High Solid-Flux Concurrent Conveying Flow Realized by Coupling a

Oct 28, 2008 - Parametric investigation further clarified that the solid circulation rate and the local solid holdup at the riser bottom of the newly ...
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Ind. Eng. Chem. Res. 2008, 47, 9703–9708

9703

High Solid-Flux Concurrent Conveying Flow Realized by Coupling a Moving Bed to the Bottom Section of a Riser Xinhua Liu,† Xin Cui,†,‡ Guang Sun,†,§ Guogang Sun,§ Toshiyuki Suda,| and Guangwen Xu*,† State Key Laboratory of Multi-Phase Complex System, Institute of Process Engineering, CAS, Beijing 100190, China, School of Chemical Engineering, Shenyang Institute of Chemical Technology, Shenyang 110142, China, Faculty of Chemical Engineering, China UniVersity Of Petroleum (Beijing), Beijing 102249, China, and Research Laboratory, IHI Corporation, Ltd. (IHI), Isogo-ku, Yokohama 235-8501, Japan

Gasification of coal and biomass is in pursuit of the technologies based on dual bed combination and a highdensity transport bed. Dual fluidized bed gasification (DFBG) relies on rapidly circulated particles between its combustor and gasifier to provide the endothermic heat required by the gasification. High-density transport bed gasification (HTBG) has to work with a high solid flux and a high particle density inside its gasifier so as to increase the heat reserve in the bed and to suppress tar evolution there. The idea of coupling a moving bed to the bottom section of the riser of a circulating fluidized bed (CFB) was proposed to realize the desired high solid-flux conveying flow inside the riser. Experiments in a 12-m high and 90 mm i.d. riser of the newly configured CFB demonstrated that at superficial gas velocities of about 9.6 m/s, a solid circulation rate as high as 370 kg/(m2 s) and an average solid holdup of about 0.12 in the bottom section of the riser were readily achieved simultaneously for the silica sand particles of 378 µm in Sauter mean diameter. Parametric investigation further clarified that the solid circulation rate and the local solid holdup at the riser bottom of the newly configured CFB were highly dependent on the moving bed aeration and the primary gas velocity of the riser, whereas changing the solid inventory in the system did not greatly affect those two variables. Adoption of a secondary air injection into the riser enabled adjustment of the solid circulation rate within a certain range, showing essentially a complementary means for controlling the gas-solid flow inside the riser of the newly configured CFB. 1. Introduction Dual fluidized bed gasification (DFBG) and high-density transport bed gasification (HTBG) are two newly developed coal and biomass gasification processes based on circulating fluidized bed (CFB) technology. An autothermal DFBG system generally consists of a highly endothermal fluidized bed pyrolyzer/gasifier of fuel and an exothermal fluidized bed combustor of char.1 The isolation of fuel pyrolysis/gasification from char combustion in the DFBG system avoids the dilution of product gas by nitrogen of combustion air, which makes it possible to produce high heating value gas in this system even using air as a gasifying agent.2,3 The DFBG is also capable of treating biomass or coal with high water content.4 The implementation of these merits relies on the transfer of heat from the combustor to the pyrolyzer/gasifier, which has to resort to sufficiently circulating a kind of solid heat carrier between the two reactors. The higher the solid circulation rate, the smaller the temperature difference between the two reactors. Thus, the operation and control of the DFBG system becomes easier. The HTBG technology generally has the advantages of modest reaction temperature and high fuel adaptability.5,6 In the HTBG system, a large particle flux and a high solid density in the transport bed increase the heat reserve in the gasifier and provide more active reaction surfaces for the cracking of tars. These in turn enhance the reaction rate of the gasification and suppress the tar evolution with the product gas. Consequently, the buildup and the control * To whom correspondence should be addressed. Tel./Fax: +8610-62550075. E-mail address: [email protected]. † Institute of Process Engineering. ‡ Shenyang Institute of Chemical Technology. § China University Of Petroleum (Beijing). | IHI Corporation.

of a high solid-flux flow in a CFB-based system are the indispensable premises for developing the DFBG and HTBG technologies. The solid circulation rate of a CFB system is subject to the imposing pressure, the pressure head of the returning leg accompanying the conveying column or the riser. In a traditional CFB with a nonmechanical loop seal at its riser bottom, a solid circulation rate of higher than about 200 kg/(m2 s), especially for Geldart B particles, is difficult to realize7 because the total pressure head of the standpipe can not be exerted directly and completely on the gas-solid flow inside the riser due to the pressure drop loss of the loop seal. Bi and Zhu8 proposed that a larger solid circulation rate and a higher particle concentration in the riser of a CFB can be achieved by involving a second riser of large cross-sectional area to lift the solids from the first riser to a higher level. Pugsley et al.9 developed a nonmechanical solid feeder for a CFB riser to achieve very high and wellcontrolled solid mass fluxes, but their design involved the replacement of the traditional standpipe or L-valve with an aerated annular bed of the solids and an innovative radial gas distributor for the riser. Li et al.10 operated a riser under high solid fluxes by coupling another downer-riser system to the CFB riser. However, all these methods of achieving high particle fluxes are either impractical or not suited enough to the above-mentioned DFBG and HTBG technologies. In this study, a moving bed coupled to the bottom section of the riser of a traditional CFB was expected to increase the feed pressure head of the returning loop and thereby to increase the solid circulation rate and the local solid holdup in the riser. For this purpose, the loop seal at the riser bottom in the traditional CFB had to be leveled up in the new design to allow an installation space for the moving bed. The involved flow

10.1021/ie801041g CCC: $40.75  2008 American Chemical Society Published on Web 10/29/2008

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Figure 1. Schematic of the experimental apparatus.

mechanism is that in the traditional CFB the pressure head induced by the particle bed in the standpipe is partially counteracted by the bottom support of the loop seal and thereby can not be converted completely into the effective driving force to push the particles flowing from the standpipe to the riser. Meanwhile, leveling up the loop seal and coupling the moving bed between the seal valve and the riser bottom hope to overcome the abovementioned shortcoming, thus making it possible to convert completely the pressure head of the moving bed into the desired feed driving force. As a consequence, both a high solid circulation rate and a high particle density are hopefully achieved simultaneously in such a newly configured CFB. The present article is devoted to demonstrating the preceding idea with experiments performed in a laboratory-scale solid conveying rig. The influences of major operating parameters, including moving bed aeration, primary gas velocity, and secondary air injection in the riser, will be also tested to show how to control the newly configured CFB under high solid fluxes. 2. Experimental Details Figure 1 shows a sketch of the employed experimental rig, a laboratory-scale circulating fluidized bed (CFB) reconfigured according to the newly proposed idea. The newly configured

Table 1. Configuration Characteristics of the Experimental CFB

i.d. (m) height (m)

riser

standpipe

moving bed

0.09 12.40

0.08 7.40

0.12 4.40 loop seal

length (m) width (m) height (m) width division ratio (-) opening area (m2)

0.35 0.20 0.60 1.69 0.02–0.06

CFB is made of plexiglas and mainly consists of a circular riser, a circular standpipe, a quadrate loop seal, and a circular moving bed. Table 1 summarizes the configuration characteristics of the CFB. The riser is 90 mm in i.d. and 12.4 m high, while the employed moving bed is 120 mm in i.d. and 4.4 m high. A throat-type bottom was adopted in the moving bed to make the particles possibly accumulate inside the bed. The loop seal in the tested CFB was mounted at a height of about 5 m above the riser bottom, which differs distinctively from the traditional CFBs. The granular flow passage at the bottom of the loop seal, the opening between the lower end of the central baffle and the loop seal distributor, was designed to be adjustable within a certain range in this design.

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Figure 2. Variation range of measured Gs in the riser under the operating conditions tested.

Figure 3. Axial profile of εs in the riser under the typical operating conditions.

As one can see from Figure 1, the upflow-side chamber of the loop seal is connected to the moving bed, while the downflow-side room is to the standpipe. More than 20 pressure taps were installed around the entire circulation loop. In order to allow the loop seal fluidizing air to flow into the riser, via a hose pipe the tap PF at the top of the loop seal was connected to the tap PB at the riser wall in the newly configured CFB. However, if the loop seal is replaced by a fluidized-bed pyrolyzer/gasifier, the gas effluence from the tap PF, as the product gas of the reactor, can also be introduced directly into the downstream line instead of being injected into the riser. All pressure taps were mounted directly on the column wall and plugged with a screen to prevent backflow of particles. Pressure transducers (CGYL-3018) were connected to the pressure taps to measure the pressures at different locations. Electric signals from the transducers were sent to a personal computer for record at a frequency of 50 Hz. Apparent solid holdup εs in the conveying column, if neglecting the effect of solids acceleration and the wall effect, can be calculated directly from the measured pressure drop through

Table 2. Particle Properties and Variation Ranges of Major Operating Parameters

εs )

∆P ⁄ ∆H (Fs - Fg)g

(1)

The solid circulation rate Gs was measured using a butterfly valve fixed at the standpipe. The loop seal aeration gas velocity Ugd in the downflow side was kept at zero in all experimental runs, because the amount of the loop seal fluidizing air (Ugu) leaking into the standpipe was large enough to maintain the particles there flowing from the downflow side to the upflow side of the seal valve. A kind of silica sand belonging to group B particles according to the Geldart classification was used as the fluidizing material. Table 2 lists the main properties of the fluidizing particles and the variation ranges of major operating parameters investigated in the article. 3. Results and Discussion 3.1. Feasibility Demonstration. As mentioned above, for Geldart B fluidizing particles, the largest solid circulation rate Gs realizable in the conventional CFBs is generally not higher than about 200 kg/(m2 s). However, as shown in Figure 2, in the newly configured CFB the solid circulation rate Gs is adjustable within the range of 100 to about 400 kg/(m2 s), and a much higher Gs than the upper limit of Gs realizable in the conventional CFBs is readily achieved. The maximal Gs of about 370 kg/(m2 s) created at the superficial gas velocity of about 9.6 m/s in the newly configured CFB clarified that the new CFB system was truly operated under high solid fluxes. Therefore,

Particle Properties dp [µm]

Fs [kg/m3]

Fb [kg/m3]

Umf [m/s]

Ut [m/s]

378

2600

1470

0.11

2.82

Variation Ranges of Operating Parameters Ugp [m/s]

Ugs [m/s]

Ugm [m/s]

Ugu [m/s]

Im [kg]

5.6-11.0

0.0–3.5

0.04–0.15

0.11–0.15

130–160

it may be deduced reasonably that the introduced moving bed contributes much to the increase of Gs in the newly configured CFB. Figure 3 shows a typical axial profile of the apparent solid holdup εs in the riser of the newly configured CFB under tested conditions. Obviously, the average solid holdup εs,0-4m in the bottom section of the riser, especially for the cases of the moving bed aeration gas velocity Ugm ) 0.05 m/s, is up to about 0.120. While, the solid circulation rate Gs is also high up to 369 kg/ (m2 s) at the superficial gas velocity Ugp ) 9.61 m/s. This experimental phenomenon demonstrates that the new bed configuration truly created the desired conveying flow with a high particle flux and a high solid density bottom region in the riser according to the definition given by Grace et al. on a high density circulating fluidized bed.11 Because the top of the standpipe of the newly configured CFB was always connected to an induced draft fun and the pressure there was kept constant, the pressure drop across the entire returning loop is a single valued function of the pressure PG at the bottom of the moving bed. Accordingly, the pressure PG at the bottom of the moving bed can be taken as a fair index to the driving force for the particles flowing from the returning loop to the riser in the new bed configuration. The higher the pressure PG at the moving bed bottom, the greater the driving force of the returning loop. The experimental data exemplified in Figure 4 show that the solid circulation rate Gs in the newly configured CFB increases proportionally with increasing PG. Because the coupling of the moving bed in the newly configured CFB raises the pressure drop across the returning loop, the solid circulation rate Gs increases high up to about twice that achievable in the traditional CFBs for the same fluidizing particles. This reveals actually the necessity and feasibility of integrating the moving bed into the bottom section of the riser in order to realize the desired high solid-flux cocurrent conveying flow. 3.2. Parametric Investigation. As shown in Figure 3 in the above, increasing the primary gas velocity Ugp from 7.86 to 9.61 m/s at a constant moving bed aeration gas velocity Ugm )

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Figure 4. Variation of Gs with the pressure at the bottom of the moving bed under the operating conditions tested.

Figure 5. Variation of Gs with Ugp at different Im and As.

0.04 m/s leads to an increase of the solid circulation rate Gs from 177 to 221 kg/(m2 s), but a slight decrease of the average solid holdup εs,0-4m in the riser bottom from 0.067 to 0.052. Actually, an increase of Ugp must result in an increase of Gs within a certain range because of an increased upflow velocity of the particles inside the riser, as shown in Figure 5. But, varying the solid inventory Im from 130 to 160 kg contributes little to the increase of Gs since the local solid holdup at the top of the riser changes little. Decreasing the bottom opening area As of the loop seal from 0.06 to 0.02 m2 also has no obvious effect on Gs, because the tested minimal As ) 0.02 m2 is still much greater than the cross-sectional area (0.011 m2) of the moving bed. Much more work may be needed on this issue in the future research. Figure 3 in the above also shows that increasing the moving bed aeration gas velocity Ugm from 0.04 to 0.05 m/s results in an obvious increase of the apparent solid holdup εs in the riser despite an increase in Ugp from 7.86 to 9.61 m/s. Because the solid flow rate from the moving bed to the riser may be mainly decided by the variation of Ugm in the newly configured CFB, the effect of Ugp on εs is completely subject to that of Ugm, making εs highly sensitive to the variation of Ugm rather than of Ugp. The increase of εs with increasing Ugm, especially at the top of the riser, also accounts reasonably for the increase of Gs with increasing Ugm at a fixed Ugp ) 8.73 m/s, as shown in Figure 6. In theory, the increase of the fluidity of the particles reserved in the moving bed due to an increase of Ugm must lead to a higher imposing pressure for the gas-solid flow inside the riser, thus to a higher Gs. But Ugm should not be much greater than the minimum fluidization velocity Umf of the employed particles in order to avoid the leakage of the fluidizing air from the riser to the moving bed. Experimental observation also showed that increasing Ugm beyond Umf () 0.11 m/s according to the Wen-Yu equation) of the silica sand particles was very

Figure 6. Variation of Gs with Ugm at different Ugu values.

likely to induce so-called plug flow in the standpipe and further to terminate normal operation of the newly configured CFB. Because the variation of the loop seal fluidizing gas velocity Ugu within the range of 0.07-0.11 m/s has little effect on the effective pressure drop across the moving bed at a fixed Ugm, the solid circulation rate Gs shown in Figure 6 also varies little with Ugu under the tested operating conditions. Li et al.12 and Xu and Gao13 took a constant εs in the top section of the riser as an indication for the point of saturation carrying capacity. Bi et al.14 reported that the saturation carrying capacity at a specified Ugp and zero Ugs should be a constant. However, Figure 3 in the above indicates that the apparent solid holdup εs in the top section of the riser of the newly configured CFB increases clearly with increasing Ugm at Ugm < Umf, implying that the solid circulation rate Gs in the riser was actually greater than the saturation carrying capacity corresponding to the superficial gas velocity. The increase of Gs with raising Ugm at a fixed Ugp ) 8.73 m/s shown in Figure 6 further consolidates the above observation and accounts reasonably for why the newly configured CFB can be operated in a high solid flux state under the operating conditions. If secondary air is injected into the riser at a fixed total superficial gas velocity Ugt and the moving bed aeration gas velocity Ugm in the newly configured CFB, the particles must become more and more concentrated into the lower part of the riser below the injection point because of a decreased Ugp. In fact, as shown in Figure 3 above, reducing Ugp from 9.61 to 6.99 m/s but raising the secondary gas velocity Ugs from 0 to 2.62 m/s correspondingly at a constant Ugm ) 0.05 m/s leads to an obvious increase of the local solid holdup from 0.32 to about 0.38 at the riser bottom. Because a lowered Ugp must transport fewer particles into the bed section above the secondary air injection point from the bottom section of the riser, the injection of the secondary air at a fixed Ugt leads to an obvious decrease of Gs in the newly configured CFB. A secondary air injection was reported to be able to reduce solid flow rate falling along the riser wall,15 while increasing as well the upward moving rate of the particles just above the secondary air injection point. Figure 7 shows that increasing Ugs at a constant Ugp of 7.86 m/s slightly increases Gs. But, this trend becomes much evident when the height Hgs of the secondary air injection point decreases from 3.5 to 2.0 m, because the particles falling along the riser wall increase at the lower Hgs and the particle flux coming from the lower bed section is also somehow higher at this lower location. At a fixed total superficial gas velocity Ugt () Ugp + Ugs) of 9.61 m/s, increasing the ratio Ugs/Ugp of the secondary to primary gas velocity from zero to about 0.37 implicates an increase of Ugs from 0 to 2.62 m/s, but a decrease of Ugp from 9.61 to 6.99 m/s. As a consequence, the measured Gs decreases gradually

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Figure 7. Variation of Gs with the secondary air ratio Ugs/Ugp at different Hgs.

Figure 8. Pressure loops around the newly configured CFB under the typical operating conditions.

with increasing Ugs/Ugp in Figure 7, though the secondary air injection may help to prevent the particles just above the injection point from falling into the bottom section of the riser. Being subject to the same reason as for the influence of Hgs clarified in the Ugp-fixed case, the decreasing degree of Gs with raising Ugs/Ugp is also more or less much obvious at a greater Hgs. 3.3. Further Clarification. Figure 8 shows the pressure loops around the entire newly configured CFB for the typical experiments. The dashed lines represent the axial pressure profiles in the entire returning loop including the standpipe, the loop seal, and the moving bed. As demarcated by the moving bed height Hm in the figure, the pressure head to drive the particles to flow from the returning loop into the riser comes mainly from two parts: the pressure drops across the bottom section of the standpipe and the moving bed. The pressure drop from the moving bed takes more than about 30% of the total pressure drop across the entire returning loop at Ugp ) 9.61 m/s and Ugm ) 0.05 m/s, indicating that the moving bed definitely contributes much to a larger Gs and a higher εs in the newly configured CFB than in the traditional CFBs. This further consolidates the proposed idea that integrating the moving bed into the bottom section of the riser helps to raise the driving force for the particles to flow from the returning loop to the riser. The experimental results exemplified in Figure 8 also indicate that a large pressure difference ∆Ptop always exists between the riser top and the primary cyclone in all experimental runs, which can be mainly attributed to the resistance losses across the riser exit ∆Pin, the elbow ∆Pelbow, and the cyclone entrance ∆Pout. But, the tested pressure drop ∆Pin across the riser exit only accounts for about 10% of total the pressure difference under

Figure 9. Variation of ∆Pf with Gs under the typical operating conditions.

the operating conditions because of the smooth exit configuration of the riser in the newly configured CFB. As mentioned above, the loop seal in the newly configured CFB should be leveled up from the bottom of the riser because of the introduction of the moving bed. By characterizing the variation of the pressure drop ∆Pf across the bottom opening of the loop seal with Gs, the effect of elevated the loop seal on its original seal function and the operation of the riser will be clarified further in following texts. If the average solid holdups in the loop seal and the bottom section of the standpipe are supposed to equal εmf of the solids, as shown in the partially enlarged detail of the loop seal configuration in Figure 9, the pressure drop ∆Psd imposed by the solids in the standpipe and the downflow-side bed of the loop seal can be estimated as ∆Psd ) (Fs - Fg)gHsdεmf ∝ Hsd

(2)

while the pressure drop ∆Psu across the upflow-side bed of the loop seal is estimated as ∆Psu ) (Fs - Fg)gHsuεmf ) const

(3)

Thus, since the height of the particle bed accumulated in the standpipe was found to seldom exceed axial location of pressure tap PD5 in the experiments, the pressure drop ∆Pf across the bottom opening of the loop seal can be determined from the experimental data as ∆Pf ) PD5 + ∆Psd - PF - ∆Psu

(4)

The calculation results plotted in Figure 9 show that ∆Pf increases obviously with increasing Gs regardless of the variation of the bottom opening area As of the loop seal. This increase of ∆Pf, however, may not alleviate the imposing pressure of the moving bed to the gas-solid flow inside the riser because the loop seal was elevated to above the moving bed in the newly configured CFB. On the contrary, it just reflects the selfadaptability of the seal function of the loop seal in the newly configured CFB and further demonstrates the technical feasibility of coupling the moving bed to the bottom section of the riser, because a higher Gs generally implies the requirement of a larger pressure drop across the returning loop. But, it is worth noting that in this study ∆Pf could not be taken as a simple function of the average orifice flow rate of the particles as expected. The nonfluidized particles near the bottom orifice of the loop seal caused by the use of zero the loop seal aeration gas velocity Ugd in all experimental runs may be responsible for this result. 4. Conclusions In a circulating fluidized bed (CFB), in order to increase the driving force for the particles to flow from the returning loop

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to the riser and thus to realize a high solid-flux concurrent conveying flow in the riser, a moving bed was proposed to be coupled to the bottom section of the riser and the loop seal of the CFB was elevated to above the moving bed correspondingly. In an experimental CFB system reconfigured according to the above idea, this article examined the possibility to achieve the desired high particle fluxes and high solid holdups in the riser and further investigated the influences of the major operating parameters on gas-solid flow inside the riser. (1) For the silica sand particles of 378 µm in Sauter mean diameter, solid circulation rates of up to 370 kg/(m2 s) and apparent solid holdups of about 0.12 in the bottom section of the riser were readily achieved at superficial gas velocities of about 9.6 m/s in the newly configured CFB. The pressure drop caused by the moving bed was found to account for about 30% of the total pressure drop across the returning loop and can act directly on the gas-solid flow inside the riser of the newly configured CFB, showing why the newly configured CFB enabled a high particle flux and a high solid holdup in the riser. (2) The gas-solid flow inside the riser of the newly configured CFB is exclusively subject to the moving bed aeration and primary gas velocity of the riser but can only be slightly affected by both the loop seal fluidizing gas velocity and the solid inventory in the system. Even a little increase of the moving bed aeration gas velocity must increase greatly both the solid circulation rate and the local solid holdup in the riser. A larger primary gas velocity allows a higher solid circulation rate but a lower local solid holdup in the riser. Under a constant total gas velocity, splitting the primary gas flow to form a secondary air injection decreases the solid circulation rate slightly. At a fixed primary gas velocity, elevating the secondary gas velocity, however, increases the solid circulation rate to a certain degree. (3) The major manipulation parameters for the gas-solid flow inside the riser of the newly configured CFB are the moving bed aeration gas velocity and the primary gas velocity of the riser. Only modest control of the solid circulation rate is possible by adoption of a secondary air injection into the riser. Acknowledgment The authors of the Chinese side are grateful for the financial support of the national 863 project under the number 2006AA05A103 and of the NSFC project under the number 20606034. Nomenclature As ) cross-sectional area of bottom opening of the loop seal (m2) dp ) particle diameter (µm) g ) gravitational acceleration (m2/s) Gs ) solid circulation rate (kg/m2s) Hgs ) height of secondary air injection point (m) Hm ) height of the moving bed (m) Hsd ) total height of the particle bed in the standpipe and the downflow-side bed of the loop seal (m) Hsu ) height of the particle bed in the upflow-side bed of the loop seal (m), Hsu ) 0.45 m Im ) solid inventory (kg) P ) static pressure (kPa) Ugu ) fluidizing gas velocity in the upflow side of the loop seal (m/s) Ugm ) moving bed aeration gas velocity (m/s)

Ugp ) primary gas velocity in the riser (m/s) Ugd ) aeration gas velocity in the downflow side of the loop seal (m/s) Ugs ) secondary gas velocity in the riser (m/s) Ugt ) total superficial gas velocity in the riser (m/s) Umf ) minimum fluidization velocity of silica sand particles (m/s) Ut ) terminal velocity of silica sand particles (m/s) ∆P/∆H ) pressure gradient along the column (kPa/m) ∆Pf ) pressure drop across the bottom opening of the loop seal (kPa) ∆Psd ) pressure drop across the standpipe and the downflow-side bed of the loop seal (kPa) ∆Psu ) pressure drop across the upflow-side bed of the loop seal (kPa) εs,0-4m ) average solid holdup in the bottom section (0-4 m) of the riser (dimensionless) εmf ) solid holdup at the minimum fluidization (dimensionless), εmf ) 0.55 εs ) apparent solid holdup (dimensionless) Fb ) bulk density of silica sand (kg/m3) Fg ) gas density (kg/m3) Fs ) real density of silica sand (kg/m3)

Literature Cited (1) Corella, J.; Toledo, J. M.; Molina, G. A review on dual fluidizedbed biomass gasifiers. Ind. Eng. Chem. Res. 2007, 46, 6831–6839. (2) Pfeifer, C.; Rauch, R.; Hofbauer, H. In-bed catalytic tar reduction in a dual fluidized bed biomass steam gasifier. Ind. Eng. Chem. Res. 2004, 43, 1634–1640. (3) Xu, G.; Murakami, T.; Suda, T.; Matsuzawa, Y.; Tani, H. The superior technical choice for dual fluidized bed gasification. Ind. Eng. Chem. Res. 2006, 45, 2281–2286. (4) Xu, G.; Murakami, T.; Suda, T.; Tani, H.; Mito, Y. Efficient gasification of wet biomass residue to produce middle caloric gas Particuology 2008, 6, 376-382. (5) Smith, P. V.; Salazar, N.; Davis B. M.; Leonard, R.; Nelson, J. M.; Breault, R. W. KBR transport gasifier. Gasification Technologies: San Francisco, 2005. (6) Abughazaleh, J.; Salazar, N.; Ariyapadi, S. Optimising hydrogen and chemicals production from the transport gasifier. International Pittsburgh Coal Conference, Johannesburg, September 10-14, 2007. (7) Zhu, J. X.; Bi, H. T. Distinctions between low density and high density circulating fluidized beds. Can. J. Chem. Eng. 1995, 743, 644– 649. (8) Bi, H. T.; Zhu, J. Static instability analysis of circulating fluidized beds and concept of high-density risers. AIChE J. 1993, 39, 1272–1280. (9) Pugsley, T. S.; Milne, B. J.; Berruti, F. An innovative non-mechanical solids feeder for high solids mass fluxes in circulating fluidized bed risers. Powder Technol. 1996, 88, 123–131. (10) Li, Z. Q.; Wu, C. N.; Wei, F.; Jin, Y. Experimental study of highdensity gas-solids flow in a new coupled circulating fluidized bed. Powder Technol. 2004, 139, 214–220. (11) Grace, J. R.; Issangya, A. S.; Bai, D. Situating the high-density circulating fluidized bed. AIChE J. 1999, 45 (10), 2108–2116. (12) Li, J.; Tang, Y.; Kwauk, M. Axial voidage profiles of fast fluidized bed in different flow regions Circulating Fluidized Bed Technology; Basu, P.; Large, J. F., Eds.; Pergamon: Oxford, 1988. (13) Xu, G.; Gao, S. Necessary parameters for specifying the hydrodynamics of circulating fluidized riserssa review and reiteration. Powder Technol. 2003, 137, 63–76. (14) Bi, H. T.; Grace, J. R.; Zhu, J. X. Types of choking in vertical pneumatic systems. Int. J. Multiphase Flow. 1993, 19 (6), 1077–1092. (15) Marzocchella, A.; Arena, U. Hydrodynamics of a circulating fluidized bed operated with different secondary air injection devices. Powder Technol. 1996, 87, 185–191.

ReceiVed for reView July 4, 2008 ReVised manuscript receiVed September 10, 2008 Accepted September 19, 2008 IE801041G