14.4 21.6

1.17 1.10 1.15

1.15 1.10 1.13 1.10

Obtained under the highest L0 (L0 ) total solids inventory expressed as the initial static bed height in the storage vessel before the start of the experiments when all solids are in the storage vessel). a

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Figure 4. Radial distributions of solids holdup at four bed levels (H ) 0.3, 0.8, 1.2, 1.7 m) at different superficial liquid velocities for (a) glass beads and (b) plastic beads. Reprinted with permission from ref 6. Copyright 2002 Elsevier.

formity index (RNI) defined by Zhu and Manyele.34 Unlike conventional fluidization and the dilute liquid transport regimes, they found the radial distribution of local liquid velocity in the LSCFB to be more nonuniform with higher liquid velocity at the axis and lower near the wall. This nonuniformity becomes more prominent with increasing liquid velocity and solids circulation rate. The radial nonuniformity index (RNI), proposed by Zhu and Manyele,34 is used to measure up the radial distribution of liquid velocity and is defined as RNI ) Figure 5. Comparison of the radial solids holdup profiles for glass beads and plastic beads under the same cross-sectional average solids holdup (εs,avg ) 0.052) at H (axial location) ) 0.8 m. Reprinted with permission from ref 6. Copyright 2002 Elsevier.

the radial nonuniformity in the LSCFB is much smaller than that observed in the GSCFB.6 Zheng and Zhu33 have carried out microscale studies and confirmed the aforementioned axial uniformity and the radial nonuniformity in the LSCFB. Further investigations were carried out by Zheng16 with the objective to characterize the effect of varying liquid viscosity. In that effort, the effects of liquid viscosity on the solids behavior and flow structure in an LSCFB were investigated. The results revealed that the flow structures in an LSCFB with varying viscosity were axially uniform but radially nonuniform. The viscosity of fluidizing liquid reduces the nonuniformity of the particle distribution in the radial direction as the higher viscosity can limit the random fluctuation of solids flow and render the dynamic movement of particles. It also divulged that particles fluctuate vigorously near the wall with a maxima at r/R ) 0.8. On the other hand, except for the work of Liang et al.,32 there was no contribution about the local liquid velocity distributions. Liang et al.32 reported from their experiments carried out under limited operating conditions that a nonuniform distribution of liquid velocity also exists in the LSCFB. Zheng and Zhu14 have extended this study and determined the local liquid velocity in a laboratory-scale circulating fluidized bed by means of a dual conductivity probe. The radial nonuniformity of the liquid velocity distribution is then characterized by a radial nonuni-

σ(Vl) ) σmax(Vl)

σ(Vl)

√(Vl,max - Vj l)2 + (Vj l - Vl,min)2

(8)

Where Vl,min ) Ul,mf and Vl,max ) 2Ul; σ(Vl) ) the standard deviation when averaging the liquid velocity in the radial j l ) the crossdirection; σmax (Vl) ) the normalizing parameter; V sectional average liquid velocity; Vl,min ) the possible minimum liquid velocity; and Vl,max ) the possible maximum liquid velocity. RNI is described as the normalized standard deviation of the cross-sectional average liquid velocity. Therefore, this radial nonuniformity index must vary between 0 and 1, with larger values indicating more nonuniformity in flow structures. In this way, the RNI quantifies the radial nonuniformity of the local liquid velocity and allows the radial distributions of liquid velocity under various operating conditions to be compared on the same base. Through this microscale study, it has been shown (Figure 6) that the nonuniform distribution of liquid velocity in the radial direction was prevailed when superficial liquid velocity was set on a higher value than a threshold. Eventually, this threshold velocity coincided with the demarcation point between the conventional liquid fluidization regime and the liquid-solid circulating fluidization regime. The radial nonuniformity of liquid velocity was found to increase with increasing liquid velocity and then decreased until another threshold liquid velocity. Thereafter, RNI remained relatively constant with a value slightly higher than zero, which indicated a flatter radial profile. The second threshold velocity at high liquid velocity described the boundary between the liquid-solid circulating fluidization regime and the dilute liquid transport regime, as defined by Liang et al.28


Figure 6. Radial nonuniformity index (RNI) for the liquid velocity under various operating conditions. Reprinted with permission from ref 14. Copyright 2003 Elsevier.

It should also be noted that the RNI values in the liquid-solid circulating fluidization regime were encountered to be much higher than those in both the particulate fluidization regime and the dilute liquid transport regime and thus indicated a more nonuniform radial distribution of local liquid velocity in the LSCFB. Roy et al.19 assessed the flow patterns and phase distributions in the riser of an LSCFB with their extensive experimental investigations by sophisticated noninvasive flow-mapping techniques, namely computer automated radioactive particle tracking (CARPT) and gamma-ray computed tomography (CT) for the first time with bigger solid particles. They primarily concluded regarding the time-averaged solids volume fraction obtained by gamma-ray CT that the solids distribution in the riser was uniform, with some minor accumulation of solids at the walls. With detailed experimental investigation with the CARPT, it was shown that in a time-averaged sense, the solids were found to be flowing up at the center and flowing down at the wall which advocates evidence of large-scale convective backmixing. In another study, Cho et al.21 investigated the radial dispersion characteristics of the continuous liquid phase in liquid-solid circulating fluidized beds. Pressure fluctuations have also been analyzed to consider the relation between the flow behavior of fluidized particles and the liquid radial dispersion in the riser. With their systematic investigation, it was concluded that the radial dispersion coefficient of the viscous liquid medium decreased with the increase in liquid velocity as well as viscosity. However, it was found to increase with an increment in the solids circulation rate as well as particle size when considered under the same solids circulating conditions. The liquid radial dispersion coefficient was observed to be strongly influenced by the dominant frequency of pressure fluctuations due to the flow behavior of fluidized particles in the bed of a viscous liquid medium. Again, this dominant frequency of pressure fluctuations appeared to decrease with increasing liquid velocity or viscosity and was increasing with the increase in size of fluidized solid particles under given solids circulating conditions.21 3.4. Heat Transfer Studies. The flow and heat transfer characteristics of liquid-solid circulating fluidized beds has been seldom studied until recent years. However, the temperature in the reactor has to be controlled effectively for its efficient industrial application despite complicated (viscous) liquid-solid contacting. From visual observation, it can be found that a LSCFB system exhibits particulate fluidization where particles are distributed homogeneously in the liquid stream throughout the riser at low superficial liquid velocities. On the other hand, at higher liquid superficial velocities, an aggregative fluidization

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appears in the riser, where the particles tend to form aggregates and fluidizing liquid channels its way among the aggregates. Within a limited operating condition, Kuramoto et al.35 measured the fluctuations of voidage and heat-transfer coefficient to analyze the particulate and aggregative fluidization in the riser of a liquid-solid circulating fluidized bed. It has been observed that a higher heat transfer rate can be obtained in the aggregative fluidization regime compared with that in the particulate fluidization regime. Shin et al.20 recently investigated the relevant information on the heat transfer in the riser containing a viscous liquid medium with the idea of a two resistances-inseries model between the immersed heater and the bed proper of the riser. The immersed heater-to-bed heat-transfer coefficient in the riser of a liquid-solid circulating fluidized bed was noticed to increase with an increasing solids circulation rate since there was an understandable increase in the solids holdup with increasing solids recirculation rate. It was also reported that the heat-transfer coefficient increased with increasing particle size due to the fact that the larger particles can have a larger inertial force to move and generate the turbulence in the viscous liquid medium. However, it was found to decrease gradually with increasing liquid viscosity. This can be attributed to the effective mobility of the particles in the bed with increasing viscosity of the liquid medium. With slower movement of the particles, there is decrease of turbulence in the bed with a higher viscosity liquid medium which in turn results in a decrease of heat-transfer coefficient between the immersed heater and the bed proper. According to this study, the thickness of liquid thin film around the heater surface which has a governing effect for the determination of overall heat-transfer coefficient was not found to change appreciably with increasing liquid velocity. This can be explained as the competing effects of the increase of turbulence and decrease of solids holdup in the bed with the increasing liquid velocity. Whereas, the thickness of the thin liquid film was found to decrease gradually with the increasing rate of solids circulation because of the increase of contacting frequency between the fluidized particles and the heater surface by increasing the solids holdup in the bed. Furthermore, it has been observed that the thickness of the thin film was increasing progressively with increasing liquid viscosity due to the enhanced adhesion force of the liquid medium to the heater surface with the increasing viscosity of the liquid phase.20 In other terms, it advocates the decrease in heat-transfer coefficient with increasing viscosity of the continuous liquid medium. Very recently, Hashizume and Kimura36 experimentally showed that the heat transfer coefficient of liquid-solid circulating fluidized beds is usually larger than that of singlephase liquid flow when operated with low liquid velocities. With increasing liquid velocity, the heat transfer coefficient (of LSCFB) gradually increases and approaches toward the heat transfer coefficient of a single-phase liquid flow. This region where the heat transfer coefficient is larger than that of the single-phase liquid flow has been termed as “the heat transfer enhanced region or enhanced heat transfer region (EHR)” by the authors. With further increment of liquid velocity, the heat transfer coefficient coincides with the heat transfer coefficient for a single-phase liquid flow. This region is termed as liquid single-phase heat transfer region (LSP). On the basis of the experimental data, they have proposed a single correlation for predicting the heat transfer coefficient in the entire region from the enhanced heat transfer region (EHR) to the liquid single phase heat transfer region (LSP), with an accuracy of (15%:

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NuD ) (NuD,EHR5 + NuD,LSP5)1/5

(9)

where NuD,EHR )

()

D RD ) 0.06ReS0.75Pr0.4 λL dp

(10)

RD ) 0.023ReDh0.8Pr0.4 λL

(11)

and NuD,LSP )

The heat transfer coefficient in the enhanced heat transfer region (EHR) has been estimated based on the slip velocity between liquid and solids particles. Res )

Usdp νl

where Us ) Ul - Up )

ul up 1 - εp εp

3.5. Studies on Modeling. The flow characteristics of liquid-solid fluidization, regardless of the flow regime and external particle circulation, are considerably different from the gas-solid processes. This has also been established by the researchers with the help of various experimental methods. Conventionally, liquid-solid fluidization has been considered as a uniformly dispersed homogeneous process which is very different from a heterogeneous gas-solid process.37 After enlightenment of the great potential of LSCFB in the ever growing fields of biotechnology, food processing, wastewater treatment, and petrochemical and metallurgical processing by Zhu et al.,1 owing to its many advantages (such as efficient liquid-solid contact, favorable mass and heat transfer, reduced backmixing of phases, and integrated reactor and regenerator design), and also with the rapid as well as successful development of computational fluid dynamics (CFD) in the multiphase reactor area, there has been a few interesting studies on modeling of hydrodynamics in LSCFBs. As discussed earlier, with the state-of-art noninvasive methods of flow characterization (γ-ray computed tomography (CT) and computer-automated particle tracking (CARPT)), Roy and Dudukovic18 showed significant nonuniformity in the radial distribution of local particle velocity, with higher solids velocities in the center and downward flow near the wall. These observations suggest that modeling studies with the assumptions of homogeneous fluidization are not applicable to LSCFBs and probably will lead to significant errors in the predictions. However, for efficient operations and to avail the optimized performance of LSCFBs especially for some cost-effective processes, the design and scale-up have to be based on the sound physical mechanisms to predict the flow fields and, subsequently, the reactor performance. With a view toward the prediction of flow fields and subsequent reactor performance, Roy and Dudukovic18 introduced a two-fluid CFD model for LSCFBs using the kinetic theory of granular flow to describe the solids phase. Adopting a two-dimensional axisymmetric geometry, they solved a two-fluid Eulerian model with proper closure for the liquid-solids momentum exchange term or drag force which envisaged both liquid and solid phases as interpenetrating continua. Their computed time-averaged solids axial velocity profile, when compared with the time-averaged velocity profiles obtained by CARPT, showed very good correspondence (Figure 7). The satisfactory comparison of the time-averaged solids radial holdup was also presented against the axially averaged, timeaveraged solids holdup profile measured by CT (Figure 8).

Figure 7. Comparison of CFD results of solids velocity profile with experimentally measured (by CARPT) axially averaged, time-averaged solids axial velocity (S/L ) solid-to-liquid flow ratio). Adapted from ref 18.

Figure 8. Comparison of CFD results of solids holdup profile with experimentally measured (by CT) axially averaged, time-averaged solids axial velocity. Adapted from ref 18.

Cheng and Zhu37 pursued the same line of thought to investigate the effects of other operating conditions, particle properties (i.e., diameter and density), and equipment sizes on the hydrodynamics and compared this against the available data in the literature to establish a reliable hydrodynamic model with a good basis for the reactor simulation (Figure 9). Armed with their successful effort to develop a robust model by validating different set of experimental data obtained under different operating conditions, they attempted to predict the influence of physical properties of fluids/solids on the two-phase flow structures (Figure 10). They have investigated 3 different systems [case I water-glass beads; case II water-steel shots; case III air-glass beads] to compare radial flow structures. In the case of only liquid-solid systems (i.e., with water-glass bead/steel bead, case I or II), it can be observed that there is hardly any difference in uniformity of the radial flow profile for different solid particles; however, quite logically, solids velocities are less in magnitude for heavier particles (steel shots). In accordance with previous discussions, it also shows that case III (gas-solid systems) has a clear nonuniform flow structure. Encouraged by this qualitatively as well as quantitatively validated model, Cheng and Zhu37 investigated the influence of bed dimension on the radial flow structure (Figure 11). With


Figure 9. Validation of model predictions by Cheng and Zhu37 with the experimental data (508 µm glass beads, x/H ) 0.86). (a) Effect of superficial liquid velocity on the flow structure. (b) Effect of solids circulation rate on the flow structure. Reprinted with permission from ref 37. Copyright 2005 John Wiley & Sons, Inc.

Figure 10. Model predictions showing the effect of fluid/solid properties on the radial flow structures in liquid-solid or gas-solid risers (x/H ) 0.8). Reprinted with permission from ref 37. Copyright 2005 John Wiley & Sons, Inc.

Figure 11. Influence of bed dimension on the radial distributions of liquid velocity (VL) and particle velocity (VS) in LSCFBs (x/H ) 0.8) [System: water-glass beads (508 µm). Operating conditions: UL ) 0.15 m/s, Gs ) 10 kg/m2 · s. Bed dimension: (D, H) ) (0.076, 3), (0.140, 3), (0.300, 12), (0.600, 24)]. Reprinted with permission from ref 37. Copyright 2005 John Wiley & Sons, Inc.

the increase of bed diameter, the computed radial distribution of liquid velocity and particle velocity were found to be more

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nonuniform. Higher negative velocities for liquid and solid phases were observed near the walls in larger diameter columns, which in turn signifies severe backmixing in the axial direction for larger diameter columns. It essentially demonstrates the complexity of scale-up which is generally dependent on experience and empirical correlations. Cheng and Zhu38 further concentrated on the CFD simulations to investigate the hydrodynamics in LSCFBs with different bed dimensions and liquid-solid systems with a primary focus on the scale-up issue, i.e., how to establish dynamic similarity between LSCFBs. In line with that, they evaluated the similitude method (or dimensional analysis) for its use in the scale-up of LSCFBs as far as its availability, simplification, and limitation, by a detailed comparison with the predicted hydrodynamics in both axial and radial directions using the experimentally validated CFD model.37 Apart from bed geometry, particle sphericity, and particle size distribution, it can be considered that the (a) riser diameter, D, (b) superficial liquid velocity, U0, (c) solids circulation rate, Gs, (d) mean particle diameter, dp, (e) solid particle density, Fs, (f) fluid density, Ff, (g) fluid viscosity, µf, and (h) acceleration due to gravity, g, are needed to determine the underlying flow patterns for the operation of any typical LSCFB. Cheng and Zhu38 tested the capability of the similitude method in scaling up LSCFBs and to establish the similarity in hydrodynamics (1) of different size LSCFBs and (2) of the same size LSCFBs but with different particle systems. The results demonstrated that matching the full set of five dimensionless groups can ensure hydrodynamic similarity in the fully developed region, except for the turbulent kinetic energy of the liquid phase. Reducing the number of dimensionless groups leads to less desirable matching. Scale-up based on the viscous-limit dimensionless group set causes larger deviations compared with the other scaling sets which implies that the fluid inertia should not be neglected in modeling and scaling of LSCFBs. With the limitation on experimental validation in practical uses, it has been suggested to use the combination of a reliable CFD model with the proper similitude scale-up for more promising and better reactor design, scale-up, and operation.38 Very recently, Razzak et al.39 carried out ample parametric numerical studies to provide a more detailed view on how an LSCFB operates under different operating parameters which include the variation of the solids circulation rate and primary and auxiliary liquid velocities. Thereafter, numerical modeling was carried out to predict the behavior of different particles with different densities upon fluidization in an LSCFB, which resolves the problem of experimentation with a wide spectrum of new particles that might have a wide variety of applications in an LSCFB (Figure 12). The simulations were carried out using glass beads of three different diameters: 200, 508, and 1000 µm. The cross-sectional average solids holdup was found to be almost identical for all three different size glass beads (200, 508, and 1000 µm). The radial parabolic profile trend of dimensionless solids holdup for different size glass beads were also similar in all cases. However, larger particles tend to have a higher degree of nonuniformity in the radial profiles of the solids holdup. Solids and liquid velocities of the smaller particles (200 µm) were much higher than the velocities of larger particles (508 and 1000 µm) in the central region because the weight of the particle increases with the increase in size. It is apparent that the heavier particles have less capability of flowing. Particle-particle interactions and drag forces are also high for larger particles. In summary, the two-fluid Eulerian model presented in all the modeling studies certainly provides a good estimation of the behavior of different types of particles in the flow system

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Figure 13. Flow regimes of an air-water-0.405-mm glass beads circulating fluidized bed. Adapted from ref 41.

Figure 12. Influence of the particle size on the flow structures (Gs ) 10 kg/m2 · s and H ) 1.7 m). Reprinted with permission from ref 39. Copyright 2008 Elsevier.

when coupled with dimensionless analysis without doing real life experiments which can substantially reduce the experimental effort to generate the large amount of data that is essential for scale-up and commercial applications. 4. Characteristics of Gas-Liquid-Solid Circulating Fluidized Beds In continuation with the earlier discussions, it can be understood that there are certain advantages of using a GLS circulating fluidized bed reactor in place of conventional fluidization systems. In particular, such places where small, porous, or light (sometimes precious) particles have to be fluidized in the viscous liquid medium in the presence of gas for many industrial processes, a three-phase circulating fluidized bed would be preferred over the conventional system. It has been pointed out that three-phase circulating fluidized beds can minimize dead zones and increase the contacting efficiency

among gas, liquid, and solid phases in the riser by enhancing the shear stress at the interfaces among individual phases.11 It has also been reported that the GLSCFB provides higher gas holdup, more uniform bubble sizes, better interphase contact, and good heat and mass transfer capabilities.40 The main reason for potential use is to have the flexibility to operate this type of CFB at a much higher liquid velocity than the minimum fluidization velocity of particles which in turn considerably increases the fractional conversion as well as production efficiency per unit cross-sectional area of the system. Furthermore, the precious deactivated catalysts can be regenerated continuously with this circulating mode.11 4.1. GLSCFB Flow Regimes and Hydrodynamics. In 1995, Liang et al.41 experimented with the flow regimes of the threephase circulating fluidized bed using 0.4 mm diameter glass beads with air and water as the gas and liquid phases, respectively. Understanding the fact that the pressure gradients in the lower and upper section of riser will become constant when the three-phase fluidized bed transforms into a transport bed, the distinctions between the three-phase circulating fluidization, transport, and expanded bed regime were demarcated (Figure 13). However, it has been observed that the pressure balance of the whole unit has great influence on the overall pressure gradient of the riser.29 Therefore, Zheng et al.22 attempted to redefine the intrinsic boundary between the expanded bed regime and the three phase circulating regime which will be affected by the physical properties of the system only. Their study revealed a new technique, independent of pressure balance of the system, where the demarcation was determined by measuring the bed emptying time in a batchoperated fluidized bed. The onset liquid velocity was found to decrease with increasing gas velocity; however, it was observed to be minimally influenced at low gas velocity. In order to develop a good understanding of the flow regime in gas-liquid-solid fluidization, Jena et al.42 have studied detailed hydrodynamics (viz. the pressure drop, minimum fluidization, bed expansion, and phase holdup) which eventually may assist better design and scale-up of industrial GLSCFB reactors. This study disclosed that the minimum liquid fluidization velocity (Vlmf) increases with increase in particle size at constant gas velocity but decreases with increase in gas velocity at constant liquid velocity (Table 3). The expansion ratio increases with increase in the liquid and gas velocity and decreases with increase in the particle size and static bed height. The gas holdup was found to increase with gas velocity and


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42

Table 3. Comparison of Minimum Fluidization Velocity for Different Particle Sizes at Different Gas Velocities dp (mm)

Vg ) 0 m/s

Vg ) 0.02 m/s

Vg ) 0.04 m/s

2.18 3.05 4.05

0.0255 0.0297 0.03040

0.0212 0.0255 0.0297

0.0170 0.0212 0.0255

particle size. However, for a fixed gas velocity, it was found to decrease with increase in liquid velocity for low liquid velocity cases. This study also revealed that further increase in liquid velocity did not have any effect on gas holdup. Very recently, pressure fluctuations and bubbling phenomena in the riser of a three-phase circulation fluidized bed bioreactor with viscous liquid medium were investigated by Son et al.25 They measured the bubble size and dynamic bubbling behaviors in the riser. It was found that the bubble size increased with increasing gas velocity and/or liquid viscosity but decreased with increasing liquid velocity. The bubbling phenomena was observed to become more complicated, and bubble size distribution tended to be broad, with increasing gas velocity and/or liquid viscosity. In line with their previous studies51 based on isotropic turbulence theory for deriving a correlation for bubble size, the particle Reynolds number has been related to the energy dissipation rate in the beds based on Kolmogoroff’s local isotropic turbulence theory. The energy dissipation rate of the liquid phase was calculated from the knowledge of individual phase holdups, fluid properties, and fluid velocities as ED )

(Ul + Ug)(εgFg + εlFl + εsFs)g εlFl

(12)

Since gas-liquid contact occurs through dispersed bubbles in three-phase fluidized beds, the particle Reynolds number was then replaced by the Reynolds number based on the mean bubble diameter for deriving the correlations of liquid side mass transfer coefficient as:

( )(

kLDb ν )R Dν Dν

m

EDDb4 ν3

)

n

) R(Sc)m(Reb)n

(13)

where m and n are correlation coefficients. Following similar kinds of studies of correlating data by isotropic turbulence theory, Son et al.25 has proposed a correlation for bubble diameter as

( )

(

Db Ug ) 36.97 dp Ul + Ug

) ( ) 0.249

µl DνFl

0.037

(14)

where, Db is bubble size and Dν is diffusivity In this context, it is worthwhile to mention that the movement of particles can be retarded in the viscous solution, thus the larger particles may require more energy to break the bubbles effectively. This phenomenon may lead to larger size bubble existence in the case of bigger particles.51 Razzak et al.23,24 employed electrical resistance tomography (ERT), a noninvasive measurement technique based on conductivity of the continuous phase, to characterize phase holdups and phase propagation velocities profiles qualitative as well as quantitatively. This imaging technique successfully enabled them to measure phase distribution and propagation in a GLSCFB. With the drawback of measuring solids holdup by this imaging technique (as it is impossible to differentiate between two nonconductive phases, e.g. solids and gas by this method), solids holdup was measured simultaneously by applying pressure transducers. It was detected that bubbles had the tendency to

Vlmf

Vg ) 0.06 m/s

Vg ) 0.08 m/s

Vg ) 0.1 m/s

0.0127 0.0170 0.0212

0.0085 0.0149 0.0181

0.0085 0.0127 0.0149

accumulate and move in the central regions. Due to wall effect and solids back mixing caused at the wall regions, the solids holdup was found to be comparably higher near the walls. Again, due to the limitations in the ERT system, phase propagation velocities can only be measured for the nonconductive phase which was the combination of the solids and gas in their studied GLSCFB system.23 The phase propagation velocities were higher in the central region compared to the wall region owing to the wall effect phenomena. It was noted that phase propagation velocities decreased with increasing liquid velocities. This was attributed to the fact that the liquid velocity increment resulted in a decrease in gas-solid phase holdup and its distortion by liquid convective currents. On the other hand, the phase propagation velocity increased by increasing the gas velocity as the nonconductive phase seemed to keep its identity for a longer period of time. However, more attention is required to explore the dependency of the phase propagation velocity on phase holdups and operating parameters in GLSCFB systems.23 Considering the inherent limitation of ERT (i.e., not able to differentiate between the other nonconductive phases, e.g. solids and gas), Razzak et al.24 later came up with a new method to determine the phase holdups of all phases for a better understanding of hydrodynamics in GLSCFB systems. Combining the advanced technology of ERT and the optical fiber probe revealed the phase holdups in GLSCFB riser more quantitatively. The optical fiber probe was capable of measuring gas holdup, which was one of the nonconductive phases. Combination of both technologies successfully employed to measure three-phase holdups distinctly. In line with previous studies, this study showed the radial distributions of gas, solids, and liquid phase holdups at different superficial gas and liquid velocities (Figure 14). It was observed that the solids holdup initially remained constant until the central location was reached; here, it started to increase radially toward the wall. Meanwhile, the opposite trend was revealed for the gas holdup as it sharply decreased near the wall region. This can be explained as the effect of nonuniform liquid velocity distribution in a cylindrical geometry. Due to higher liquid velocity in the center, there exists an inward pressure drop which forces the lighter gas bubbles to move toward the center of the riser. However, this pressure is not effective in the case of heavier solid particles. Again, the cross-sectional average of gas holdup was found to increase with the increases of superficial gas velocity and was observed to decrease with the increases of superficial liquid velocity (Figure 15). On the contrary, the remaining phase (liquid phase) showed the reverse trend. Liu et al.43 studied the hydrodynamics in a slurry airlift reactor at high solids concentrations. The influence of the average solids concentrations, superficial gas velocities, and particle sizes on the radial and axial profiles was addressed in their work. They measured solids holdup by using electrical conductivity probe. In their investigations, they found that, at lower solids concentrations, solids holdups are uniform in the radial direction but, at higher solids concentration situations, radial nonuniformity is profound. The axial profile of cross-sectional average solids concentrations was found to be uniform at all conditions even when operated with higher solids concentrations.

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Figure 14. Radial distribution of gas, liquid, and solids holdup measured for different superficial gas velocities at a superficial liquid velocity (Ul) ) 5.6 cm/s and auxiliary liquid velocity (Ua) ) 1.4 in a GLSCFB riser using the combination of ERT and optical probe [axial location (H) ) 2.02 m above the distributor]. Reprinted with permission from ref 24. Copyright 2009 Elsevier.

4.2. Mass Transfer Studies. GLSFBs or conventional threephase fluidized beds nowadays are widely used in different industrial applications due to their good mass transfer characteristics over packed bed systems. They have a larger interfacial area and can handle smaller size particles. Their only limitation is that they cannot handle wider ranges of particle sizes under higher velocities. GLSCFBs on the other hand have better hydrodynamics behavior over smaller particles and can handle even larger and more dense particles as well. In many application processes, solids are used as catalysts or inert as carrier materials where gas is adsorbed on the solid phase. When adsorption is the rate-limiting step, it is necessary to fully understand gas-liquid-solid mass transfer.44 Yang et al.45 studied gas to liquid volumetric mass transfer behavior using an oxygen dissolution method in GLSCFB system. They proposed a one-dimensional axial dispersion model (ADM) to describe the gas-to-liquid mass transfer behavior in gas-liquid-solid three-phase CFB reactors, in which the gasto-liquid volumetric mass transfer coefficient, kLa, has been determined by measuring the axial distribution of dissolved oxygen concentrations. To measure the dissolved oxygen concentrations in the liquid, experiments were carried out in a Plexiglas riser column 140 mm in internal diameter with a 3 m height. Water and air were used as the liquid and gas phases where glass beads of 0.4 mm in size were used as solid particles.

Figure 15. Cross-sectional average gas, liquid, and solids holdup measured for different superficial liquid velocities and gas velocities at an auxiliary liquid velocity (Ua) ) 1.4 cm/s in a GLSCFB riser using the combination of ERT and optical probe. Reprinted with permission from ref 24. Copyright 2009 Elsevier.

Due to the particle circulation in GLSCFBs, mass transfer between gas and liquid is different than that in the conventional fluidized system. Yang et al.45 found some different characteristics under different operating conditions. Usually in a conventional three-phase system, mass transfer is mainly controlled by supercritical gas flow. In their studies, they found superficial liquid velocities also have some degree of influence (Figure 16). The effect of superficial liquid velocity on the volumetric mass transfer coefficient, kLa, is presented shown in Figure 16. Experimentally measured kLa values under different liquid superficial velocities suggest that the mass transfer coefficient generally increases with increasing liquid superficial velocity except in cases of small gas flux through the bed where the influence is negligible. This is because an increase in liquid velocity may result in an increase in turbulence and a consequent vigorous solid mixing can happen which may break bubbles effectively. Thereby, the gas-liquid interfacial area increases due to the decrease in bubble size which in turn results in increasing kLa with increasing liquid velocity.51 Gas-liquid mass transfer is closely related to the hydrodynamic behavior of the GLSCFB system, and this behavior is different in a conventional bed system. In conventional bed system, mass transfer was dominant mainly with axial dispersion where the bed seems to be homogeneous radially. However, in the case of GLSCFBs, radial nonuniformity needs to be considered as one of the mass transfer controlling steps. Liu et al.44 studied both hydrodynamics and mass transfer simultaneously under different operating conditions. They validated


kLa ) 0.069Ul

0.042

Figure 16. Influence of superficial liquid velocity on mass transfer coefficient. (a) Secondary liquid velocity ) 0.0181 m/s. (b) Secondary liquid velocity ) 0.0145 m/s. Adapted from ref 45.

their experimental data with the axial dispersion model (ADM) which is suitable with the GLSCFB system. Local gas-liquid volumetric mass transfer coefficients have been evaluated by measuring oxygen concentrations in seven radial positions. Thereafter, they fit their experimental data with the ADM model. It has been reported that the influence of superficial gas velocities over the mass transfer coefficient, (kLa)r, is higher than the superficial liquid velocity and solids circulation rate. The mass transfer rate was found to be higher in the central region than in the wall region. It has been reported that bubble sizes are smaller in the center than near the wall, and this difference is eliminated in cases of lower gas velocities. This phenomena results in nonuniformity of radial mass transfer with the increase of gas velocities. Increasing solids circulation rate will result in the presence of more particles in the riser column which in turn lead to a better rate of bubble coalescence and reduction of gas-liquid interfacial area. Therefore, it was evident that mass transfer coefficients dropped with an increasing solids circulating rate. In another study, Son et al.26 has presented the gas holdup measurement to analyze the effect of bubbling phenomena on the volumetric gas-liquid mass transfer coefficient in a threephase circulating fluidized-bed bioreactor. The effects of the gas and liquid velocities and the holdup of the biofilm media on the volumetric gas-liquid mass transfer coefficient was studied to disclose the prerequisite knowledge for the design and scale-up of three-phase circulating fluidized-bed bioreactors or contactors. However, in line with the previous studies, they have also ascertained the fact that the value of the gas holdup decreases slightly with increasing liquid velocity or holdup of the biofilm media. The correlations for gas holdup and the volumetric gas-liquid mass transfer coefficient in the riser were given as, respectively,26 εg ) 0.15U1-0.047Ug0.303εs-0.05

(15)

Ug

0.162

0.136

εs

7887

(16)

These equations were modeled to cover the following ranges: 0.01 e Ul e 0.03 m/s 0.005 e Ug e 0.05 m/s 0.05 e εs e 0.2 Therefore, it can be seen that the gas velocity is a vital parameter in deciding the value of the mass transfer coefficient in the riser. Surprisingly, the volumetric gas-liquid mass transfer coefficient seemed to have very sparing effect with variation of the liquid velocity in the riser unlike with increasing gas velocity or holdup of the biofilm media which influenced considerably the increase of mass transfer coefficient. 4.3. Heat Transfer Studies. There is very scant literature available on the heat transfer characteristics in three-phase circulating fluidized beds. Cho et al.11 investigated characteristics of heat transfer coefficient and temperature fluctuations in the riser of a three-phase circulating fluidized bed. Highly complicated, nonlinear temperature fluctuations and heat transfer in three-phase circulating fluidized beds were analyzed and described with the help of deterministic chaos theory. It was remarked after the analysis of bubbling phenomena in GLSCFBs that the bubble size is smaller and its distribution is narrower in the circulating fluidized beds compared to the conventional three-phase fluidized beds. This can be attributed to the increase of turbulence in the circulating beds owing to the high velocity of the liquid phase to circulate the solid particles. The bubble rising velocity was found to increase with increasing Ug (gas velocity) and Ul (liquid velocity), but it decreased slightly with increasing Gs (solids circulation rate). It has also been observed that the mean value of LV (bubble chord length) increased with Ug but decreased progressively with increasing Ul or Gs, whereas the bubble frequency (FB) increased with increasing UG, UL, or Gs. It is generally understood that the turbulence which is effective to decrease the bubble size and to increase the bubble frequency for a given gas velocity, increases with increasing liquid velocity and solids circulation rate. However, with increase in liquid velocity, upward force acting on the rising bubbles increases and thus the bubble rising velocity also increases. Therefore, increased liquid velocity eventually results in a slight increase of bubble size in spite of decreasing the size since the bubble rising velocity is proportional to the size of bubbles. They showed that the overall heat transfer coefficient increased with increasing Ug and Gs, but it appeared to be unaltered by the effect of Ul. The following correlation was proposed to relate heat transfer coefficient with the flow velocities: h ) 2776Ul0.0317Ug0.0799Gs0.0654

(17)

The values of h in circulating fluidized beds were found to be somewhat higher than those in the conventional three-phase fluidized beds.11 4.4. Studies on Modeling. A closed Eulerian-Eulerian-La grangian model for simulating gas-liquid-solid three-phase local flow properties of a GLSCFB has been recently developed by Cao et al.27 with the combination of the two-fluid model (TFM) and distinct element method (DEM). This model was based on the fundamental equations of fluid mechanics. The motion of particles was described in the Lagrangian coordinates, while the gas phase and the liquid phase were described within the Eulerian coordinates. The effect of particles on the gas phase and liquid phase was achieved by implementing momentum exchange terms into the TFM. Thereafter, the model was

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Figure 17. Comparison between the CFD predicted and the experimental local liquid velocity radial profiles for various superficial gas velocities, 0.6 m above gas distributor, solids circulation rate 4 kg/m2 · s, superficial liquid velocity 4.0 cm/s, air-SCMC-styrene resin system. Adapted from ref 27.

validated with their in-house experimental data sets for radial distributions of local phase holdups and local liquid velocity which were obtained by the microconductivity probe technique and the electrolyte tracer technique for the system studied with styrene resin as solid phase and 0.05 and 0.20 (wt %) carboxymethyl cellulose sodium (CMCS) as liquid phases in a GLSCFB riser. According to their studies, in a GLSCFB riser, the motion of a solid particle at any moment in time is random but there are some patterns when the velocity profiles are time averaged. The results (Figure 17) revealed a general trend of increasing local liquid velocity in the reactor center with increasing superficial gas velocity. At the same time, the nonuniformity of the liquid velocity distribution in the radial direction is found to be increased with increasing gas velocity. This observation was explained by the fact that the increase of gas velocity leads to more bubbles and has stronger turbulence in the bed. The effect of the solids circulation rate on the radial distribution of the local liquid velocity can also be obtained from this study (Figure 18). With an increasing solids circulation rate, the liquid velocity at the center region of the bed was found to decrease and the radial distribution of the liquid velocity was trying to be more uniform. This can occur because of an increase in solids holdup which can uphold the contacting and splitting probability among gas, liquid, and solid phases. The variation of local solids holdup with various superficial gas velocities was also studied (Figure 19), which showed that there was a significant difference in the local solids holdup near the reactor center region as compared to those near the wall region. In the center region, the local solids holdup was more uniform. This can also imply that the interaction between the bubbles and the solid particles causes significant bubble break-up, resulting in smaller sizes and more uniform bubbles in the center region. Near the wall region, the local solids holdup profile showed a prevalent peak of the distribution. It can be noted that the prediction of the radial distribution of the local liquid velocity at lower liquid velocities by this model cannot account for the flow regimes encountered in a GLSCFB riser. However, all these CFD results were in good qualitative agreement with the previous studies as well as the experimental observations, thus forming a sound basis for further investigation on the modeling of GLSCFBs which can hold quantitative information as well.

Figure 18. Comparison between the CFD predicted and the experimental local liquid velocity radial profiles for various solids circulation rates, 0.6 m above gas distributor, superficial gas velocity 4 cm/s, superficial liquid velocity 5.0 cm/s, air-0.05% SCMC-styrene resin system. Adapted from ref 27.

Figure 19. Comparison between the CFD predicted and the experimental local solids holdup radial profiles for different superficial gas velocities, 0.6 m above the gas distributor, solids circulation rate 6 kg/m2 · s, superficial liquid velocity 2.0 cm/s, air-0.05% SCMC-styrene resin system. Adapted from ref 27.

5. Applications of (G)-LSCFBs Significantly improved interfacial contact efficiency, as well as good mixing and efficient heat and mass transfer have made fluidized bed reactors a unique choice in many industries. (G)-LSCFBs emerged as very promising multiphase contactors not only for their advantages but also their controllable residence time and operating conditions. Applications and potential applications of (G)-LSCFBs are increasingly being reported in industrial fields such as wastewater treatment, protein separation, pharmaceutical, and petrochemical.23 The (G)-LSCFB reactor system has two important wings with distinct compartments, namely, the riser and downer. These two zones can have completely different characteristics of retention time distribution and sheer stress for interphase interactions and therefore are particularly suitable for several biological processes. Another unique feature which adds more value to this system is the independent control of the downer. The following table (Table 4) summarizes the characteristics of the riser and downer. 5.1. Wastewater Treatment. A new wastewater treatment process based on the very promising “fixed-film” BNR (biologi-

Ind. Eng. Chem. Res., Vol. 48, No. 17, 2009 Table 4. Comparison of Characteristics between the Riser and Downer1 mean retention time (MRT) zone

liquid

solid

ratio of MRT (liquid) and MRT (solid)

shear stress

riser downer

short long

short long

∼1 variable

high low

cal nutrient removal) technology has been developed in recent years, where the riser is used for the anoxic process and the downer for the aerobic process, with 500-1000 µm porous particles loaded with biofilm circulating between the downer and riser.8 This successful CFBBR (circulating fluidized bed bioreactor) or LSCFBBR technology has gone through laboratory-scale study4,7 as well as pilot-scale study8 and is now ready for commercialization. The results of the lab-scale evaluation of CFBBR7 running in anoxic-aerobic mode using synthetic wastewater and employing lava rock as a carrier media demonstrated that use of the liquid-solid circulating fluidized bed system (LSCFB) is not only technically feasible but also very advantageous for municipal wastewater treatment. Owing to enhanced mass transfer, COD removal efficiency, in the range of 86.6-98.5%, and ammonia removals, in excess of 99%, was asserted for certain operating conditions within very favorable time limit in comparison with conventional systems where to achieve similar performance it used to take much more time. It has also been shown that this process offers favorable nitrification and denitrification kinetics in terms of volumetric nitrification rate and volumetric denitrification rate, respectively. Afterward, researchers from the same group4,17,46 further reported the performance of the circulating fluidized bed bioreactor (CFBBR) as anoxic and aerobic beds for the simultaneous removal of carbon, nitrogen, and phosphorus from municipal wastewater. They achieved relatively higher overall nitrogen removal efficiencies without bioparticle recirculation; however, better biological P removal was realized with bioparticle recirculation. It was also noticed that denitrification occurred simultaneously with P release in the anoxic bed, while nitrification took place in the aerobic bed. Negligible ammonia removal was observed in the anoxic bed. In terms of chemical oxygen demand (COD) and total suspended solids (TSS) removals, the CFB bioreactor was found to respond very positively. More recently, Chowdhury et al.8,47 reported the results of their 5000 L/day pilot-scale study at London city’s wastewater treatment plant in Canada. The pilot-scale LSCFBBR has a 20 cm i.d. and 6.0 m high riser column and a 50 cm i.d. and 3.0 m high downer column. The system runs at a feeding rate of 5000 L/d with a corresponding organic loading rate (OLR) of 3.65 kg COD/m3 · d and a nitrogen loading rate (NLR) of 0.37 kg N/m3 · d. The results show the that the system is able to remove 92%, 89%, and 76% of the influent organic, nitrogen, and phosphorus, respectively. The effluent generated by this system was characterized by

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