Novel Concept of Fluidized Bed Reactor Design for Highly Efficient

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A novel concept of fluidized bed reactor design for high efficient internal heat transfer Xiao Wang, and Hui Si Ind. Eng. Chem. Res., Just Accepted Manuscript • Publication Date (Web): 12 Aug 2016 Downloaded from http://pubs.acs.org on August 13, 2016

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Industrial & Engineering Chemistry Research

A novel concept of fluidized bed reactor design for high efficient internal heat transfer Xiao Wang, Hui Si* School of Technology, Beijing Forestry University, No.35 Tsinghua East Road Haidian District, Beijing, P.R.China ABSTRACT: A novel double pipe fluidized beds (DPFB) reactor was used in the present study to supply heat energy during fast pyrolysis. The steady state heat transfer model of DPFB was developed where classical equations of heat transfer inside fluidized bed and that of wall-to-bed was considered. The sensitivity analysis showed that contact time was critical parameter to overall heat transfer coefficient. The model provided good predicting accuracy on temperature calculating by deviating from measured data within +20%. The simulation of DPFB was conducted based on the model and results showed that the heat exchange mainly occurred in the bed phase. Both reducing contact time and increasing Reynolds number would improve heat transfer significantly. The effect of static bed height on outlet temperatures of DPFB could be observed under high inlet temperature difference, which illustrated the importance of annular fluidized bed in heat transfer enhancement. keywords: annular fluidized bed; internal heat transfer; reactor design; modeling 1. INTRODUCTION Fluidized beds find a widespread application in thermo-chemical processes such as pyrolysis for the excellent heat transfer properties. Since pyrolysis is an endothermic reaction, heat need to be supplied during pyrolysis reaction1, leading to many studies focusing on the development of pyrolysis reactor in recent decades. It has been found that the form of wall-to-bed heat transfer in fluidized bed is an effective approach to deliver energy from outside to the reaction region. There are mainly two approaches to exploit the wall-to-bed heat transfer. One is having heating tubes immersed in the bed which is proven to be effective according to literatures2-4. However, these heating tubes are attrited seriously in fluidized beds and have an impact on the hydrodynamics of fluidized bed. Another way is equipped the fluidized bed column with electrical jacketed heaters which was widely used in both experimental studies and industrial applications5, 6. However, it appears that electrical heating is not effective compared with combustion heating. In addition, the waste heat could hardly be utilized in the electrical heating system which made it not energy efficient. To solve this problem, double pipe fluidized beds (DPFB) reactor was come up which utilized the heat of combustion gas instead of electric energy. In this reactor, the conventional fluidized bed column was enclosed with another fluidized bed with larger diameter, resulting two fluidized beds being divided by a vessel wall. If the temperature of annular bed was higher than the cylindrical one, heat would transport to the reaction region to replenish the energy loss due to the biomass conversion. The high temperature in annular bed could be achieved by using a combustor instead of electrical heater, through which the internal heat transfer would be significantly improved. From this point of view, the annular bed of DPFB would play an important role in heat 1

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transfer since it determined the heat transfer coefficient. It was also found that the annular bed was used widely in industrial applications. Cheng et al.7 carried out an analysis for a fully-developed, forced convective flow through a packed sphere bed between concentric cylinders maintained at different temperatures. Doyle III et al.8 developed a model to account for pinning behavior in an annular moving bed reactor of finite length. Raupp et al.9 studied effects of system parameters on process performance for a gas–solid lamp-in-tube annular photocatalytic oxidation reactor where the annular space was filled with photocatalyst coated packing. Yasar Demirel10 analyzed the convective heat transfer in an annular packed bed heated by constant heat fluxes and verified the fully developed temperature profile and the plug flow conditions with experimental data. Collin et.al11 conducted the study based on laboratory and the pilot annular fluidized bed, and characterized the flow pattern in the system. Recently, Qiu et al.12 proposed a new type of annular combustion chamber to improve the secondary air penetration and provided adequate space to arrange the heat transfer surface. However, there were few studies that reported the application of such a DPFB reactor. Therefore, some fundamental investigations had to be carried out to help us understanding the features of this reactor, especially the heat transfer characteristics. It was certain at present that there were two main forms of heat transfer inside DPFB. One was the heat transfer between particles and fluidizing gas. There were many empirical correlations available for the estimation of fluid to particle heat transfer coefficient13, 14. Among them, the correlation of D.J Gunn was able to describe heat transfer of both fixed and fluidized beds of particles within the porosity range of 0.35-1.0. The second one was the heat transfer between bed bulk and surface (wall-to-bed) which was considered to affect the behavior of the fluidized bed more significant especially for small particles ( 8!

And

Γ where

7 1! A 7 A 1! 1 C A 1!F 9! ln A A @1 B @1 AB 7 A @1 B 2 7 7 7 2

7

A -H I

+ 10! +

1 , J ,

/K

11!

where + and + were conductivities of solid and fluid respectively, -H was an

experimental-determined particle shape factor, and -H 1.4 for particles with arbitrary shape. As a result, correlations of heat transfer inside bed phase of DPFB could be obtained based on energy conservation: L 7 M M 12! L ) 7N M M) ! 13! 2/+, ON L M M) ! 14! ln ⁄ ! )

L ) 7N M) M ! 15! L 7 M M 16! Among them, equation(12) and (16) were heat transfer between fluid and solid in reaction and annular regions respectively, where M and M were fluid temperatures in two regions, M

and M represented mean solid temperatures, 7 and 7 were surface areas of particles in two regions; (13) and (15) represented heat transfer of wall-to-bed in two regions, where M) and M) were inner and outer surface temperatures of the vessel wall and 7N and 7N were surface areas of inner and outer wall immersed in beds; while equation(14) was heat transfer inside the vessel wall, where ON was the length of vessel immersed in two beds. Combining above correlations with conservation equations of mass and energy shown as (17) to (19), the heat transfer model of DPFB was completed. %,PQ $,PQ %,:RS $,:RS 17! %,PQ $,PQ %,:RS $,:RS 18! L -. T M,PQ M,:RS -. T M,:RS M,PQ 19! where subscripts 1 and 2 represented parameters in reaction and annular regions respectively unless otherwise state, UV and W1 represented inlet and outlet. It was assumed that the total mass and the energy in solid particles would not vary during the operation, thus the conservation equations of solid phase was not presented here. In addition, it had to note that gas properties in 5



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this model were dependent on the temperature. 3.2 Model sensitivity In order to investigate heat transfer characteristics of DPFB, one dimensional finite difference method was employed to solve these equations. The control volume was divided into several elements with the same length along the axial direction, and a function of gas properties to temperature was established to renew the properties in each element to minimize the error. By given the boundary condition which included inlet temperatures and gas velocities, the calculation could be started at the inlet of DPFB. The effects of parameters including packet contact time, bed voidage particle diameter and inlet gas velocity on heat transfer performance were studied numerically. There were no mathematical relationships between the these parameters at present, therefore the sensitivity of heat transfer coefficient to each of these parameters could be obtained. 3.2.1 Contact time According to packet renewal model16, the rate of heat transfer was controlled by the rate of replacement of packets near the wall surface from the core area of bed which was described by contact time in the model. Contact times in reaction and annular region were set to equal tentatively (12 12), and by traversing 12 within [0.1, 10s] while fixing other parameters ( M,PQ 100℃; M,PQ 200℃; ℎ 0.51T; ℎ 0.33T; $,PQ 0.3 T⁄[ ; $,PQ 0.3 T⁄[ ; 0.44; "# 0.256TT) where ℎ and ℎ were static bed heights of reaction and annular region respectively, the variation of the overall heat transfer coefficient of DPFB \, with 12 could be obtained. To eliminate the effect of grid number on the modeling, the calculation results were compared by using three grid numbers (800, 1000 and 1200), as show in Figure 3. And \, was calculated using the follow expression: \, L/7) / 1, ! 20! where L was the total heat transferred which could be got by accumulating the heat transferred in each grid, 7) was the total heat transfer area of DPFB, 1, was the logarithmic mean temperature difference (LMTD). And 7) and 1, were defined as: 7) 7N 7 21! 1, ∆M:RS ∆MPQ !⁄ln ∆M:RS /∆MPQ ! 22! where 7 was the surface area in the freeboard of outer wall, ∆M:RS and ∆MPQ were

temperature differences of the outlet and inlet of DPFB respectively: ∆M:RS M,:RS M,:RS 23! ∆MPQ M,PQ M,PQ 24!

As could be seen in Figure 3, hardly any differences were observed among calculation results of these three grid numbers. Accordingly, without a significant sacrifice in accuracy, the medium grid (1000) was selected in the following study. Meanwhile, it was observed that \, decreased with contact time, which was consisted with the theory. As described in packet renewal model, it was particles movement along the vessel wall that improved the wall-to-bed heat transfer. If particle moved quickly, particle replacement rate would increase, leading to a larger temperature difference between the bed and the wall which improve the heat transfer, and it was the reason that contact time had a negative impact on overall heat transfer coefficient. In addition, there was always a turning point in the curve, it was the point that wall-to-bed heat transfer coefficient started to be lower than fluid to solid heat transfer coefficient along with the increase in contact 6


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time. According to the following expression which could be derived from equation(12)-(16), the overall heat transfer coefficient would be significantly influenced by the smaller factor, and such a phenomenon was much easier to be observed in numerical simulation.

^_

`abc

@

def dbc

B

@

def

`bgc dec

B

hf ijhf ⁄hc k_

`abc

@

def dbc

B

@

def

`bgc dec

B

(25)

Figure 3. Variation of overall heat transfer coefficient of DPFB with contact time. 3.2.2 Bed voidage The influence of bed voidage on heat transfer was also investigated. It had to be emphasized that the voidage in the core area of fluidized bed was different from that in the packets , which was equal to the value of fixed bed in this model. Firstly, by traversing within [0.01, 0.6] while fixing other parameters (ℎ 0.51T; ℎ 0.33T; $ 0.3 T⁄[ ; $ 0.3 T⁄[ ; 12 12 5[; "# 0.256TT), the variation of the overall heat transfer coefficient of DPFB \, with could be illustrated. According to Figure 4, hardly any influence of on \, could be

observed. It was mainly because that the overall heat transfer coefficient was influence by both fluid to solid and wall-to-bed heat transfer. However, the variation of only change the rate of fluid to solid heat transfer, while had no impact on wall-to-bed heat transfer model.

Figure 4. Variation of overall heat transfer coefficient of DPFB with voidage of fluidized bed under different inlet temperatures. In the same way, the effect of , was indicated in Figure 5. It was found that , had a negative impact on \, . This was due to that larger , led to smaller +, which mainly determined the wall-to-bed heat transfer. Meanwhile, it was speculated that , in this range caused smaller wall-to-bed heat transfer coefficient compared with fluid to solid heat transfer 7



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coefficient, as a result, the trend of \, was consistent with that of wall-to-bed heat transfer coefficient which was similar with the result of 3.2.1. Besides, it was found that the inlet temperatures of fluids did not have a significant impact on \, , for they only influenced the gas properties which had a little effect on heat transfer compared with solid.

Figure 5. Variation of overall heat transfer coefficient of DPFB with voidage of packet under different inlet temperatures. 3.2.3 Particle diameter The effect of particle diameter on \, was demonstrated in Figure 6. It seemed that \, did not vary with "# significantly, though the findings in literatures showed that wall-to-bed heat

transfer would improve by decreasing particle diameter. It was because that particle diameter was only a factor that influence heat convection between fluid and solid as indicated in equation (1) and (2), while it had no contribution to wall-to-bed heat transfer in this model (its influence on 12 was not taken into consideration in this section). As a result, the effect of particle diameter on the overall heat transfer coefficient was not significant. It also seemed that there was little effect of inlet gas velocity on heat transfer. This was mainly because that the correlations of inlet gas velocity and other parameters such as contact time of packet had not been added to the model as mentioned earlier. According to this result, it can be concluded that gas velocity itself was not a critical factor of heat transfer in DPFB without regard to its relationship with other parameters.

Figure 6. Variation of overall heat transfer coefficient of DPFB with particle diameter under different inlet gas velocities. 3.3 Experimental validation Among all parameters in this model, only contact time of packet was uncertain which had to be 8


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determined by experiment. It was found that 12 was the function of several parameters such as particle diameter and gas velocity, and we ignored the effect of temperature and cross section of the fluidized bed on 12 . Therefore 12 would equal to 12 if all other parameters in reaction and annular regions were the same. As a result, 12 (or 12 ) could be derived using equations (12) to (19) by given the outlet temperatures of DPFB which could be measured through experiment. The fluid and solid used in the experiment were air and silica sand whose properties were listed in Table 1. Table 1 also contained operation parameters unless otherwise state. Earlier testing results showed that static bed height had little effect on 12 , thus only the relationship of 12 and gas velocity was considered in this study. By measuring outlet temperatures using different inlet gas velocities, the correlation of 12 (or 12 ) and $,PQ (or $,PQ ) was derived through regression analysis:

12

1.1

1.28 26! *$,PQ and this equation was only suitable for $,PQ within 50.2 0.6T/[9. By combining equation (26),

the model of DPFB was completed which could be used to predict the heat transfer performance in steady state. To validate the model, some more experiments were carried out under relatively high temperature using the same bed material and fluidizing gas. Figure 7. showed the comparison of predicted and measured outlet temperatures of reaction region. The reasonable agreement within +20% deviation was observed. The overestimated of the model might be ascribed to the heat loss of the reactor which was particularly significant under high temperature condition. Therefore, to improve the model, the energy dissipation had to be considered in the future work. Table 1. Operation parameters of the experiment Solid bulk density, \l⁄T Solid conductivity, m ⁄T ∙ ℃ Solid heat capacity, o⁄\l ∙ ℃ Mean particle diameter, TT Bed voidage, Static be height in reaction region, T Static be height in annular region, T Gas velocity at inlet, T/s Parameter

Value

1450 7.6 840 0.256 0.44 0.055 0.083 0.2-0.6

9



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Figure 7. Comparison of the outlet temperatures of reaction region between experiment and calculation values. 4 HEAT TRANSFER SIMULATION By adding the correlation of inlet gas velocity and contact time (equation 26), the model was able to predict the heat transfer behavior of DPBF. The heat transfer characteristics of DPBF was simulated in this section. Effects of operation parameters including inlet gas velocities and static bed height on heat transfer performance were considered. 4.1 Temperature distribution in DPFB The temperature distribution inside both reaction and annular fluidized bed of DPFB was investigated firstly to examine the heat transfer performance. Since the calculation was achieved by one dimensional finite difference method, temperature distribution along the axial direction could be plotted. The inlet temperatures of reaction and annular regions were set to be 375 and 800℃ respectively which were the actual values from experiment. As shown in Figure 8 (a), with increase in axial position, M and M increased and decrease dramatically at first, then there were nearly no changes in them when axial position rose further. The rapid changing of M and M occurred at lower position in DPFB where particles occupied the major space. Increasing $,PQ and $,PQ from 0.3T/[ to 0.5T/[ while remaining other parameters unchanged, the temperature difference of M and M at outlet of DPFB increased as well, as shown in (b). When static bed heights were increased from ℎ 0.055T and ℎ 0.083T to ℎ 0.206T and ℎ 0.167T respectively, little temperature difference at outlet was observed as shown in (c). The above finding indicated that since solid particles would make a great contribution to heat transfer inside DPFB, adding bed materials to both reaction and annular regions of DPFB would improve heat transfer significantly.

10


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(a)

(b)

(c)

Figure 8. Temperatures of fluid in reaction (M ) and annular region (M ) along the axial direction for (a) $,PQ $,PQ 0.3T/[; ℎ 0.055T; ℎ 0.083T, (b) $,PQ $,PQ 0.5T/[; ℎ 0.055T; ℎ 0.083T and (c) $,PQ 0.5T/[; ℎ 0.206T; ℎ 0.167T.

4.2 Effect of operation parameters on heat transfer in DPFB According to literatures, parameters such as gas velocity and properties of fluid and solid had significant influence on wall-to-bed heat transfer. Therefore, the effect of operation parameters including inlet gas velocity and static bed height on heat transfer characteristics in DPFB was investigated. Firstly, by setting $,PQ $,PQ and varying them together, heat transfer coefficient as a function of the inlet gas velocity was derived numerically. As could be seen in Figure 9, all , , ) and ) increased with inlet gas velocity which was consistent with finding of 11



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other researchers. The increase of fluid to solid heat transfer coefficients( and ) was

attributed to the increase in Reynolds number in fluidized bed. Wall-to-bed heat transfer coefficients in both regions () and ) ) were more sensitive to the gas velocity. It was attributed to the impact of contact time of packet which decreased with gas velocity significantly. However, temperatures of fluid and solid in reaction region decreased with inlet gas velocity as shown in Figure 10. It was due to that the increase in gas velocity resulted in increase in mass flowrate, which would reduce the average temperature.

Figure 9. Effect of inlet gas velocity on heat transfer coefficients inside DPFB.

Figure 10. Effect of inlet gas velocity on temperatures of fluid and solid in DPFB. Then, the effect of static bed height on heat transfer was illustrated in Figure 11. Generally, the quantity of bed materials in reaction region was fixed which was determined by process demand. Therefore only the effect of bed height in annular region was considered. According to Figure 11 (a), temperatures of fluid and solid did not change much with static bed height for M,:RS and M,:RS were almost the same, and the heat could not exchange anymore. (b) showed the plot when M,PQ was reduced to 175℃. As could be seen, the temperature difference between M,:RS and M,:RS increased compared with (a) when static bed height was relatively low, and the

impact of static bed height on temperature became more remarkable especially for solid. The above findings indicated that the annular fluidized bed was playing an important role in heat transfer inside DPFB. The heat transfer inside it would be dramatically enhanced compared with the case where no particles were added. And only when the temperature difference between the inlets of DPFB was large enough, could the effect of static bed height on M,:RS and M,:RS be 12


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clearly observed. (a)

(b)

Figure 11. Effect of static bed height in annular region on temperatures of fluid and solid in DPFB for (a) M,PQ 375℃, M,PQ 800℃ and (b) M,PQ 175℃, M,PQ 800℃. 4. CONCLUSIONS In this study, a novel double pipe fluidized beds (DPFB) reactor was designed whose cylindrical reaction region was covered by an annular bed fluidized in order to replenish energy during fast pyrolysis process. A spiral heat exchanger was mounted to the reactor to further utilize the heat of the vent gas. The construction of the new reactor was elaborated. After that, the steady state heat transfer model of DPFB was developed theoretically based on classical equations which described heat transfer inside fluidized bed and that of wall-to-bed respectively, and one dimensional finite difference method was employed to solve the equations. The model sensitivity analysis and validation were carried out. The model was then used to predict the heat transfer characteristics of DPFB. The main conclusions were drawn as follows: (1) The model would not influenced by the grid number of the control volume, and sensitivity analysis showed that contact time and bed voidage had a negative impact on the overall heat transfer while the effect of particle diameter was not significant. (2) The comparison of measured and calculated outlet temperature of reaction region showed reasonable agreement within +20% deviation. The error of the model might be ascribed to the heat loss of the reactor under high temperature condition. (3) Heat exchange mainly occurred in the bed phase where solid particles were playing an important role. 13



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(4) Both reducing contact time and increasing Reynolds number would improve heat transfer significantly which could be achieved by increasing inlet gas velocity. But increasing inlet gas velocity would not raise the temperature in reaction region because the mass flowrate would increase as well. The effect of static bed height of annular region on outlet temperatures of fluid could only be observed when inlet temperature difference was large enough. AUTHOR INFORMATION Corresponding Author *Hui Si. E-mail: [email protected] Notes The authors declare no competing financial interest. ACKNOWLEDGMENTS We would like to express our sincere gratitude to the Fundamental Research Funds for the Central Universities (BLYJ201515) and the 948 project of State Forestry Administration of the People's Republic of China (2012-4-19) for providing financial support for this study. NOMENCLATURE Acronyms DPFB = double pipe fluidized beds LMTD = log mean temperature difference Symbols 7N = inner surface area of vessel wall immersed in bed (T ) 7N = outer surface area of vessel wall immersed in bed (T ) 7) outer surface area of vessel wall T ! 7 = surface area of particles in reaction region (T ) 7 = surface area of particles in annular region (T ) 7 = outer surface area of vessel wall in the freeboard (T )

-. = specific heat (o/ \l℃!) "# = particle diameter (TT)

\, = overall heat transfer coefficient of DPFB (m/ T ℃!) = Nusselt number (-) = Reynolds number (-) = Prandtl number (-) L = total thermal load of DPFB (W) M,PQ = fluid inlet temperature of reaction region (℃) M,PQ = fluid outlet temperature of annular region (℃) M,:RS = fluid outlet temperature of reaction region (℃) M,:RS = fluid inlet temperature of annular region (℃) 12 = contact time (s) $,PQ = inlet gas velocity of reaction region (℃) $,PQ = inlet gas velocity of annular region (℃) $,:RS = outlet gas velocity of reaction region (℃) $,:RS = outlet gas velocity of annular region (℃)

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Greek letters = heat transfer coefficient of between fluidizing gas and solid (m/ T℃!)

) = heat transfer coefficient of wall-to-bed (m/ T ℃!) = bed viodage (-) , = bed viodage of fixed bed (-) + = conductivity (m/ T℃!) & = viscosity (\l/ T s!) % = density (\l/T)

REFERENCES (1) Chen, J. C.; Grace, J. R.; Golriz, M. R., Heat transfer in fluidized beds: design methods. Powder Technol. 2005, 150, 123-132. (2) Sunderesan, S. R.; Clark, N. N., Local heat transfer coefficients on the circumference of a tube in a gas fluidized bed. Int. J. Multiphase Flow 1995, 21, 1003-1024. (3) AL-Busoul, M. A.; Abu-Ein, S. K., Local heat transfer coefficients around a horizontal heated tube immersed in a gas fluidized bed. Heat Mass Transfer 2003, 39, 355-358. (4) Masoumifard, N.; Mostoufi, N.; Hamidi, A.-A.; Sotudeh-Gharebagh, R., Investigation of heat transfer between a horizontal tube and gas–solid fluidized bed. Int. J. Heat Fluid Flow 2008, 29, 1504-1511. (5) Mandal, D.; Sathiyamoorthy, D.; Vinjamur, M., Heat Transfer Characteristics of Lithium Titanate Particles in Gas-Solid Packed Fluidized Beds. Fusion Sci. Technol. 2012, 62, 150-156. (6) Mandal, D.; Vinjamur, M.; Sathiyamoorthy, D., Hydrodynamics of beds of small particles in the voids of coarse particles. Powder Technol. 2013, 235, 256-262. (7) Cheng, P.; Hsu, C. T., Fully-developed, forced convective flow through an annular packed-sphere bed with wall effects. Int. J. Heat Mass Transfer 1986, 29, 1843-1853. (8) III, F. J. D.; Jackson, R.; Ginestra, J. C., The phenomenon of pinning in an annular moving bed reactor with crossflow of gas. Chem. Eng. Sci. 1986, 41, 1485-1495. (9) Raupp, G. B.; Nico, J. A.; Annangi, S.; Changrani, R.; Annapragada, R., Two-flux radiation-field model for an annular packed-bed. AIChE J. 1997, 43, 792-801. (10) Demirel, Y.; Kahraman, R., Thermodynamic analysis of convective heat transfer in an annular packed bed. Int. J. Heat Fluid Flow 2000, 21, 442-448. (11) Collin, A.; Wirth, K.-E.; Ströder, M., Experimental characterization of the flow pattern in an annular fluidized bed. Can. J. Chem. Eng. 2008, 86, 536-542. (12) Qiu, G.; Ye, J.; Wang, H., Investigation of gas–solids flow characteristics in a circulating fluidized bed with annular combustion chamber by pressure measurements and CPFD simulation. Chem. Eng. Sci. 2015, 134, 433-447. (13) Gunn, D. J., Transfer of heat or mass to particles in fixed and fluidised beds. Int. J. Heat Mass Transfer 1977, 21, 467-476. (14) Nsofor, E. C.; Adebiyi, G. A., Measurements of the gas-particle convective heat transfer coefficient in a packed bed for high-temperature energy storage. Exp. Therm. Fluid Sci. 2001, 24, 1-9. (15) Saxena, S.; Gabor, J., Mechanisms of heat transfer between a surface and a gas-fluidized bed for combustor application. Prog. Energy Combust. Sci. 1981, 7, 73-102. 15



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(16) Mickley, H. S.; Fairbanks, D. F., Mechanism of heat transfer to fluidized beds. AIChE J. 1955, 1, 374-384. (17) Zehner, P.; Schlünder, E. U., Wärmeleitfähigkeit von Schüttungen bei mäßigen Temperaturen. Chem. Ing. Tech. 1970, 42, 933-941.

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Novel Concept of Fluidized Bed Reactor Design for Highly Efficient

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