- Email: [email protected]
Steady-state simulation of internal heat-transfer characteristics in a double tube reactor
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Contents lists available at ScienceDirect
Chemical Engineering & Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Steady-state simulation of internal heat-transfer characteristics in a double tube reactor
T
Yong Baia, Hui Sia, , Xiao Wangb ⁎
a b
School of Technology, Beijing Forestry University, No. 35 Tsinghua East Road Haidian District, Beijing 100083, PR China Chinese Academy of Forestry, Houxiangshan Road, Summer Palace, Haidian District, Beijing 100091, PR China
ARTICLE INFO
ABSTRACT
Keywords: Double tube reactor Fixed bed Fluidized bed Heat transfer Numerical simulation
A novel annular reactor-double tube reactor (DTR) was utilized innovatively to provide thermal energy for thermochemical reactions. A steady-state heat transfer model of DTR was established where classical heat transfer equations between the fixed bed, the fluidized bed and the wall were considered. The independence analysis showed that the different number of meshes was irrelevant. Compared with the experimental values, it was found that the calculated values of the model deviated within +15%, which illustrated good accuracy for temperature prediction. Based on the model, the numerical simulation of DTR showed that the heat exchange mainly occurred in the area with packing through the wall on both zones. Particle diameter in the reaction zone had a greater effect on heat transfer than the particle void ratio, and the bed-wall heat transfer coefficients were more sensitive to the increase of inlet gas velocities. Heat transfer would achieve quickly in DTR with only small packing heights. When temperature differences between the inlet gases of the two zones reached 600℃, the effect of the initial packing height on the fluid and solid temperature in DTR could be significantly observed, indicating the importance of the annular fixed bed in strengthening the heat transfer process.
1. Introduction Annular reactors with the advantages of simple structure, excellent heat and mass transfer properties have been widely used in chemical and industrial production in the last four decades [1]. For most thermochemical reactions, a stable reaction environment and a continuous heat source are required. Currently, there are two main methods to implement this process in annular reactors. One method is to use the annular zone of the annular reactors as the reaction zone and the inner zone as the heating zone [2]. Shedid et al. developed an annular reactor to heat an internal fluidized bed using an electric heater [3]. Berruti et al. utilized a built-in cylindrical fluidized bed as the combustion chamber and the reaction zone as an annular zone [4]. However, the biggest problem of this method in practice was that the control of the internal heating zone was difficult, resulting in a large temperature gradient in the temperature of the annular zone. The phenomenon that the uniformity of the distribution of the feedstock in the annular zone was serious[5]. Another method was to use the annular zone as the heating zone and the inner zone as the reaction zone. A number of studies have found that the best form of heat transfer in annular reactors was to transfer heat from the outside to the inside. Qiu et al. developed a new type of annular combustion chamber to improve the ⁎
air permeability and provide enough space to arrange the heat transfer surface [6]. Wang et al. designed a double-fluidized bed reactor filled with quartz sand between two concentric cylinders and studied its heat transfer characteristic [7]. Husan et al. developed a circular doubletube moving bed reactor for studying the chemical cycle gasification reaction of iron-based oxygen carriers to carbon [8]. Besides, Husan et al. had used rice husk as solid fuel for chemical cycle combustion in a circular double-tube moving bed reactor, and verified the internal temperature distribution with experimental data [9]. Alarifi et al. designed a new type of circular multi-tube reactor for the production of methanol and carried out its steady-state simulation numerically [10]. In recent decades, many studies have focused on the development of annular reactors, instead of the form of heating. According to the literature [11], one method was to incorporate a combustion chamber into the reaction zone. However, it severely affected the hydrodynamic properties in the reaction zone and increased the inconvenience of packing. Another method was to equip the periphery of the reaction zone with an electric heater, which had been widely used for experimental research. However, electric heating consumed more energy than combustion heating. Heat dissipation in electric heating was severe and it was almost impossible to utilize excessive heat [12]. To solve this problem, a double tube reactor (DTR), essentially an
Corresponding author. E-mail address: [email protected] (H. Si).
https://doi.org/10.1016/j.cep.2019.107642 Received 9 May 2019; Received in revised form 20 August 2019; Accepted 21 August 2019 Available online 30 August 2019 0255-2701/ © 2019 Published by Elsevier B.V.
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Nomenclature
tm v
Acronyms DTR
Creek Letters
Double tube reactor
αfs αws
Symbols Ags1 Ags2 Asw1 Asw2 As CF c d h lw K Nu Pr Q Re r1 r2 T tc
Logarithmic mean temperature, ℃ Velocity, m/s
Outer surface area of vessel wall immersed in bed, m2 Outer surface area of vessel wall, m2 Surface area of particles in reaction zone, m2 Surface area of particles in annular zone, m2 Effective heat transfer area in the bed phase, m2 Particle shape factor, − Specific heat capacity, J/(kg·℃) Diameter, mm Packing height, m Inner tube length, m Overall heat transfer coefficient in DTR, W/(m2·℃) Nusselt number, − Prandtl number, − Energy synthesis of all grid cells, W Reynolds number, − Radius of the reaction zone, mm Radius of the annular zone, mm Temperature, ℃ Contact time, s
φg φg,0 λ δ ρ μg μw1
Heat transfer coefficient between gas and solid, W/(m2·℃) Heat transfer coefficient between the vessel and solid, W/ (m2·℃) Particle void ratio, − Packet void ratio, − Conductivity, W·(m·℃) Wall thickness, m Density, kg/m3 The dynamic viscosity of the fluid, Pa·s The dynamic viscosity of the fluid at wall's temperature, Pa·s
Subscripts f g in out s w 1 2
annular reactor, was designed to replace electrical energy by the heat of the combustion gases. In this reactor, the reaction zone was a fluidized bed with excellent heat transfer and reaction characteristics; the annular zone was a fixed bed with almost no packing-back phenomenon, where the fluid and packing were in full contact [13]. By using a combustion chamber instead of an electric heater, a high temperature environment in the annular fixed bed could be achieved, thereby further significantly improving internal heat transfer. Therefore, the annular fixed bed played an important role in heat transfer because it determined the heat transfer coefficient largely. However, few studies have reported the principles and characteristics of such reactors above, especially the heat transfer behavior. Based on the boundary layer theory, the heat transfer research method was only applied to a single fluidized bed or a fixed bed reactor to obtain internal heat transfer coefficients, temperature distribution and fluid characteristics under the constant wall temperature condition [7,10]. There was little research on the heat transfer between the annular and the reaction zones in an annular reactor [3,5,7,11,12]. In other studies, experiments and numerical simulations were only used to predict temperature changes at specific locations in reactors at different conditions to verify the rationality of the reactor structure, with little consideration for temperature distribution, heat transfer coefficients between fluids and overall heat transfer coefficients [1–3,5,9].Moreover, to the best of author's knowledge, there was little literature on the fixed bed as the heating zone for an annular reactor. Therefore, in order to obtain good operating parameters of this reactor, especially for heat transfer characteristics, some fundamental investigations must be conducted to demonstrate the efficiency of heat transfer between the fixed bed and the fluidized bed. The purpose of this paper was to study the heat transfer characteristics of DTR. First, the principle and structural design of DTR were described. It used the annular zone (a fixed bed) and the cylindrical zone (a fluidized bed) as the heating zone and the reaction zone respectively. It was our innovative design for the structure of annular reactors. Then, a first attempt on modeling the internal heat transfer of DTR was presented and the model was verified through experimental
Fluid Gas Gas inlet parameters gas outlet parameters Solid The vessel wall Parameters in reaction zone Parameters in annular zone
data. Finally, based on the model coupled with classical heat transfer and energy equations, the numerical simulation of DTR was carried out by the commercial CFD software-Fluent V6.3.26, and the effects of various operating parameters on the internal heat transfer of DTR were studied. 2. Experimental setup 2.1. The double tube reactor The structure of DTR was shown in Fig.1.The inner tube acted as the reaction zone for thermochemical reactions, and the pre-filled quartz sand particles were in a fluidized state by air through an air distribution plate at the bottom of DTR while the annular zone filled with many Pall rings was used as the heating fixed bed. They could increase the residence time of high-temperature combustion gas and continue to heat the inner tube. A combustion furnace was tangentially mounted on the annular zone to generate combustion gas. With a spiral plate heat exchanger located below the DTR, the excess heat energy of the combustion gas out of the annular zone could preheat the fluidizing gas entering the reaction zone. The circulation of combustion gas in the two zones was achieved by a return tube. The design and packing specifications of the DTR have been summarized in Table 1. 2.2. Apparatus The DTR above and other assisted devices were used to determine the operating conditions required to achieve the objectives of the study. The experimental apparatus was designed and constructed as shown in Fig.2. The liquefied petroleum gas tank supplied petroleum gas to the furnace and the gas pressure was controlled through a pressure regulator. 50kPa- compressed air exited from the compressed gas tank into the two branches with independent gas flow was adjusted by ball valves, and the flowmeters were used to measure amount of the flow rate. One branch provided compressed air as a fluidized carrier gas for the reaction zone. The air flowed through the distribution plate to 2
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Table 1 Design and Packing Specifications of the DTR. Parameter
Value
inner tube diameter outer tube diameter inner tube length outer tube length tubes material others material packing type
180 mm o.d., 170 mm i.d. 273 mm o.d., 263 mm i.d. 1.4m 1.55m 310S 304 inner tube: quartz sand; outer tube: Pall rings 0.256 mm 25 × 0.3 × 0.3 mm SiO2 310S
quartz sand particles specification Pall rings specification quartz sand material pall rings material
of DTR, the both zones needed to be divided differently. Along the axial direction, the reaction zone was essentially a type of bubbling bed, which could be divided into a dense bed phase and a freeboard. The annular zone belonged to a fixed bed, which was divided into a packing area and an unpacking area, as shown in Fig.3. Since DTR was a revolving body, it was examined to be a simplified model in which the zone on one side of the axis was used as the control body. There were many particles in the dense bed phase of the reaction zone with many bubbles rising during the fluidization process. Its height was related to the parameters of fluidized gas velocity and particles material characteristics. The heat transfer efficiency of the fluidized gas, the inner wall surface and the movement of the particles was closely related. In the freeboard, there were fewer particles where convective heat transfer mainly between the gas and the wall occurred, and temperature gradient was generated in the radial direction. The packet renewal theory was used to divide the fluidized bed in the reaction zone into the core area and the thermal boundary layer [7]. Different from the core area, the thermal boundary layer had a small particle layer where particles in an irregular motion state collided with the wall in the form of "packets" or “clusters” and heat exchange occurred rapidly, resulting in an intense temperature gradient [14]. The packing in the annular zone was fixed and the heat transfer process in the unpacking area was the same as the freeboard in the reaction zone. 3.2. Heat transfer inside DTR 3.2.1. Heat transfer in the reaction zone The heat transfer of fluidized bed in the reaction zone included (1) heat exchange between gas and packing in bed phase and (2) heat transfer between packing and vessel wall in bed phase and (3) convective heat exchange of gas in the freeboard of the reaction zone and the unpacking area of the annular zone. For heat exchange between gas and packing in bed phase, gas to solid heat transfer coefficient αgs could be solved by Gunn's empirical equation (adapted to fluidized bed and fixed bed) [15]:
Fig. 1. Schematic diagram of DTR: (1) return tube; (2) annular distributor; (3) quartz sand particles; (4) spiral heat exchanger; (5) circular distributor; (6) feedstock inlet;(7)Pall rings; (8) flange; (9) reaction zone; (10) annular zone.
produce a uniform air flow into the DTR, which ultimately exited as heated air. The other one provided compressed air for the furnace. NiCr/Ni-Al thermocouple probes (type K) were used to measure the temperature of the fluid throughout DTR. The temperature signals were collected by an A/D converter. The measurements were monitored and recorded by SIEMENS S7-200 PLC software to calculate the internal heat transfer coefficient.
Nu = (7
10 + 5 2)(1
0.7Re 0.2Pr 1/3) + (1.33
2.4 + 1.2 2) Re 0.7Pr 1/3 (1)
3. Model development
where Re was the Reynolds number of gas, Pr was Prandtl number of gas in bed phase, φ was the bed void ratio of fluidized bed in dense phase. The effective heat transfer area As of bed phase could be calculated by the specific surface area of the packed bed or the total area of the particles per unit volume, assuming particles were spherical [16]:
3.1. Heat transfer model of DTR
As = V
For this reactor structure, it was necessary to analyze the heat transfer in different parts. Due to the different properties of the packing
For heat transfer between packing and vessel wall in bed phase, its heat transfer coefficient αsw could be calculated by conventional packet 3
6(1 ds
) (2)
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Fig. 2. Schematic diagram of the experimental apparatus.
Fig. 3. Physical model of heat transfer inside the DTR.
renewal model [14]: sw
=2
capacity, φg,0 was the bed void ratio stationarily. The effective thermal conductivity was the most difficult parameter to determine in the fluidized bed-to-wall heat transfer model. There were various derivation formulas in the literature works. Among them, the formula proposed by Zehner was mostly applied [18]. The formula also considered the effect of the relative thermal conductivity of the solid phase and gas phase:
(3)
pc /( tc )
where λ was the effective thermal conductivity of bed phase, ρ was the density of bed phase, c was effective specific heat capacity, tc was the particle contact time, characterizing the average period of contact of a single particle with vessel wall. And λ, c, ρ represented the physical properties of the whole bed phase, instead of the physical properties of solid particles, and their relationship was as follows[17]:
=
s
+
g
1-
g,0
where
(4)
s
=
where ρs was the solid particle density, cs was the solid particle specific heat capacity, ρg was the gas density, cg was the gas specific heat
g
= (1
c = (1-
g,0 ) s c s
+
g,0 s cs
(5)
4
[ A + (1 1-
g,0 ) g,0
) ]
g
(6) (7)
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
(8)
= 7.26 × 10-3
3.2.3. Energy equations According to the above assumption, the fluidized bed dense phase and the fixed bed packing area had the following relationships (subscripts 1, 2 indicated the reaction zone and the annular zone, the same as below):
and
=
2
(A B ) A
(1
1) B A ln A B
B 2 ) A
(1
(B
1)
(1
B ) A
B+1 2
(9)
where
A=
s
(10)
g
B = CF (
1-
g,0 10/9 )
where λs and λg were the thermal conductivities of solid and gas respectively, and CF was the particle shape factor. For convective heat exchange of gas in the freeboard of the reaction zone and the unpacking area of the annular zone, its heat transfer coefficient α1 could be calculated by the equation of forced convection heat transfer in the tube [11]: 1
=
f
d1
Nu
f
2 1
dt
1/3
d1 l
0.14
µ1
(13)
µ w1
(20)
Q2 =
gs 2 As 2 (Ts2
Tg 2)
(21)
Q2 =
sw 2 Ab2 (Tg 2
Ts2)
(22)
2 w lw (Tw1 ln(r2/ r1)
Tw2)
(b ) y
+
4d p L
+(
(15)
(16) 2)
1
B
0.26 0.216
= 2.58
dp G µ
1/3
cp µ
1/3
dp G µ
0.4
cp µ
(25)
g 2, in vg 2, in
=
g 2, out vg 2, out
(26)
K=
(17)
+ 0.094
g1, out vg1, out
Tg1, out )
(27)
Ts2, in)
(28)
According to the packet renewal theory [21], the heat transfer efficiency between the fluidized bed and the wall surface in the reaction zone was mainly affected by the replacement rate of packets near the wall, which was described by contact time. The parameters were assumed as follows: Tg1, in = 100℃, Tg2, in = 200℃；h1 = 0.08 m, h2 = 0.12 m；vg1,in= vg2,in = 0.5 m/s；φg = 0.44；ds = 0.256 mm [7]. By traversing tc within [0.1, 10 s], the relationship between the overall heat transfer coefficient and the contact time of the packet in the reaction zone was obtained. Divide the DTR model with quadrilateral meshes by Fluent V6.3.26, for it consisted essentially of two quadrilateral faces. The quadrilateral mesh had fast production speed, good quality, simple data structure, smooth area, and the simulation result was closer to the true value [22–25]. Different grid numbers were selected and the grid independence was verified. At the same time, the calculation results were compared, as shown in Fig.4. The overall heat transfer coefficient was calculated using the following Eq. (29):
s
s
=
3.3. Model grid independence test
hf d w2 2dt dp
g1, in vg1, in
where in and out represented inlet and outlet.
(14)
dt2 Pr Rep
(23) (24)
Q = cg 2 mg 2 (Tg 2, out
Q A1 tm
(29)
where Q was the energy synthesis of all grid cells. In Fig.4, the prediction results of the five grid numbers (200, 400, 600, 800 and 1000) showed almost no difference. Considering the calculation accuracy and cost, the subsequent calculations latter used a medium grid number (600). It can be seen that K decreased as the contact time increased, which was in agreement with the theory [21]. The faster the particle packet moved, the higher the update frequency and the greater the temperature gradient between the fluidized bed and
where hw was the heat flow coefficient for wall, it could be calculated by the following Eq. (28) [20]
hf d p
Tw1)
Q = cg1 mg1 (Tg1, in
where y and b both were the dimensionless numbers, and Φ represented the influence of the fluid film on the heat transfer at the contact point of the packing. Both a1 and Φ(b) were functions of b, and Φ1 and Φ2 could be found in the manual[19]:
2
sw1 Ab1 (Ts1
For the second and third cases, the Eq. (1)could be used as an alternative. Meanwhile, according to the conservation of mass and energy, the following equations could be obtained [7]:
3.2.2. Heat transfer in the annular zone Heat transfer in the packing area of the fixed bed in the annular zone mainly included (1) heat transfer between fluid and packing and (2) contact heat conduction - packing contacted with vessel wall surface and (3) internal heat transfer between the packing. For heat transfer between fluid and packing, the convective heat transfer coefficient αgs could also be obtained by Eq. (1) For contact heat conduction - the packing contacted with the vessel wall surface, heat transfer coefficient hp could be calculated by Eq. (14) [15]:
=
Q1 =
Q = Q1 = Q 2
where μg was the dynamic viscosity of the fluid, μw1 was the dynamic viscosity of the fluid at wall's temperature, d1 was the tube inner diameter and l was the tube length.
b=
(19)
and
(12)
Nu = 1.86 Re Pr
y=
Ts1)
Q=
and
hp =
gs1 As1 (Tg1
The above four equations could be solved in series only when there was packing in both zones. According to the presence or absence of packing, it could be divided into three types: both of the two zones had packing, only one side had packing and no packing each side. For the first case, it could be solved by adding the wall internal heat conduction correlation [9]:
(11)
g,0
Q1 =
0.4
(18)
For internal heat transfer between the packing, the heat transfer coefficient was the thermal conductivity of the material. 5
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
distribution and overall heat transfer coefficient of different parts at specific working conditions. 4.1. Axial temperature distribution in DTR The simulated values of the axial gas temperature distribution inside the reaction zone and annular zone of DTR at different inlet velocities and packing heights was shown in Fig.6. It could be seen that at the assumed condition (Tg1, in = 300℃, Tg2, in = 800℃, ds = 0.256 mm, φg = 0.44), as the axial position increased, T f1 and T f2 respectively increased and decreased rapidly, and reached a steady state finally. The rapid change in fluid temperature was mainly ascribed to the heat transfer between the gas, wall and packing, which improved the heat transfer efficiency. Comparing Fig.6. (a) and (b), vg1,in and vg2,in increased from 0.3m⋅s−1 to 0.5m⋅s-1, it could be seen that the temperature difference between the two zones at steady state increased from 50℃ to 100℃, which attributed to the larger average temperature difference between the heat transfer and the smaller convective heat transfer coefficient between the fluid and the wall when vg1,in and vg2,in increased, indicating that heat exchange existed, but the efficiency was significantly reduced. Similarly, comparing Fig.6.(b) and (c), h1, h2 respectively were increased to 0.08 m, 0.12 m to 0.16 m, 0.2 m, it could be found that the temperature difference between the two zones tended to zero, indicating that the two zones has sufficient heat transfer and reached thermal equilibrium. When h1, h2 increased, the contact frequency between the wall and packets increased in the reaction zone, resulting in a significant increase in the average convective heat transfer coefficient of the solid phase [21].
Fig. 4. Relationship between overall heat transfer coefficient and contact time of particle packet in reaction zone under different grid numbers.
the wall surface in the reaction zone, which intensified the heat transfer efficiency. The overall heat transfer coefficient curve had two turning points (A and B) and the curvature decreased. This phenomenon was caused by the ratio change of fluidized bed-to-wall heat transfer coefficient, the fixed bed heat transfer coefficient and the gas-solid heat transfer coefficient. Combining Eqs. (19)–(24), the formula for calculating the overall heat transfer coefficient was as follows:
K=
1 d2 gs1 d1
+
d2 gs2 d1
+
d2 sw1d1
+
1 sw2
+
d2 d1
(30)
where αgs1、αgs2、αsw1 and αsw2 were variables, so K was sensitive to small values. Since αsw1 and αsw2 were affected by tc, as tc increased, the decreasing trend of K would be more obvious. In Fig.4.the A and B points happened respectively when the bed-to-wall heat transfer coefficient was equal to the gas-to-solid heat transfer coefficient and the fixed bed heat transfer coefficient.
4.2. Effects of operating parameters on internal heat transfer in DTR Operating parameters, including inlet gas velocity, packing specification, packing height and others, also affected the heat transfer of DTR. Therefore, these parameters on the heat transfer characteristics of DTR were investigated next.
3.4. Experimental validation
4.2.1. Bed void ratio in the reaction zone In the packet renewal theory, the bed void ratio included the particle void ratio and packet void ratio in a fluidized bed. Therefore, the effect of bed void ratio in the reaction zone on K was investigated firstly. According to the division of the model grid, the particle void ratio was different from the packet void ratio which was only equal to the particle void ratio stationarily [26]. By traversing φg and φg,0 within [0.4, 0.8] while fixing other parameters(h1 = 0.08 m, h2 = 0.12 m；vg1,in= vg2,in = 0.5 m/s；tc1= tc2 = 5 s; ds = 0.256 mm), the relationship of K with φg and φg,0 of the reaction zone at different inlet temperatures was illustrated in Fig.7. and Fig.8. respectively. It could be seen that φg had little effect on K, whose change interval was relatively small, while K decreased as the increase of φg,0. K was mainly affected by the fluidized bed-wall heat transfer coefficient and gas-solid heat transfer coefficient
The heating test of DTR was carried out by using the apparatus in section 2.2 When the heating process reached a steady state, the temperature of the gas outlet of the reaction zone was recorded. And the gas and solid in the reaction zone were air and quartz sand whose properties were listed in Table 2. with operation parameters, considering our previous research results [7] and experimental data in some literature [3,8,9]. In order to verify the reliability of the model, the calculated reactor outlet temperature values were compared with some more experimental values in Fig.5. The deviation of the predicted value from the measured value was within 15%, indicating that the model had high reliability. The overestimation of the calculated values may be attributed to the thermal energy dissipation, dynamic energy dissipation during the experiment, which began to become apparent at over 200℃. In addition, the model needed to be further improved, such as considering the effect of friction and uneven distribution of packing or fluid. Considering the predicted values of the outlet temperature, the error of DTR was lower than that of some devices, such as a reactor with annular double fluidized beds in [7], indicating that the heat transfer mode of DTR had certain advantages.
Table 2 Operation Parameters of the Simulation and Experiment.
4. Internal heat-transfer simulation The thermal mathematical model of DTR was completed by constructing the physical model and establishing the internal heat transfer relationship of DTR by means of the classical heat transfer and energy equations. It has been verified that the established model has high reliability and could predict heat transfer behavior better. This section used numerical simulation to predict parameters such as temperature 6
Parameter
Value
gas density, kg/m3 solid density, kg/m3 gas heat capacity, J/(kg·℃) solid heat capacity, J/(kg·℃) gas conductivity, W·(m·℃) solid conductivity, W·(m·℃) mean particle diameter, mm bed void ratio stationarily,— static height in reaction zone, m height in annular zone, m inlet gas velocity, m/s
1.225 2500 1008 840 0.025 7.6 0.256 0.44 0.08 0.12 0.1—0.8
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Fig. 5. Comparison of the outlet temperatures of reaction zone between calculated and experimental values.
simultaneously, but the φg only affected the gas-solid heat transfer coefficient, so K changed very little. The increase of φg,0 led to a decrease in the effective thermal conductivity in the bed phase, and K was positively correlated with it. In addition, φg,0 made the fluidized bedwall heat transfer coefficient smaller than the gas-solid heat transfer coefficient in the interval [0.4, 0.8]. Therefore, K was more sensitive to the fluidized bed-wall heat transfer coefficient. 4.2.2. Particle diameter in the reaction zone The effect of the particle diameter in the reaction zone on K was demonstrated in Fig.9(keeping h1 = 0.08 m, h2 = 0.12 m, Tg1, in = 300℃, Tg2, in = 800℃, φg = 0.44). It could be seen that the simulation and experimental results had almost the same trend, and K decreased significantly as dsincreased within [0.2, 1 mm]. At the same inlet gas velocity, the turbulence of small-diameter particle inside DTR was larger, and the collision frequency of particles and wall was higher, resulting in the stronger heat transfer. In addition, as the inlet gas velocity increased, K increased at the same particle diameter, but the difference was not significant. This was because small-diameter particles (the minimum fluidization velocity was lower) were more easily fluidized than large-diameter particles. Small-diameter particles encountered little resistance and got more kinetic energy from heated air, taking more heat away from the wall surface after colliding with the wall. In addition, the experimental results were lower than the simulation results, probably due to the thermal energy dissipation, dynamic energy dissipation. 4.2.3. Inlet gas velocity The relationship of heat transfer coefficient and different gas inlet velocities (vg1,in= vg2,in) was showed in Fig.10 (keeping h1 = 0.08 m, h2 = 0.12 m；Tg1, Tg2, φg = 0.44, in = 300℃, in = 800℃, ds = 0.256 mm). It could be seen that αfs1, αfs2, sw1 and αsw2 all increased with the increase of inlet gas velocities. The increase in fluid -solid heat transfer coefficients (αfs1 and αfs2) was attributed to the increase of vg1,in and vg2,in, resulting in the increase of the gas Reynolds number and more severe turbulence. Relatively speaking, the solid-wall heat transfer coefficients (αsw1 and αsw2) were more sensitive to vg1,in and vg2,in, increasing from 500 to 1700 W/(m2·℃) when vg1,in and vg2,in, increased from 0 to 1 m/s. As the vg1,in and vg2,in, increased, the average temperature difference between wall and fluid heat transfer decreased, and the convective heat transfer coefficients between the wall and the fluid increased, resulting in larger turbulent of the particles in the fluidized bed, which enhanced heat transfer between the particles.
Fig. 6. Axial distribution of gas temperatures in the annular zone (T f1)and reaction zone (T f 2) (a) h1 = 0.08 m, h2 = 0.12 m, vg1,in= vg2,in = 0.3m⋅s−1, (b) h1 = 0.08 m,h2 = 0.12 m, vg1,in= vg2,in = 0.5m⋅s−1, (c) h1 = 0.16 m,h2 = 0.2 m, vg1,in= vg2,in = 0.5m⋅s−1.
usually fixed to study the variation of the heat transfer effect with the height (h2) in the annular zone. The results were showed in Fig.11. As h2 increased, the temperature of the gas and solids in the reaction zone increased while that in the annular zone decreased. The reason may be that heat was transferred from the annular zone to the reaction zone, and the packing improved the heat transfer efficiency. Comparing Fig.11. (a) and (b), when the temperature difference between the two zones was increased from 400 to 600℃, the temperature at a steady state of Tf1, out and Tf2, out was reduced from about 500 to 400℃, while the temperature difference of Ts1 and Ts2 became larger. As shown in 11(b), it can be seen that the temperature gradient of the packing increased significantly, while the gas temperature did not change significantly. The above phenomena proved that the increase of the packing height in the annular zone could promote the heat exchange
4.2.4. Packing height in the annular zone Similarly, the heat transfer coefficient was calculated at different packing heights. In practice, the packing height in the two zones was often determined by the process requirements. Therefore, h1 was 7
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Fig. 10. Effect of inlet gas velocities on heat transfer coefficient of fluid and solid.
Fig. 7. Relationship between the overall heat transfer coefficient and the particle void ratio in the reaction zone at different inlet temperatures (Tf1,in: inlet fluid temperature in reaction zone; Tf2,in: inlet fluid temperature in annular zone).
Fig. 8. Relationship between the overall heat transfer coefficient and packet void ratio in the reaction zone at different inlet temperatures.
Fig. 11. Effect of the packing height in the annular zone on the internal temperatures of fluid and solid (a) Tg1, in = 400℃, Tg2, in = 800℃, h1 = 0.08 m, (b) Tg1, in = 200℃, Tg2, in = 800℃, h1 = 0.08 m.
Fig. 9. Effect of particle diameter in reaction zone on the overall heat transfer coefficient at different inlet gas velocities compared by experiment results.
efficiency between the two zones, which could quickly reach the thermal equilibrium. Besides, the effect of the packing height in the annular zone on the overall heat transfer coefficient was obtained, as shown in Fig.12. It can be seen that K increased as h2 increased and K increased with v1,in and v2,in.The reason may be that the annular zone was a fixed bed, and the gas can only pass through their gap during the period. As the height of the packing increased, the gas-solid heat exchange capacity increased and the “insulation” function of the entire annular zone became stronger. The increase of vg1,in and vg2,in caused the relatively lowtemperature gas after heat exchange to leave the annular zone quickly and the high-temperature gas was quickly replenished. The heat loss
Fig. 12. Effect of packing height in annular zone on overall heat transfer coefficient at different inlet gas velocities. 8
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
caused by the shortage of high-temperature gas and the slow increase of K were reduced, so the heat exchange rate between the two zones was increased slowly.
(2017) 519–529. [2] T.J. Brooms, M.S. Onyango, A. Ochieng, Photodegradation of phenol using TiO2, ZnO and TiO2/ZnO catalysts in an annular reactor, J Water Chem Technol. 39 (2017) 155–160. [3] M.H. Shedid, M.A.M. Hassan, Heat transfer characteristics of the fluidized bed through the annulus, Heat Mass Transfer. 52 (2016) 1943–1952. [4] F.M. Berruti, C.L. Briens, Novel intermittent solid slug feeder for fast pyrolysis reactors: fundamental and modeling, Powder Technol. 247 (2017) 95–105. [5] J.S.M. Botterill, M. Desai, Limiting factors in gas-fluidized bed heat transfer, Powder Technol 6 (4) (1972) 231–238. [6] G. Qiu, J. Ye, H. Wang, Investigation of gas−solids flow characteristics in a circulating fluidized bed with annular combustion chamber by pressure measurements and CPFD simulation, Chem. Eng. Sci. 134 (2015) 433–447. [7] X. Wang, H. Si, Novel concept of fluidized bed reactor design for highly efficient internal heat transfer, Ind. Eng. Chem. Res. 55 (2016) 9276–9283. [8] C.W. Hsuan, K. Young, Chemical Looping Gasification of Charcoal with Iron-Based Oxygen Carriers in an Annular Dual-Tube Moving Bed Reactor, Aerosol air quality research 16 (2015) 1093–1103. [9] C.W. Hsuan, K. Young, T. Hsin-Hsien, K. Yu-Lin, T. Yao-Hsuan, Rice husk as solid fuel for chemical looping combustion in an annular dual-tube moving bed reactor, Int. J. Chem Eng. 280 (2015) 82–89. [10] A. Alarifi, A. Elkamel, E. Croist, Steady-state simulation of a novel annular multitubular reactor for enhanced methanol production, Ind. Eng. Chem. Res. 52 (2013) 15387–15393. [11] H. Taofeeq, M. Al-Dahhan, Heat transfer and hydrodynamics in a gas-solid fluidized bed with vertical immersed internals, Int. J. Heat Mass Transfer 122 (2018) 229–251. [12] A. Singhal, S. Cloete, S. Radl, F.R. Quinta, S. Amini, Heat transfer to a gas from densely packed beds of cylindrical particles, Chem. Eng. Sci. 172 (2017) 1–12. [13] H.A. Aghdas, S.B. Ralph, Y.A. Habibollah, Effect of process conditions on product yield and composition of fast pyrolysis of Eucalyptus grandis in fluidized bed reactor, Ind. J. Eng. Chem. 7 (2014) 2594–2602. [14] R. Zarghami, N. Mostoufi, R. Sotudeh-Gharebagh, J. Chaouki, Analysis and modeling of Particle-wall contact time in gas fluidized beds, Chem Eng Sci. 62 (2007) 4573–4578. [15] D.J. Gunn, Transfer of heat or mass to particles in fixed and fluidised beds, Int. J. Heat Mass Transfer. 21 (1978) 467–476. [16] Cormos, Dynamic simulation of fluidized bed chemical looping combustion process with iron based oxygen carrier, Fuel 214 (2018) 436–445. [17] H.M. Abdelmotalib, M.A. Youssef, Im, Heat transfer process in gas–solid fluidized bed combustors: a review, Int J Heat Mass Transfer. 89 (2015) 567–575. [18] P. Zehner, E.U. Schlünder, Wärmeleitfähigkeit von Schüttungen bei mäßigen Temperaturen, Chem. Ing. Tech 42 (1970) 933–941. [19] Z.W. Zi, Petrochemical Design Manual, Chemical Industry Press, Beijing, 2015 (in Chinese). [20] A. Kumar, H.M. Jena, Removal of methylene blue and phenol onto prepared activated carbon from Fox nutshell by chemical activation in batch and fixed-bed column, J. Clean Prod. 137 (2016) 1246–1259. [21] S.E. Olsson, A.E. Almstedt, Local instantaneous and time-averaged heat transfer in a pressurized fluidized bed with horizontal tubes: influence of pressure, fluidization velocity and tube-bank geometry, Chem. Eng. Sci. 50 (1995) 3231–3245. [22] M. Sosnowski, J. Krzywanski, K. Grabowska, R. Gnatowska, Polyhedral meshing in numerical analysis of conjugate heat transfer, EPJ Web of Conferences vol. 180, (2018) 02096. [23] M. Sosnowski, J. Krzywanski, R. Gnatowska, Polyhedral meshing as an innovative approach to computational domain discretization of a cyclone in a fluidized bed CLC unit, E3S Web of Conferences vol. 14, (2017) 01027. [24] M. Spiegel, T. Redel, Y.J. Zhang, T. Struffert, J. Hornegger, R.G. Grossman, C. Karmonik, Tetrahedral vs. polyhedral mesh size evaluation on flow velocity and wall shear stress for cerebral hemodynamic simulation, Comput Method Biomec 14 (2011) 9–22. [25] J. Su, Z. Gu, C. Chen, X.Y. Xu, A two-layer mesh method for discrete element simulation of gas‐particle systems with arbitrarily polyhedral mesh, Int J Numer Meth Eng 103 (2015) 759–780. [26] H.S. Mickley, D.F. Fairbanks, Mechanism of heat transfer to fluidized beds, AIChE J 1 (1955) 374–384.
5. Conclusions A novel double tube reactor (DTR) was designed in which the internal reaction zone was wrapped by an annular fixed bed to supply energy for thermochemical reactions. A spiral heat exchanger was installed under the lower portion of DTR to further utilize the heat of the hot exhaust gas. The construction and principle of the reactor were elaborated. Then the heat transfer model of DTR at steady-state conditions was mathematically established. Based on the classical heat transfer and energy equations, the heat transfer process of fluidized bed and fixed bed was described. Model independence analysis and verification were also performed. Finally this model was used to predict the heat transfer characteristics of DTR. The main conclusions were as follows: (1) The independence analysis using five grid numbers (200, 400, 600, 800 and 1000) showed that the model was not affected by the number of control volume grids. Considering the accuracy and time of the calculation, a medium quantity grid (600) was finally selected. (2) The difference of the experimental and calculated values of the outlet temperature in the reaction zone was within ± 15%, which had high rationality. The error may be the heat loss of the reactor at high temperature conditions during the experiment. (3) The heat exchange of the reactor mainly occurred in the packing area of the two zones. Increasing packing height could effectively promote heat transfer. With only small packing heights (h1 = 0.16 m, h2 = 0.2 m), two fluids with large temperature differences could quickly complete the heat transfer process. (4) In general, different from the packet void ratio, the particle void ratio in the reaction zone had a small effect on the heat transfer process. When vg1,in and vg2,in increased, the changing degree of the solid-wall heat transfer coefficients was about 5 times that of the fluid-solid heat transfer coefficients. Only when the inlet gas temperature differences between the two zones were sufficiently large (such as 600℃) could the effect of h2 on the gas-solid fluid temperature in DTR be clearly observed. Acknowledgments The authors are grateful to the Project of Beijing Municipal Science and Technology Commission of People's Republic of China (Z161100001316004)for providing financial support for this study. References [1] M. Cheng, Y. Li, Z.S. Li, N.S. Cai, An integrated fuel reactor coupled with an annular carbon stripper for coal-fired chemical looping combustion, Powder Technol. 320
9
Contents lists available at ScienceDirect
Chemical Engineering & Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep
Steady-state simulation of internal heat-transfer characteristics in a double tube reactor
T
Yong Baia, Hui Sia, , Xiao Wangb ⁎
a b
School of Technology, Beijing Forestry University, No. 35 Tsinghua East Road Haidian District, Beijing 100083, PR China Chinese Academy of Forestry, Houxiangshan Road, Summer Palace, Haidian District, Beijing 100091, PR China
ARTICLE INFO
ABSTRACT
Keywords: Double tube reactor Fixed bed Fluidized bed Heat transfer Numerical simulation
A novel annular reactor-double tube reactor (DTR) was utilized innovatively to provide thermal energy for thermochemical reactions. A steady-state heat transfer model of DTR was established where classical heat transfer equations between the fixed bed, the fluidized bed and the wall were considered. The independence analysis showed that the different number of meshes was irrelevant. Compared with the experimental values, it was found that the calculated values of the model deviated within +15%, which illustrated good accuracy for temperature prediction. Based on the model, the numerical simulation of DTR showed that the heat exchange mainly occurred in the area with packing through the wall on both zones. Particle diameter in the reaction zone had a greater effect on heat transfer than the particle void ratio, and the bed-wall heat transfer coefficients were more sensitive to the increase of inlet gas velocities. Heat transfer would achieve quickly in DTR with only small packing heights. When temperature differences between the inlet gases of the two zones reached 600℃, the effect of the initial packing height on the fluid and solid temperature in DTR could be significantly observed, indicating the importance of the annular fixed bed in strengthening the heat transfer process.
1. Introduction Annular reactors with the advantages of simple structure, excellent heat and mass transfer properties have been widely used in chemical and industrial production in the last four decades [1]. For most thermochemical reactions, a stable reaction environment and a continuous heat source are required. Currently, there are two main methods to implement this process in annular reactors. One method is to use the annular zone of the annular reactors as the reaction zone and the inner zone as the heating zone [2]. Shedid et al. developed an annular reactor to heat an internal fluidized bed using an electric heater [3]. Berruti et al. utilized a built-in cylindrical fluidized bed as the combustion chamber and the reaction zone as an annular zone [4]. However, the biggest problem of this method in practice was that the control of the internal heating zone was difficult, resulting in a large temperature gradient in the temperature of the annular zone. The phenomenon that the uniformity of the distribution of the feedstock in the annular zone was serious[5]. Another method was to use the annular zone as the heating zone and the inner zone as the reaction zone. A number of studies have found that the best form of heat transfer in annular reactors was to transfer heat from the outside to the inside. Qiu et al. developed a new type of annular combustion chamber to improve the ⁎
air permeability and provide enough space to arrange the heat transfer surface [6]. Wang et al. designed a double-fluidized bed reactor filled with quartz sand between two concentric cylinders and studied its heat transfer characteristic [7]. Husan et al. developed a circular doubletube moving bed reactor for studying the chemical cycle gasification reaction of iron-based oxygen carriers to carbon [8]. Besides, Husan et al. had used rice husk as solid fuel for chemical cycle combustion in a circular double-tube moving bed reactor, and verified the internal temperature distribution with experimental data [9]. Alarifi et al. designed a new type of circular multi-tube reactor for the production of methanol and carried out its steady-state simulation numerically [10]. In recent decades, many studies have focused on the development of annular reactors, instead of the form of heating. According to the literature [11], one method was to incorporate a combustion chamber into the reaction zone. However, it severely affected the hydrodynamic properties in the reaction zone and increased the inconvenience of packing. Another method was to equip the periphery of the reaction zone with an electric heater, which had been widely used for experimental research. However, electric heating consumed more energy than combustion heating. Heat dissipation in electric heating was severe and it was almost impossible to utilize excessive heat [12]. To solve this problem, a double tube reactor (DTR), essentially an
Corresponding author. E-mail address: [email protected] (H. Si).
https://doi.org/10.1016/j.cep.2019.107642 Received 9 May 2019; Received in revised form 20 August 2019; Accepted 21 August 2019 Available online 30 August 2019 0255-2701/ © 2019 Published by Elsevier B.V.
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Nomenclature
tm v
Acronyms DTR
Creek Letters
Double tube reactor
αfs αws
Symbols Ags1 Ags2 Asw1 Asw2 As CF c d h lw K Nu Pr Q Re r1 r2 T tc
Logarithmic mean temperature, ℃ Velocity, m/s
Outer surface area of vessel wall immersed in bed, m2 Outer surface area of vessel wall, m2 Surface area of particles in reaction zone, m2 Surface area of particles in annular zone, m2 Effective heat transfer area in the bed phase, m2 Particle shape factor, − Specific heat capacity, J/(kg·℃) Diameter, mm Packing height, m Inner tube length, m Overall heat transfer coefficient in DTR, W/(m2·℃) Nusselt number, − Prandtl number, − Energy synthesis of all grid cells, W Reynolds number, − Radius of the reaction zone, mm Radius of the annular zone, mm Temperature, ℃ Contact time, s
φg φg,0 λ δ ρ μg μw1
Heat transfer coefficient between gas and solid, W/(m2·℃) Heat transfer coefficient between the vessel and solid, W/ (m2·℃) Particle void ratio, − Packet void ratio, − Conductivity, W·(m·℃) Wall thickness, m Density, kg/m3 The dynamic viscosity of the fluid, Pa·s The dynamic viscosity of the fluid at wall's temperature, Pa·s
Subscripts f g in out s w 1 2
annular reactor, was designed to replace electrical energy by the heat of the combustion gases. In this reactor, the reaction zone was a fluidized bed with excellent heat transfer and reaction characteristics; the annular zone was a fixed bed with almost no packing-back phenomenon, where the fluid and packing were in full contact [13]. By using a combustion chamber instead of an electric heater, a high temperature environment in the annular fixed bed could be achieved, thereby further significantly improving internal heat transfer. Therefore, the annular fixed bed played an important role in heat transfer because it determined the heat transfer coefficient largely. However, few studies have reported the principles and characteristics of such reactors above, especially the heat transfer behavior. Based on the boundary layer theory, the heat transfer research method was only applied to a single fluidized bed or a fixed bed reactor to obtain internal heat transfer coefficients, temperature distribution and fluid characteristics under the constant wall temperature condition [7,10]. There was little research on the heat transfer between the annular and the reaction zones in an annular reactor [3,5,7,11,12]. In other studies, experiments and numerical simulations were only used to predict temperature changes at specific locations in reactors at different conditions to verify the rationality of the reactor structure, with little consideration for temperature distribution, heat transfer coefficients between fluids and overall heat transfer coefficients [1–3,5,9].Moreover, to the best of author's knowledge, there was little literature on the fixed bed as the heating zone for an annular reactor. Therefore, in order to obtain good operating parameters of this reactor, especially for heat transfer characteristics, some fundamental investigations must be conducted to demonstrate the efficiency of heat transfer between the fixed bed and the fluidized bed. The purpose of this paper was to study the heat transfer characteristics of DTR. First, the principle and structural design of DTR were described. It used the annular zone (a fixed bed) and the cylindrical zone (a fluidized bed) as the heating zone and the reaction zone respectively. It was our innovative design for the structure of annular reactors. Then, a first attempt on modeling the internal heat transfer of DTR was presented and the model was verified through experimental
Fluid Gas Gas inlet parameters gas outlet parameters Solid The vessel wall Parameters in reaction zone Parameters in annular zone
data. Finally, based on the model coupled with classical heat transfer and energy equations, the numerical simulation of DTR was carried out by the commercial CFD software-Fluent V6.3.26, and the effects of various operating parameters on the internal heat transfer of DTR were studied. 2. Experimental setup 2.1. The double tube reactor The structure of DTR was shown in Fig.1.The inner tube acted as the reaction zone for thermochemical reactions, and the pre-filled quartz sand particles were in a fluidized state by air through an air distribution plate at the bottom of DTR while the annular zone filled with many Pall rings was used as the heating fixed bed. They could increase the residence time of high-temperature combustion gas and continue to heat the inner tube. A combustion furnace was tangentially mounted on the annular zone to generate combustion gas. With a spiral plate heat exchanger located below the DTR, the excess heat energy of the combustion gas out of the annular zone could preheat the fluidizing gas entering the reaction zone. The circulation of combustion gas in the two zones was achieved by a return tube. The design and packing specifications of the DTR have been summarized in Table 1. 2.2. Apparatus The DTR above and other assisted devices were used to determine the operating conditions required to achieve the objectives of the study. The experimental apparatus was designed and constructed as shown in Fig.2. The liquefied petroleum gas tank supplied petroleum gas to the furnace and the gas pressure was controlled through a pressure regulator. 50kPa- compressed air exited from the compressed gas tank into the two branches with independent gas flow was adjusted by ball valves, and the flowmeters were used to measure amount of the flow rate. One branch provided compressed air as a fluidized carrier gas for the reaction zone. The air flowed through the distribution plate to 2
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Table 1 Design and Packing Specifications of the DTR. Parameter
Value
inner tube diameter outer tube diameter inner tube length outer tube length tubes material others material packing type
180 mm o.d., 170 mm i.d. 273 mm o.d., 263 mm i.d. 1.4m 1.55m 310S 304 inner tube: quartz sand; outer tube: Pall rings 0.256 mm 25 × 0.3 × 0.3 mm SiO2 310S
quartz sand particles specification Pall rings specification quartz sand material pall rings material
of DTR, the both zones needed to be divided differently. Along the axial direction, the reaction zone was essentially a type of bubbling bed, which could be divided into a dense bed phase and a freeboard. The annular zone belonged to a fixed bed, which was divided into a packing area and an unpacking area, as shown in Fig.3. Since DTR was a revolving body, it was examined to be a simplified model in which the zone on one side of the axis was used as the control body. There were many particles in the dense bed phase of the reaction zone with many bubbles rising during the fluidization process. Its height was related to the parameters of fluidized gas velocity and particles material characteristics. The heat transfer efficiency of the fluidized gas, the inner wall surface and the movement of the particles was closely related. In the freeboard, there were fewer particles where convective heat transfer mainly between the gas and the wall occurred, and temperature gradient was generated in the radial direction. The packet renewal theory was used to divide the fluidized bed in the reaction zone into the core area and the thermal boundary layer [7]. Different from the core area, the thermal boundary layer had a small particle layer where particles in an irregular motion state collided with the wall in the form of "packets" or “clusters” and heat exchange occurred rapidly, resulting in an intense temperature gradient [14]. The packing in the annular zone was fixed and the heat transfer process in the unpacking area was the same as the freeboard in the reaction zone. 3.2. Heat transfer inside DTR 3.2.1. Heat transfer in the reaction zone The heat transfer of fluidized bed in the reaction zone included (1) heat exchange between gas and packing in bed phase and (2) heat transfer between packing and vessel wall in bed phase and (3) convective heat exchange of gas in the freeboard of the reaction zone and the unpacking area of the annular zone. For heat exchange between gas and packing in bed phase, gas to solid heat transfer coefficient αgs could be solved by Gunn's empirical equation (adapted to fluidized bed and fixed bed) [15]:
Fig. 1. Schematic diagram of DTR: (1) return tube; (2) annular distributor; (3) quartz sand particles; (4) spiral heat exchanger; (5) circular distributor; (6) feedstock inlet;(7)Pall rings; (8) flange; (9) reaction zone; (10) annular zone.
produce a uniform air flow into the DTR, which ultimately exited as heated air. The other one provided compressed air for the furnace. NiCr/Ni-Al thermocouple probes (type K) were used to measure the temperature of the fluid throughout DTR. The temperature signals were collected by an A/D converter. The measurements were monitored and recorded by SIEMENS S7-200 PLC software to calculate the internal heat transfer coefficient.
Nu = (7
10 + 5 2)(1
0.7Re 0.2Pr 1/3) + (1.33
2.4 + 1.2 2) Re 0.7Pr 1/3 (1)
3. Model development
where Re was the Reynolds number of gas, Pr was Prandtl number of gas in bed phase, φ was the bed void ratio of fluidized bed in dense phase. The effective heat transfer area As of bed phase could be calculated by the specific surface area of the packed bed or the total area of the particles per unit volume, assuming particles were spherical [16]:
3.1. Heat transfer model of DTR
As = V
For this reactor structure, it was necessary to analyze the heat transfer in different parts. Due to the different properties of the packing
For heat transfer between packing and vessel wall in bed phase, its heat transfer coefficient αsw could be calculated by conventional packet 3
6(1 ds
) (2)
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Fig. 2. Schematic diagram of the experimental apparatus.
Fig. 3. Physical model of heat transfer inside the DTR.
renewal model [14]: sw
=2
capacity, φg,0 was the bed void ratio stationarily. The effective thermal conductivity was the most difficult parameter to determine in the fluidized bed-to-wall heat transfer model. There were various derivation formulas in the literature works. Among them, the formula proposed by Zehner was mostly applied [18]. The formula also considered the effect of the relative thermal conductivity of the solid phase and gas phase:
(3)
pc /( tc )
where λ was the effective thermal conductivity of bed phase, ρ was the density of bed phase, c was effective specific heat capacity, tc was the particle contact time, characterizing the average period of contact of a single particle with vessel wall. And λ, c, ρ represented the physical properties of the whole bed phase, instead of the physical properties of solid particles, and their relationship was as follows[17]:
=
s
+
g
1-
g,0
where
(4)
s
=
where ρs was the solid particle density, cs was the solid particle specific heat capacity, ρg was the gas density, cg was the gas specific heat
g
= (1
c = (1-
g,0 ) s c s
+
g,0 s cs
(5)
4
[ A + (1 1-
g,0 ) g,0
) ]
g
(6) (7)
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
(8)
= 7.26 × 10-3
3.2.3. Energy equations According to the above assumption, the fluidized bed dense phase and the fixed bed packing area had the following relationships (subscripts 1, 2 indicated the reaction zone and the annular zone, the same as below):
and
=
2
(A B ) A
(1
1) B A ln A B
B 2 ) A
(1
(B
1)
(1
B ) A
B+1 2
(9)
where
A=
s
(10)
g
B = CF (
1-
g,0 10/9 )
where λs and λg were the thermal conductivities of solid and gas respectively, and CF was the particle shape factor. For convective heat exchange of gas in the freeboard of the reaction zone and the unpacking area of the annular zone, its heat transfer coefficient α1 could be calculated by the equation of forced convection heat transfer in the tube [11]: 1
=
f
d1
Nu
f
2 1
dt
1/3
d1 l
0.14
µ1
(13)
µ w1
(20)
Q2 =
gs 2 As 2 (Ts2
Tg 2)
(21)
Q2 =
sw 2 Ab2 (Tg 2
Ts2)
(22)
2 w lw (Tw1 ln(r2/ r1)
Tw2)
(b ) y
+
4d p L
+(
(15)
(16) 2)
1
B
0.26 0.216
= 2.58
dp G µ
1/3
cp µ
1/3
dp G µ
0.4
cp µ
(25)
g 2, in vg 2, in
=
g 2, out vg 2, out
(26)
K=
(17)
+ 0.094
g1, out vg1, out
Tg1, out )
(27)
Ts2, in)
(28)
According to the packet renewal theory [21], the heat transfer efficiency between the fluidized bed and the wall surface in the reaction zone was mainly affected by the replacement rate of packets near the wall, which was described by contact time. The parameters were assumed as follows: Tg1, in = 100℃, Tg2, in = 200℃；h1 = 0.08 m, h2 = 0.12 m；vg1,in= vg2,in = 0.5 m/s；φg = 0.44；ds = 0.256 mm [7]. By traversing tc within [0.1, 10 s], the relationship between the overall heat transfer coefficient and the contact time of the packet in the reaction zone was obtained. Divide the DTR model with quadrilateral meshes by Fluent V6.3.26, for it consisted essentially of two quadrilateral faces. The quadrilateral mesh had fast production speed, good quality, simple data structure, smooth area, and the simulation result was closer to the true value [22–25]. Different grid numbers were selected and the grid independence was verified. At the same time, the calculation results were compared, as shown in Fig.4. The overall heat transfer coefficient was calculated using the following Eq. (29):
s
s
=
3.3. Model grid independence test
hf d w2 2dt dp
g1, in vg1, in
where in and out represented inlet and outlet.
(14)
dt2 Pr Rep
(23) (24)
Q = cg 2 mg 2 (Tg 2, out
Q A1 tm
(29)
where Q was the energy synthesis of all grid cells. In Fig.4, the prediction results of the five grid numbers (200, 400, 600, 800 and 1000) showed almost no difference. Considering the calculation accuracy and cost, the subsequent calculations latter used a medium grid number (600). It can be seen that K decreased as the contact time increased, which was in agreement with the theory [21]. The faster the particle packet moved, the higher the update frequency and the greater the temperature gradient between the fluidized bed and
where hw was the heat flow coefficient for wall, it could be calculated by the following Eq. (28) [20]
hf d p
Tw1)
Q = cg1 mg1 (Tg1, in
where y and b both were the dimensionless numbers, and Φ represented the influence of the fluid film on the heat transfer at the contact point of the packing. Both a1 and Φ(b) were functions of b, and Φ1 and Φ2 could be found in the manual[19]:
2
sw1 Ab1 (Ts1
For the second and third cases, the Eq. (1)could be used as an alternative. Meanwhile, according to the conservation of mass and energy, the following equations could be obtained [7]:
3.2.2. Heat transfer in the annular zone Heat transfer in the packing area of the fixed bed in the annular zone mainly included (1) heat transfer between fluid and packing and (2) contact heat conduction - packing contacted with vessel wall surface and (3) internal heat transfer between the packing. For heat transfer between fluid and packing, the convective heat transfer coefficient αgs could also be obtained by Eq. (1) For contact heat conduction - the packing contacted with the vessel wall surface, heat transfer coefficient hp could be calculated by Eq. (14) [15]:
=
Q1 =
Q = Q1 = Q 2
where μg was the dynamic viscosity of the fluid, μw1 was the dynamic viscosity of the fluid at wall's temperature, d1 was the tube inner diameter and l was the tube length.
b=
(19)
and
(12)
Nu = 1.86 Re Pr
y=
Ts1)
Q=
and
hp =
gs1 As1 (Tg1
The above four equations could be solved in series only when there was packing in both zones. According to the presence or absence of packing, it could be divided into three types: both of the two zones had packing, only one side had packing and no packing each side. For the first case, it could be solved by adding the wall internal heat conduction correlation [9]:
(11)
g,0
Q1 =
0.4
(18)
For internal heat transfer between the packing, the heat transfer coefficient was the thermal conductivity of the material. 5
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
distribution and overall heat transfer coefficient of different parts at specific working conditions. 4.1. Axial temperature distribution in DTR The simulated values of the axial gas temperature distribution inside the reaction zone and annular zone of DTR at different inlet velocities and packing heights was shown in Fig.6. It could be seen that at the assumed condition (Tg1, in = 300℃, Tg2, in = 800℃, ds = 0.256 mm, φg = 0.44), as the axial position increased, T f1 and T f2 respectively increased and decreased rapidly, and reached a steady state finally. The rapid change in fluid temperature was mainly ascribed to the heat transfer between the gas, wall and packing, which improved the heat transfer efficiency. Comparing Fig.6. (a) and (b), vg1,in and vg2,in increased from 0.3m⋅s−1 to 0.5m⋅s-1, it could be seen that the temperature difference between the two zones at steady state increased from 50℃ to 100℃, which attributed to the larger average temperature difference between the heat transfer and the smaller convective heat transfer coefficient between the fluid and the wall when vg1,in and vg2,in increased, indicating that heat exchange existed, but the efficiency was significantly reduced. Similarly, comparing Fig.6.(b) and (c), h1, h2 respectively were increased to 0.08 m, 0.12 m to 0.16 m, 0.2 m, it could be found that the temperature difference between the two zones tended to zero, indicating that the two zones has sufficient heat transfer and reached thermal equilibrium. When h1, h2 increased, the contact frequency between the wall and packets increased in the reaction zone, resulting in a significant increase in the average convective heat transfer coefficient of the solid phase [21].
Fig. 4. Relationship between overall heat transfer coefficient and contact time of particle packet in reaction zone under different grid numbers.
the wall surface in the reaction zone, which intensified the heat transfer efficiency. The overall heat transfer coefficient curve had two turning points (A and B) and the curvature decreased. This phenomenon was caused by the ratio change of fluidized bed-to-wall heat transfer coefficient, the fixed bed heat transfer coefficient and the gas-solid heat transfer coefficient. Combining Eqs. (19)–(24), the formula for calculating the overall heat transfer coefficient was as follows:
K=
1 d2 gs1 d1
+
d2 gs2 d1
+
d2 sw1d1
+
1 sw2
+
d2 d1
(30)
where αgs1、αgs2、αsw1 and αsw2 were variables, so K was sensitive to small values. Since αsw1 and αsw2 were affected by tc, as tc increased, the decreasing trend of K would be more obvious. In Fig.4.the A and B points happened respectively when the bed-to-wall heat transfer coefficient was equal to the gas-to-solid heat transfer coefficient and the fixed bed heat transfer coefficient.
4.2. Effects of operating parameters on internal heat transfer in DTR Operating parameters, including inlet gas velocity, packing specification, packing height and others, also affected the heat transfer of DTR. Therefore, these parameters on the heat transfer characteristics of DTR were investigated next.
3.4. Experimental validation
4.2.1. Bed void ratio in the reaction zone In the packet renewal theory, the bed void ratio included the particle void ratio and packet void ratio in a fluidized bed. Therefore, the effect of bed void ratio in the reaction zone on K was investigated firstly. According to the division of the model grid, the particle void ratio was different from the packet void ratio which was only equal to the particle void ratio stationarily [26]. By traversing φg and φg,0 within [0.4, 0.8] while fixing other parameters(h1 = 0.08 m, h2 = 0.12 m；vg1,in= vg2,in = 0.5 m/s；tc1= tc2 = 5 s; ds = 0.256 mm), the relationship of K with φg and φg,0 of the reaction zone at different inlet temperatures was illustrated in Fig.7. and Fig.8. respectively. It could be seen that φg had little effect on K, whose change interval was relatively small, while K decreased as the increase of φg,0. K was mainly affected by the fluidized bed-wall heat transfer coefficient and gas-solid heat transfer coefficient
The heating test of DTR was carried out by using the apparatus in section 2.2 When the heating process reached a steady state, the temperature of the gas outlet of the reaction zone was recorded. And the gas and solid in the reaction zone were air and quartz sand whose properties were listed in Table 2. with operation parameters, considering our previous research results [7] and experimental data in some literature [3,8,9]. In order to verify the reliability of the model, the calculated reactor outlet temperature values were compared with some more experimental values in Fig.5. The deviation of the predicted value from the measured value was within 15%, indicating that the model had high reliability. The overestimation of the calculated values may be attributed to the thermal energy dissipation, dynamic energy dissipation during the experiment, which began to become apparent at over 200℃. In addition, the model needed to be further improved, such as considering the effect of friction and uneven distribution of packing or fluid. Considering the predicted values of the outlet temperature, the error of DTR was lower than that of some devices, such as a reactor with annular double fluidized beds in [7], indicating that the heat transfer mode of DTR had certain advantages.
Table 2 Operation Parameters of the Simulation and Experiment.
4. Internal heat-transfer simulation The thermal mathematical model of DTR was completed by constructing the physical model and establishing the internal heat transfer relationship of DTR by means of the classical heat transfer and energy equations. It has been verified that the established model has high reliability and could predict heat transfer behavior better. This section used numerical simulation to predict parameters such as temperature 6
Parameter
Value
gas density, kg/m3 solid density, kg/m3 gas heat capacity, J/(kg·℃) solid heat capacity, J/(kg·℃) gas conductivity, W·(m·℃) solid conductivity, W·(m·℃) mean particle diameter, mm bed void ratio stationarily,— static height in reaction zone, m height in annular zone, m inlet gas velocity, m/s
1.225 2500 1008 840 0.025 7.6 0.256 0.44 0.08 0.12 0.1—0.8
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Fig. 5. Comparison of the outlet temperatures of reaction zone between calculated and experimental values.
simultaneously, but the φg only affected the gas-solid heat transfer coefficient, so K changed very little. The increase of φg,0 led to a decrease in the effective thermal conductivity in the bed phase, and K was positively correlated with it. In addition, φg,0 made the fluidized bedwall heat transfer coefficient smaller than the gas-solid heat transfer coefficient in the interval [0.4, 0.8]. Therefore, K was more sensitive to the fluidized bed-wall heat transfer coefficient. 4.2.2. Particle diameter in the reaction zone The effect of the particle diameter in the reaction zone on K was demonstrated in Fig.9(keeping h1 = 0.08 m, h2 = 0.12 m, Tg1, in = 300℃, Tg2, in = 800℃, φg = 0.44). It could be seen that the simulation and experimental results had almost the same trend, and K decreased significantly as dsincreased within [0.2, 1 mm]. At the same inlet gas velocity, the turbulence of small-diameter particle inside DTR was larger, and the collision frequency of particles and wall was higher, resulting in the stronger heat transfer. In addition, as the inlet gas velocity increased, K increased at the same particle diameter, but the difference was not significant. This was because small-diameter particles (the minimum fluidization velocity was lower) were more easily fluidized than large-diameter particles. Small-diameter particles encountered little resistance and got more kinetic energy from heated air, taking more heat away from the wall surface after colliding with the wall. In addition, the experimental results were lower than the simulation results, probably due to the thermal energy dissipation, dynamic energy dissipation. 4.2.3. Inlet gas velocity The relationship of heat transfer coefficient and different gas inlet velocities (vg1,in= vg2,in) was showed in Fig.10 (keeping h1 = 0.08 m, h2 = 0.12 m；Tg1, Tg2, φg = 0.44, in = 300℃, in = 800℃, ds = 0.256 mm). It could be seen that αfs1, αfs2, sw1 and αsw2 all increased with the increase of inlet gas velocities. The increase in fluid -solid heat transfer coefficients (αfs1 and αfs2) was attributed to the increase of vg1,in and vg2,in, resulting in the increase of the gas Reynolds number and more severe turbulence. Relatively speaking, the solid-wall heat transfer coefficients (αsw1 and αsw2) were more sensitive to vg1,in and vg2,in, increasing from 500 to 1700 W/(m2·℃) when vg1,in and vg2,in, increased from 0 to 1 m/s. As the vg1,in and vg2,in, increased, the average temperature difference between wall and fluid heat transfer decreased, and the convective heat transfer coefficients between the wall and the fluid increased, resulting in larger turbulent of the particles in the fluidized bed, which enhanced heat transfer between the particles.
Fig. 6. Axial distribution of gas temperatures in the annular zone (T f1)and reaction zone (T f 2) (a) h1 = 0.08 m, h2 = 0.12 m, vg1,in= vg2,in = 0.3m⋅s−1, (b) h1 = 0.08 m,h2 = 0.12 m, vg1,in= vg2,in = 0.5m⋅s−1, (c) h1 = 0.16 m,h2 = 0.2 m, vg1,in= vg2,in = 0.5m⋅s−1.
usually fixed to study the variation of the heat transfer effect with the height (h2) in the annular zone. The results were showed in Fig.11. As h2 increased, the temperature of the gas and solids in the reaction zone increased while that in the annular zone decreased. The reason may be that heat was transferred from the annular zone to the reaction zone, and the packing improved the heat transfer efficiency. Comparing Fig.11. (a) and (b), when the temperature difference between the two zones was increased from 400 to 600℃, the temperature at a steady state of Tf1, out and Tf2, out was reduced from about 500 to 400℃, while the temperature difference of Ts1 and Ts2 became larger. As shown in 11(b), it can be seen that the temperature gradient of the packing increased significantly, while the gas temperature did not change significantly. The above phenomena proved that the increase of the packing height in the annular zone could promote the heat exchange
4.2.4. Packing height in the annular zone Similarly, the heat transfer coefficient was calculated at different packing heights. In practice, the packing height in the two zones was often determined by the process requirements. Therefore, h1 was 7
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
Fig. 10. Effect of inlet gas velocities on heat transfer coefficient of fluid and solid.
Fig. 7. Relationship between the overall heat transfer coefficient and the particle void ratio in the reaction zone at different inlet temperatures (Tf1,in: inlet fluid temperature in reaction zone; Tf2,in: inlet fluid temperature in annular zone).
Fig. 8. Relationship between the overall heat transfer coefficient and packet void ratio in the reaction zone at different inlet temperatures.
Fig. 11. Effect of the packing height in the annular zone on the internal temperatures of fluid and solid (a) Tg1, in = 400℃, Tg2, in = 800℃, h1 = 0.08 m, (b) Tg1, in = 200℃, Tg2, in = 800℃, h1 = 0.08 m.
Fig. 9. Effect of particle diameter in reaction zone on the overall heat transfer coefficient at different inlet gas velocities compared by experiment results.
efficiency between the two zones, which could quickly reach the thermal equilibrium. Besides, the effect of the packing height in the annular zone on the overall heat transfer coefficient was obtained, as shown in Fig.12. It can be seen that K increased as h2 increased and K increased with v1,in and v2,in.The reason may be that the annular zone was a fixed bed, and the gas can only pass through their gap during the period. As the height of the packing increased, the gas-solid heat exchange capacity increased and the “insulation” function of the entire annular zone became stronger. The increase of vg1,in and vg2,in caused the relatively lowtemperature gas after heat exchange to leave the annular zone quickly and the high-temperature gas was quickly replenished. The heat loss
Fig. 12. Effect of packing height in annular zone on overall heat transfer coefficient at different inlet gas velocities. 8
Chemical Engineering & Processing: Process Intensification 144 (2019) 107642
Y. Bai, et al.
caused by the shortage of high-temperature gas and the slow increase of K were reduced, so the heat exchange rate between the two zones was increased slowly.
(2017) 519–529. [2] T.J. Brooms, M.S. Onyango, A. Ochieng, Photodegradation of phenol using TiO2, ZnO and TiO2/ZnO catalysts in an annular reactor, J Water Chem Technol. 39 (2017) 155–160. [3] M.H. Shedid, M.A.M. Hassan, Heat transfer characteristics of the fluidized bed through the annulus, Heat Mass Transfer. 52 (2016) 1943–1952. [4] F.M. Berruti, C.L. Briens, Novel intermittent solid slug feeder for fast pyrolysis reactors: fundamental and modeling, Powder Technol. 247 (2017) 95–105. [5] J.S.M. Botterill, M. Desai, Limiting factors in gas-fluidized bed heat transfer, Powder Technol 6 (4) (1972) 231–238. [6] G. Qiu, J. Ye, H. Wang, Investigation of gas−solids flow characteristics in a circulating fluidized bed with annular combustion chamber by pressure measurements and CPFD simulation, Chem. Eng. Sci. 134 (2015) 433–447. [7] X. Wang, H. Si, Novel concept of fluidized bed reactor design for highly efficient internal heat transfer, Ind. Eng. Chem. Res. 55 (2016) 9276–9283. [8] C.W. Hsuan, K. Young, Chemical Looping Gasification of Charcoal with Iron-Based Oxygen Carriers in an Annular Dual-Tube Moving Bed Reactor, Aerosol air quality research 16 (2015) 1093–1103. [9] C.W. Hsuan, K. Young, T. Hsin-Hsien, K. Yu-Lin, T. Yao-Hsuan, Rice husk as solid fuel for chemical looping combustion in an annular dual-tube moving bed reactor, Int. J. Chem Eng. 280 (2015) 82–89. [10] A. Alarifi, A. Elkamel, E. Croist, Steady-state simulation of a novel annular multitubular reactor for enhanced methanol production, Ind. Eng. Chem. Res. 52 (2013) 15387–15393. [11] H. Taofeeq, M. Al-Dahhan, Heat transfer and hydrodynamics in a gas-solid fluidized bed with vertical immersed internals, Int. J. Heat Mass Transfer 122 (2018) 229–251. [12] A. Singhal, S. Cloete, S. Radl, F.R. Quinta, S. Amini, Heat transfer to a gas from densely packed beds of cylindrical particles, Chem. Eng. Sci. 172 (2017) 1–12. [13] H.A. Aghdas, S.B. Ralph, Y.A. Habibollah, Effect of process conditions on product yield and composition of fast pyrolysis of Eucalyptus grandis in fluidized bed reactor, Ind. J. Eng. Chem. 7 (2014) 2594–2602. [14] R. Zarghami, N. Mostoufi, R. Sotudeh-Gharebagh, J. Chaouki, Analysis and modeling of Particle-wall contact time in gas fluidized beds, Chem Eng Sci. 62 (2007) 4573–4578. [15] D.J. Gunn, Transfer of heat or mass to particles in fixed and fluidised beds, Int. J. Heat Mass Transfer. 21 (1978) 467–476. [16] Cormos, Dynamic simulation of fluidized bed chemical looping combustion process with iron based oxygen carrier, Fuel 214 (2018) 436–445. [17] H.M. Abdelmotalib, M.A. Youssef, Im, Heat transfer process in gas–solid fluidized bed combustors: a review, Int J Heat Mass Transfer. 89 (2015) 567–575. [18] P. Zehner, E.U. Schlünder, Wärmeleitfähigkeit von Schüttungen bei mäßigen Temperaturen, Chem. Ing. Tech 42 (1970) 933–941. [19] Z.W. Zi, Petrochemical Design Manual, Chemical Industry Press, Beijing, 2015 (in Chinese). [20] A. Kumar, H.M. Jena, Removal of methylene blue and phenol onto prepared activated carbon from Fox nutshell by chemical activation in batch and fixed-bed column, J. Clean Prod. 137 (2016) 1246–1259. [21] S.E. Olsson, A.E. Almstedt, Local instantaneous and time-averaged heat transfer in a pressurized fluidized bed with horizontal tubes: influence of pressure, fluidization velocity and tube-bank geometry, Chem. Eng. Sci. 50 (1995) 3231–3245. [22] M. Sosnowski, J. Krzywanski, K. Grabowska, R. Gnatowska, Polyhedral meshing in numerical analysis of conjugate heat transfer, EPJ Web of Conferences vol. 180, (2018) 02096. [23] M. Sosnowski, J. Krzywanski, R. Gnatowska, Polyhedral meshing as an innovative approach to computational domain discretization of a cyclone in a fluidized bed CLC unit, E3S Web of Conferences vol. 14, (2017) 01027. [24] M. Spiegel, T. Redel, Y.J. Zhang, T. Struffert, J. Hornegger, R.G. Grossman, C. Karmonik, Tetrahedral vs. polyhedral mesh size evaluation on flow velocity and wall shear stress for cerebral hemodynamic simulation, Comput Method Biomec 14 (2011) 9–22. [25] J. Su, Z. Gu, C. Chen, X.Y. Xu, A two-layer mesh method for discrete element simulation of gas‐particle systems with arbitrarily polyhedral mesh, Int J Numer Meth Eng 103 (2015) 759–780. [26] H.S. Mickley, D.F. Fairbanks, Mechanism of heat transfer to fluidized beds, AIChE J 1 (1955) 374–384.
5. Conclusions A novel double tube reactor (DTR) was designed in which the internal reaction zone was wrapped by an annular fixed bed to supply energy for thermochemical reactions. A spiral heat exchanger was installed under the lower portion of DTR to further utilize the heat of the hot exhaust gas. The construction and principle of the reactor were elaborated. Then the heat transfer model of DTR at steady-state conditions was mathematically established. Based on the classical heat transfer and energy equations, the heat transfer process of fluidized bed and fixed bed was described. Model independence analysis and verification were also performed. Finally this model was used to predict the heat transfer characteristics of DTR. The main conclusions were as follows: (1) The independence analysis using five grid numbers (200, 400, 600, 800 and 1000) showed that the model was not affected by the number of control volume grids. Considering the accuracy and time of the calculation, a medium quantity grid (600) was finally selected. (2) The difference of the experimental and calculated values of the outlet temperature in the reaction zone was within ± 15%, which had high rationality. The error may be the heat loss of the reactor at high temperature conditions during the experiment. (3) The heat exchange of the reactor mainly occurred in the packing area of the two zones. Increasing packing height could effectively promote heat transfer. With only small packing heights (h1 = 0.16 m, h2 = 0.2 m), two fluids with large temperature differences could quickly complete the heat transfer process. (4) In general, different from the packet void ratio, the particle void ratio in the reaction zone had a small effect on the heat transfer process. When vg1,in and vg2,in increased, the changing degree of the solid-wall heat transfer coefficients was about 5 times that of the fluid-solid heat transfer coefficients. Only when the inlet gas temperature differences between the two zones were sufficiently large (such as 600℃) could the effect of h2 on the gas-solid fluid temperature in DTR be clearly observed. Acknowledgments The authors are grateful to the Project of Beijing Municipal Science and Technology Commission of People's Republic of China (Z161100001316004)for providing financial support for this study. References [1] M. Cheng, Y. Li, Z.S. Li, N.S. Cai, An integrated fuel reactor coupled with an annular carbon stripper for coal-fired chemical looping combustion, Powder Technol. 320
9