Energy Fuels 2010, 24, 5634–5651 Published on Web 09/10/2010
: DOI:10.1021/ef1008852
Geometrical Optimization of a Fast Pyrolysis Bubbling Fluidized Bed Reactor Using Comutational Fluid Dynamics K. Papadikis,† S. Gu,*,† and A.V. Bridgwater‡ †
School of Engineering Sciences, University of Southampton, Highfield, Southampton, SO17 1BJ, United Kingdom, and ‡ School of Engineering and Applied Science, Aston University, Aston Triangle, Birmingham B4 7ET, United Kingdom Received July 10, 2010. Revised Manuscript Received August 24, 2010
This paper analyzes the physical phenomena that take place inside an 1 kg/h bubbling fluidized bed reactor located at Aston University and presents a geometrically modified version of it, in order to improve certain hydrodynamic and gas flow characteristics. The bed uses, in its current operation, 40 L/min of N2 at 520 °C fed through a distributor plate and 15 L/min purge gas stream, i.e., N2 at 20 °C, via the feeding tube. The Eulerian model of FLUENT 6.3 is used for the simulation of the bed hydrodynamics, while the k - ɛ model accounts for the effect of the turbulence field of one phase on the other. The three-dimensional simulation of the current operation of the reactor showed that a stationary bubble was formed next to the feeding tube. The size of the permanent bubble reaches up to the splash zone of the reactor, without any fluidizaton taking place underneath the feeder. The gas flow dynamics in the freeboard of the reactor is also analyzed. A modified version of the reactor is presented, simulated, and analyzed, together with a discussion on the impact of the flow dynamics on the fast pyrolysis of biomass.
et al.,8 as well as the work of Li et al.,9 where the effect of the change of the volumetric flow rate due to phase change in a fluidized bed was investigated. In the study of Chiesa et al.,10 a computational study of the flow behavior of a labscale fluidized bed was performed by using both EulerianLagrangian and Eulerian-Eulerian models. The results were in very good agreement with experiments performed at a two-dimensional lab-scale bubbling fluidized bed reactor. There are several studies investigating the effect of secondary gas jets on the hydrodynamics of the fluidized beds, something that comes closer to reality, since most of the fluidized beds for fast pyrolysis use some sort of secondary gas fed into the reactor for various reasons. In many cases this gas jet is a part of the feeding process, as well as for cooling purposes and prevention of any backflow of pyrolysis vapors into the feeding system. Li et al.11 studied with threedimensional simulations in a rectangular fluidized bed the behavior of single and multiple horizontal gas jet injections. It was found that the secondary gas injection mainly affects the hydrodynamics of the upper section above the injection level, while its effect is nearly negligible below the injection. At another study of the same authors,12 a three-dimensional
1. Introduction Fluidized beds have been the center of modeling attention for many years, and various simulation approaches have been developed, ranging from Lattice-Boltzmann methods (LBM)1 to discrete particle (DPMs, Eulerian-Lagrangian),2,3 and twofluid models (TFMs, Eulerian-Eulerian).4,5 More recently, discrete bubble models (DBMs)6 have been developed and applied to fluidized beds, increasing the potential of industrial scale simulations beyond the capabilities of TFMs. Several computational models have been presented in the literature that simulate the hydrodynamics of the particulate phase under certain fluidization conditions. Most of the models concern bench scale reactors, in which the fluidizing gas is supplied through a distributor plate at the bottom of the sand bed, such as the works of Taghipour et al.7 and Pain *To whom correspondence should be addressed. Telephone: 023 8059 8520. Fax: 023 8059 3230. E-mail:
[email protected]. (1) Ladd, A. J. C.; Veberg, R. Lattice-Boltzmann simulations of particle fluid suspensions. J. Stat. Phys. 2001, 104, 1191. (2) Feng, Y. Q.; Xu, B. H.; Zhang, S. J.; Yu, A. B.; Zulli, P. Discrete particle simulation of gas fluidization of particle mixtures. AIChE J. 2004, 50 (8), 1713–1728. (3) Hoomans, B. P. B.; Kuipers, J. A. M.; Mohd Salleh, M. A.; Stein, M.; Seville, J. P. K. Experimental validation of granular dynamics simulations of gas-fluidised beds with homogeneous in-flow conditions using positron emission particle tracking. Powder Technol. 2001, 116 (2-3), 166–177. (4) Gidaspow, D. Multiphase Flow and Fluidization: Continuum and Kinetic Theory Descriptions. Academic Press: New York, 1994. (5) Enwald, H.; Peirano, E.; Almstedt, A. E.; Leckner, B. Simulation of the fluid dynamics of a bubbling fluidized bed. Experimental validation of the two-fluid model and evaluation of a parallel multiblock solver. Chem. Eng. Sci. 1999, 54, 311–328. (6) Bokkers, G. A.; Laverman, J. A.; Van Sint Annaland, M.; Kuipers, J. A. M. Modelling of large-scale dense gas-solid bubbling fluidised beds using a novel discrete bubble model. Chem. Eng. Sci. 2006, 61, 5590–5602. (7) Taghipour, F.; Ellis, N.; Wong, C. Experimental and computational study of gas-solid fluidized bed hydrodynamics. Chem. Eng. Sci. 2005, 60, 6857–6867. r 2010 American Chemical Society
(8) Pain, C. C.; Mansoorzadeh, S.; Gomes, J. L. M.; de Oliveira, C. R. E. A numerical investigation of bubbling gas-solid fluidized bed dynamics in 2-D geometries. Powder Technol. 2002, 128 (1), 56–77. (9) Li, T.; Mahecha-Botero, A.; Grace, J. R. Computational Fluid Dynamic Investigation of Change of Volumetric Flow in Fluidized-Bed Reactors, Ind. Eng. Chem. Res., DOI: 10.1021/ie901676d. (10) Chiesa, M.; Mathiesen, V.; Melheim, J. A.; Halvorsen, B. Numerical simulation of particulate flow by the Eulerian-Lagrangian and the Eulerian-Eulerian approach with application to a fluidized bed. Comput. Chem. Eng. 2005, 29 (2), 291–304. (11) Li, T.; Pougatch, K.; Salcudean, M.; Grecov, D. Numerical simulation of single and multiple gas jets in bubbling fluidized beds. Chem. Eng. Sci. 2009, 64 (23), 4884–4898. (12) Li, T.; Pougatch, K.; Salcudean, M.; Grecov, D. Numerical simulation of horizontal jet penetration in a three-dimensional fluidized bed. Powder Technol. 2008, 184 (1), 89–99.
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Figure 1. The 1 kg/h fluidized bed reactor: (left) original design, (right) modified design. All dimensions are shown in millimeters.
simulation of a single horizontal gas jet into a cylindrical gas-solid lab-scale fluidized bed was performed, investigating the jet penetration lengths of different jet velocities. Christensen et al.13 studied the distributed secondary gas injection via a fractal injector in a 3-D lab-scale fluidized bed to determine its effect on bubble size, bubble fraction, residence time, mixing, and conversion. The results showed that distributed secondary gas injection improves the mass transfer and gas-solid contact, which consequently increase the performance of the reactor. Utikar and Ranade14 performed a computational fluid dynamics (CFD) study with rectangular fluidized beds operated with a central jet. Fluidization of two types of particles, glass and polypropylene (PP), was studied at two different initial bed heights with three central jet velocities. Although the results were in good agreement with the experimental data and previously published correlations, complete agreement could not be found in some aspects. The scope of the current paper is to investigate the sand bed hydrodynamics and gas flow dynamics of an 1 kg/h fast
pyrolysis bubbling fluidized bed reactor and to propose an optimized modified version using CFD. The modification of the reactor is based on the results of the simulation of the original design in order to improve certain hydrodynamic and gas flow characteristics. The main target is to determine the physical phenomena occurring inside the reactor and to suggest solutions that could overcome certain problems that reduce the performance of the reactor. The analysis is based on the application of the specific fluidized bed reactor for fast pyrolysis purposes. 2. Model Description 2.1. Operational Parameters. The original and the modified dimensions of the 1 kg/h bubbling fluidized bed reactor can be seen in Figure 1. The rig is located at the laboratory of the Bioenergy Research Group (BERG) of Aston University. At its current state, the reactor is operated using 40 L/min of nitrogen through a distributor plate at 520 °C and 15 L/min of nitrogen at 20 °C through the feeding tube. The sand bed consists of 1 kg of quartz sand with a bulk density of 1660 kg/ m3, preheated to 520 °C. Sand particle sizes range from 600 to 710 μm. The operational parameters of the original reactor are given in Table 1. The three-dimensional simulation is based on these specific operational parameters, where for the sand particle size an average value of 650 μm has been chosen. The discussion and analysis of the geometrical modification of the reactor is discussed in section 4 and compared with the current configuration.
(13) Christensen, D.; Nijenhuis, J.; van Ommen, J. R.; Coppens, M.-O. Influence of Distributed Secondary Gas Injection on the Performance of a Bubbling Fluidized-Bed Reactor. Ind. Eng. Chem. Res. 2008, 47 (10), 3601–3618. (14) Utikar, P. R.; Ranade, V. V. Single jet fluidized beds: Experiments and CFD simulations with glass and polypropylene particles. Chem. Eng. Sci. 2007, 62 (1-2), 167–183.
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Table 1. Simulation Parameters of the Original Reactor property
value
comment
superficial velocity (distributor), U0 nitrogen temperature (distributor) secondary gas injection (feeding tube), Ujet nitrogen temperature (feeding tube) solids particle density, Fs sand specific heat capacity, Cp,s sand thermal conductivity, ks mean solids particle diameter, ds restitution coefficient, ess initial solids packing, εs static bed height bed width
0.16 m/s 520 °C 3.5 m/s 20 °C 2650 kg/m3 835 J/kg K 0.35 W/mK 650 μm 0.9 0.62 0.144 m 0.0729 m
fixed value fixed value fixed value fixed value quartz sand fixed fixed uniform distribution value in literature fixed value fixed value fixed value
Table 2. Conservation Equations of the Eulerian Model
Table 3. Constitutive Equations
gas mass conservation
solids stress tensor 2 τ s ¼ εs μs ðrυs þ rυTs Þ þ εs λs - μs r 3 υs I s 3
∂ðεg Fg Þ þ r 3 ðεg Fg υg Þ ¼ 0 ∂t solids mass conservation
bulk viscosity
∂ðεs Fs Þ þ r 3 ðεs Fs υs Þ ¼ 0 ∂t
λs ¼
gas momentum conservation
rffiffiffiffiffiffi 4 Θs εs Fs ds g0, ss ð1 þ ess Þ 3 π
solids shear viscosity
∂ðεg Fg υg Þ þ r 3 ðεg Fg υg Xυg Þ ¼ -εg 3 rp þ r 3 τ g þ εg Fg g þ Kgs ðug - us Þ ∂t
μs ¼ μs, col þ μs, kin þ μs, fr collisional viscosity
solids momentum conservation ∂ðεs Fs υs Þ þ r 3 ðεs Fs υs Xυs Þ ¼ -εs 3 rp - rps þ r 3 τ s þ εs Fs g þ Kgs ðug - us Þ ∂t
μs, col ¼ frictional viscosity
energy conservation ∂ðεq Fq hq Þ þ r 3 ðεq Fq υq hq Þ ∂t ¼ -εq
rffiffiffiffiffiffi 4 Θs εs Fs ds g0, ss ð1 þ ess Þ 5 π
μs, fr ¼
n X ∂pq þ τ q : rυq - r 3 qq þ Sq þ ðQpq þ m_ pq hpq - m_ qp hqp Þ ∂t p¼1
ps sinðφÞ pffiffiffiffiffiffiffi 2 I2D
kinetic viscosity
The simulation parameters of the modified case are exactly the same as for the original one, with the only difference being the superficial velocity of nitrogen in the distributor plate. The superficial velocity for the modified case was defined at U0 = 0.28 m/s, which comes to a volumetric flow rate of Q0 = 70 L/min. 2.2. Numerical Setup. The simulation was set according to the following parameters: (1) The geometry was discretized in 956 000 tetrahedral cells with a minimum volume of 2.892 068 10-10 and maximum volume of 8.199 684 10-9 m3. (2) Velocity inlet boundary conditions were defined for the distributor and feeding tube inlets, while the outlet was defined as the pressure outlet. At the wall of the reactor, a fixed temperature of 793 K was defined. (3) The geometry was partitioned in 14 cores Intel Xeon at 2.27 GHz, and the simulation time step was defined at 0.000 01 s.
μs, kin ¼
pffiffiffiffiffiffiffiffiffi 2 10Fs ds Θs π 4 1 þ εs g0, ss ð1 þ ess Þ 5 96εs g0, ss ð1 þ ess Þ
solids pressure ps ¼ εs Fs Θs þ 2Fs ð1 þ ess Þεs 2 g0, ss Θs radial distribution function 2 g0, ss
¼ 41 -
εs εs, max
!1=3 3-1 5
while the kinetic theory of granular flow was applied for the conservation of the solid’s fluctuation energy. Table 2 contains the conservation equations for the gas and solid phases, with the model’s constitutive equations shown in Table 3. The Gidaspow4 interphase exchange coefficient is given in Table 4, while the fluctuating energy conservation of solid
3. Mathematical Model 3.1. Multiphase Flow Governing Equations. The simulations of the bubbling behavior of the fluidized bed were performed by solving the equations of motion of a multifluid system. An Eulerian model for the mass and momentum for the gas (nitrogen) and fluid phases was applied, 5636
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Table 4. Interphase Exchange Coefficient
Table 7. Constitutive Equations for Turbulence Model
Gidaspow interphase exchange coefficient
Kgs ¼
3 εs εg Fg jυs - υg j -2:65 Cd εg 4 ds
Reynolds stress tensor T 2 Bq ÞI þ Fq μt, q ðrV Bq þ rV Bq Þ τq ¼ - ðFq kq þ Fq μt, q r 3 V 3
¼ 00
for εg > 0:8
production of turbulent kinetic energy Kgs ¼ 150
εs 2 μg εg ds
2
þ 1:75
εs Fg jυs - υg j ds
Gk, q ¼ - Fui uj
for εg e 0:8
∂ui ∂xj
interphase turbulent momentum transfer drag coefficient N X
24 Cd ¼ ½1 þ 0:15ðεg Res Þ0:687 εg Res
N X
f Þ ¼ Klq ðvfl - υ q
l¼1
f
N X
f
Klq ðVl - Vq Þ -
l¼1
l ¼1
Klq f υdr, lq
drift velocity Reynolds number
Res ¼
f υdr, lq
ds Fg jυs - υg j μg
Dq Dl ¼ rεl rεq σ lq εl σ lq εq
turbulent viscosity
Table 5. Fluctuation Energy Conservation of Solid Particles
μt, q ¼ Fq Cμ
fluctuation energy conservation of solid particles 3 ∂ ðεs Fs Θs Þ þ r 3 ðεs Fs υs Θs Þ ¼ ð- ps I s þ τ s Þ : r 3 υs 2 ∂t
τt, lq τF, lq
ηlq ¼
characteristic particle relaxation time connected with inertial effects acting on dispersed phase
diffusion coefficient of granular temperature rffiffiffiffiffiffi pffiffiffiffiffiffiffiffiffi 2 150Fs ds Θs π 6 Θs ¼ 1 þ εs g0, ss ð1 þ ess Þ þ 2Fs ds εs 2 g0, ss ð1 þ ess Þ 384ð1 þ ess Þg0, ss 5 π
τF, lq ¼ εl Fq Klq
-1
Fl þ CV Fq
!
Lagrangian integral time scale
collision dissipation energy
γΘs ¼
kq 2 εq
ratio of characteristic times
þ r 3 ðkΘs 3 r 3 Θs Þ - γΘs þ φgs
kΘs
!
τ t, q τt, lq ¼ qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi ð1 þ Cβ ξ2 Þ
12ð1 - ess 2 Þg0, ss pffiffiffi Fs εs 2 Θs 3=2 ds π
characteristic time of the energetic turbulent eddies transfer of kinetic energy τ t, q ¼
φgs ¼ -3Kgs Θs
Table 6. Turbulence Model: k - ε Model for Each Phase
length scale of turbulent eddies
turbulent kinetic energy μt, q ∂ðεq Fq kq Þ þ r 3 ðεq Fq V Bq kq Þ ¼ r 3 εq rkq þ ðεq Gk, q - εq Fq εq Þ ∂t σk þ
N X
Klq ðClq kl - Cql kq Þ -
l¼1
þ
N X l¼1
N X l¼1
Klq ðV Bl - V Bq Þ 3
Klq ðV Bl - V Bq Þ 3
L t, q ¼
Cβ ¼ 1:8 - 1:35 cos2 θ
μt, q rεq εq σ q ξ ¼
dissipation rate
N X
N X μ t, l μt, q Klq ðV Bl - V Bq Þ 3 rεl þ Klq ðV Bl - V Bq Þ 3 rεq ε σ ε l l q σq l¼1 l¼1
rffiffiffi 3 kq 3=2 Cμ 2 εq
constitutive dimensionless quantities
μt, l rεl εl σl
μt, q ∂ðεq Fq Eq Þ þ r 3 ðεq Fq V Bq Eq Þ ¼ r 3 εq rEq ∂t σE " N X Eq þ C1E Eq Gk, q - C2E εq Fq Eq þ C3ε Klq ðClq kl - Cql kq Þ kq l¼1
3 kq Cμ 2 εq
Clq ¼ 2,
!
jf υlq jτt, q Lt, q
Cql ¼ 2
ηlq 1 þ ηlq
!
particles and its constitutive equations are given in Table 5. An analytical discussion of the solid-phase properties can be found in Boemer et al.15
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Figure 2. Fluidized bed hydrodynamics at different times.
Figure 3. Jet velocity at 3 s of simulation.
The effect of turbulence of one phase on the other is modeled using the k - ɛ turbulence model. The transport equations for the kinetic energy and its dissipation rate of each phase are given in Table 6, while the constitutive terms are defined in Table 7. The model constants are equal to Cμ = 0.09, C1ɛ = 1.44, C2ɛ = 1.92, C3ɛ = 1.3. The addedmass coefficient, CV, is equal to CV = 0.5. The dispersion Prandtl number, σlq, takes the value of σlq = 0.75, while the turbulent kinetic energy and turbulent dissipation rate Prandtl numbers are σk = 1, σɛ = 1.3, respectively.
10 s. There are certain problems that one can immediately spot at a first glance on the hydrodynamic behavior of the bed. Considering the sand bed underneath the feeding tube, one can immediately notice that no fluidization takes place and it is only the gas injected from the feeder that causes a stationary bubble to be formed next to the feeding point and up to the splash zone of the reactor. This phenomenon cannot be easily predicted if someone takes into account the total volumetric flow rate of nitrogen inside the reactor. Under the operating conditions specified, the minimum fluidization velocity of the sand bed comes close to Umf = 0.19 m/s, whereas the superficial velocity of the fluidizing gas coming through the distribution is approximately U0 = 0.16 m/s. The velocity of the gas jet from the feeding tube is approximately Ujet = 3.6 m/s. Hence, someone would expect that the side gas injection would enhance the fluidization behavior of the bed. However, as the study of Li et al.11 showed, the gas jet has negligible influence on the hydrodynamics of the bed below the injection level. This phenomenon was also observed in the current simulation, since the gas jet did not actually provide any enhancement in the fluidization quality of the
4. Results and Discussion 4.1. Fluidized Bed Hydrodynamics: Original Reactor. Figure 2 illustrates the hydrodynamics of the 1 kg/h fluidized bed reactor at its original operating conditions. The sand volume fraction contours are shown for only 3 s since the flow pattern remained identical for the rest of the simulation which in total lasted for (15) Boemer, A.; Qi, H.; Renz, U. Eulerian simulation of bubble formation at a jet in a two-dimensional fluidised beds. Int. J. Multiphas. Flow 1997, 23 (5), 927–944.
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Figure 4. X-Y sand velocity components above the feeding tube.
Figure 5. Volume fraction of nitrogen at different heights above the feeding tube. Indication of jet penetration into the bubbling bed and bubble diameter.
bed. The superficial gas velocity through the distributor was less than the minimum fluidization velocity of the bed, and therefore no fluidization was observed below the level of the feeding tube. In contrast, the gas jet from the feeding tube is responsible for the stationary bubble formed next to the feeding point and up to the splash zone. Figure 3 shows the jet velocity at different heights above the feeding tube, which can locally reach 10 m/s. This has a negative influence on the mixing of the sand, since it forces the latter to move toward the walls of the reactor as it is shown by the sand velocity contours in Figure 4. All of the above phenomena result in a poor performance of the reactor since the mixing of solids is minimum and thus behaving more like a static bed instead of a bubbling bed. The size of the bubble and consequently the jet penetration length can be determined by the gas volume fraction distribution above the feeding tube as it is shown in Figure 5.
As we can see, the maximum bubble size can reach approximately 0.04 m, forming a narrow path from which the gas escapes providing no fluidization to the upper part of the bed. Another reason for this can also be the position of the feeding tube in the reactor compared to the static bed height. We can see that the bed height above the feeding tube is approximately 0.043 m while the bed height below the feeding tube is 0.101 m. This results in a minor expansion and penetration of the gas jet inside the bed. Considering the fact that no fluidization and sand mixing is provided from the lower part of the reactor through the distributor, in order to disturb the jet gas flow field, the stationary bubble formation becomes almost inevitable. One immediately can spot the weaknesses of the current configuration for fast pyrolysis applications. One of the most important features of fluidized beds is the mixing of solids that provide high heat transfer coefficients and uniform 5639
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Figure 6. Contours of nitrogen temperature from the feeding tube of the original reactor.
Figure 7. Y-velocity contours at the freeboard of the reactor.
temperature inside the reactor. In this particular case, the mixing of solids that act as heat carriers from the wall toward the center of the bed is completely lost. Therefore, the biomass particles that are going to be injected through the feeding tube will only travel along the narrow path of the stationary bubble up to the splash zone of the reactor. In this way, the contact with the solids of the bed is minimized and biomass is pyrolyzed mainly by the jet gas that comes from the feeding tube. The temperature contours of the jet gas are shown in Figure 6, and one can immediately notice that the gas needs to travel almost half way to the splash zone to reach the temperature of the bed. This is expected if one considers the weak mixing of solids and that the gas does not come in full contact with the large heat transfer area of the sand bed, since a stationary bubble has been formed. Therefore, the pyrolysis of biomass will mainly depend on convective heat transfer from the jet gas instead of the proper combination of gas convective and particle conductive effects that occur in fluidized beds. 4.2. Freeboard Gas Flow Dynamics: Original Reactor. Figure 7 shows the y-velocity contours in the freeboard of
the reactor. The gas velocity distribution has a parabolic profile at the 0.2 and 0.6 s of the simulation, something that can be also seen in Figure 8, which shows the y-velocity component of nitrogen above the splash zone of the reactor. However, as soon as the stationary bubble is formed, the velocity contours indicate a low velocity region above the splash zone which is in comparable size to the bubble. This region acts as a flow separation region and forces the ejected gas to flow around it. This phenomenon can be seen on Figure 8 where the y-velocity distribution across the diameter of the reactor is shown as well as on Figure 9 where the static pressure at different heights above the splash zone is compared. The phenomenon can be explained in exactly the same way as the boundary layer separation occurs in aerofoils. The gas that flows through the narrow path of the stationary bubble is ejected with high velocity to the freeboard where it has to expand rapidly to a larger diameter tube. From conservation of mass, the gas is slowed down to keep the mass flow constant. This causes a negative pressure gradient acting from the higher levels of the freeboard toward the splash zone where the bubble erupts. This phenomenon forces the gas to 5640
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Figure 8. Y-velocity components of nitrogen at different heights above the splash zone of the reactor.
separate and flow around this counter pressure gradient region. This has an additional effect on the final velocity of the gases in the freeboard, since the effective gas flow area is reduced and the gases are accelerated toward the outlet making the enlargement of the diameter of the freeboard less effective for gas deceleration and solids separation. This phenomenon is permanent only because a stationary bubble has been formed which continuously ejects high velocity gas
into the freeboard. The separation zone cannot be easily predicted if the sand bed hydrodynamics is not known and therefore make the simple calculations for freeboard gas velocities diverge from reality. The gas flow dynamics in the freeboard of the reactor is very significant in fast pyrolysis for various reasons. The freeboard acts as a disengaging zone for char particles, which are ejected with high velocity from the sand bed when 5641
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Figure 9. Static pressure of nitrogen at different heights above the splash zone of the reactor.
bubbles erupt. The optimum residence time of char particles is approximately 2 s, in order to avoid extensive secondary cracking of pyrolysis vapors to noncondensable volatiles. The flow separation region above the splash zone of the reactor can have a significant negative effect in the entrainment of char particles. The low gas velocities that are present in this region can act as a deceleration factor by inducing negative drag to the higher velocity char particles. This may result in the return of the char particles to the bulk of the bed
and consequently increase the residence time. This means that the char particles will have to move closest to the wall of the reactor, where higher gas velocities are present. In this case the negative effect of the stationary bubbles becomes visible, since its presence results in minor mixing of solids, making the circulation of char particles around the bed difficult. 4.3. Effect of Turbulence: Original Reactor. Figure 10 shows the turbulent viscosity ratio of nitrogen along the 5642
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Figure 10. Turbulent viscosity ratio of nitrogen along the height of the reactor, starting from the feeding tube.
Figure 11. Y-velocity component of nitrogen without turbulence model.
height of the reactor. We can see that turbulence is only significant at the inlet of the feeding tube, μr = 28, since it is the only actual fully turbulent region in the reactor as well as the separation region, μr = 7.5, and the outlet region, μr = 8, which are more in the transitional stage. This is indicative of the influence of the turbulent viscosity, which has a significant contribution on the diffusion of turbulent kinetic energy. No differences were observed on the bed hydrodynamics, even when the model was tested without the inclusion of a turbulence model. However, differences were noticeable in the separation region, where negative values of the y-velocity component of nitrogen were observed (Figure 11). In contrast to the turbulent case, this is the result of the exclusion of the effect of the small scale eddies on the flow, since the diffusive effects are solely based on molecular viscosity. Although, a small difference on the flow pattern in the separation zone was observed, the size of the zone remained approximately the same for both the laminar and turbulent
case. This resulted in a minor difference in the gas velocities in the rest of the freeboard. 4.4. Fluidized Bed Hydrodynamics: Modified Reactor. Figure 12 shows the fluidized bed hydrodynamics of the geometrically modified reactor for 5 s of simulation. Despite the fact that 5 s is a very limited time compared to the hours of operation of a bubbling fluidized bed, it is quite indicative of the flow pattern that is going to be developed. As it is shown in Figure 1, the modified version of the reactor mainly concerns three parts. The feeding tube has been brought down to 0.070 m from the distributor, which is approximately in the middle of the static bed height of the sand bed. This was done in order to provide a larger path to the side gas injection, toward the splash zone. In this way, the gas jet will have to travel a longer distance to reach the top of the bed, something that will lead to higher dissipation of the jet and effectively reduce the bubble size. Also, the increase in the superficial velocity in the distributor above the minimum fluidization velocity (U0 = 0.28 m/s) will provide extra gas bubbles 5643
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Figure 12. Fluidized bed hydrodynamics of the modified reactor at different times.
which coalesce with the central jet, making it dissipate more quickly into the emulsion and lose some of its momentum, as it is stated by Guo et al.16
A second modification was to create an enlargement in the diameter of the reactor exactly at the top of the static bed. In this way, the sand will be able to flow in a transverse direction and collide with the wall of the reactor when the bed expands due to bubble formation. This modification takes advantage of the momentum and inertia gained by the sand bed in the
(16) Guo, Q.; Liu, Z.; Zhang, J. Flow Characteristics in a Large Jetting Fluidized Bed with Two Nozzles. Ind. Eng. Chem. Res. 2000, 39 (3), 746–751.
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Figure 13. Volume fraction of nitrogen at different heights above the feeding tube of the modified reactor. Indication of jet penetration into the bubbling bed and bubble diameter.
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Figure 14. Contours of nitrogen temperature from the feeding tube of the modified reactor.
transverse direction as well as the gravitational force (solids flow down to a 45° inclined plane), which makes the bed flow back to main bulk after a bubble has erupted. In this way, the formation of a stationary bubble is avoided and mixing is enhanced. All of these phenomena are clearly visible in Figure 12 where the fluidization quality has been significantly improved. There is no stationary bubble formed next to the feeding point, and the increase in the superficial velocity provides the necessary solids mixing and bubble coalescence for a more evenly distributed bubbling. Figure 13 shows the nitrogen volume fraction at different heights above the feeding tube, which is a clear indication of the bubble diameter. The largest bubble was noted at 1.1 s (intense red dot marker on the plot), with a diameter of approximately 0.050 m, which is a product of bubble coalescence from the feeder and the distributor. In general terms, the bubble sizes are of comparable size with the original case despite the fact that the volumetric flow rate has almost been doubled (Q0 = 70 L/min). This indicates that a lot of gas escapes through the sides of the bed, forming bubbles close to the wall as it is clearly obvious in Figures 12 and 13. The modified configuration becomes more attractive for fast pyrolysis applications. The enhanced mixing of solids provides the necessary convective and conductive heat transfer and temperature uniformity that fast pyrolysis demands. The biomass particles are able to freely move inside the bed and float at the splash zone of the reactor, something that makes the ejection and disengagement of char much more effective. Figure 14 shows the temperature contours of the jet gas from the feeding tube. Compared to the original case, where the gas had to travel a significant distance inside the bed to be heated to the bed temperature, the good solids mixing provided by the modified version overcomes this problem. The gas needs to travel only a few millimeters inside the bed to reach the bed temperature, due to the fact that it is fully exposed to the large heat transfer area of the solids. This also has a great impact to fast pyrolysis, especially if we take into account that it lasts for only 2 s. This fact makes the immediate heating of the gases a necessity.
4.5. Freeboard Gas Flow Dynamics: Modified Reactor. Figure 15 shows the y-velocity contours in the freeboard of the modified reactor. The third modification was the narrowing of the freeboard of the reactor to 0.095 from 0.0999 m. This was done by simply using the mass conservation equation, in order to calculate the necessary gas velocities in the freeboard that will be able to entrain the remained char particles after pyrolysis. The char particle sizes considered were in the range of 250-1000 μm with terminal velocities, Ut, ranging from 0.13 to 1.1 μm, using an average char particle density of 150 kg/m3.17 As it is shown in Figures 15 and 16, the maximum velocities in the freeboard range from 0.5 to 0.6 m/s. These velocities are much less than the terminal velocity of 1000 μm particles; however, someone has to take into account the shrinkage and attrition phenomena that are present in bubbling fluidized beds and have a significant effect in the final size and shape of the particles. The effect of the quality of fluidization on the formation of a separation zone in the freeboard of the reactor is also shown in Figures 15 and 16. The y-velocity is much more evenly distributed across the diameter of the reactor with slight recirculation taking place only when a bubble erupts. Therefore, the effect of the stationary bubble that was present in the original design on the freeboard flow dynamics is by a great percentage eliminated. This makes simple calculations of freeboard velocities come closer to reality. Another important indication of the quality of fluidization and design is that the gas velocity in the freeboard is of the same magnitude and in some cases even smaller than the original design, despite the fact that the volumetric flow rate from the distributor has almost doubled. This can also be seen in Figure 17 where the counter pressure gradients in the separation zones have been greatly reduced, compared to the original design. As it was discussed in section 4.2 for the original case, the freeboard gas dynamics affect the entrainment of char particles. In the modified case, the gas velocity at the freeboard (17) Papadikis, K.; Gu, S.; Fivga, A.; Bridgwater, A. V. Numerical comparison of the drag models of granular flows applied to the fast pyrolysis of biomass. Energy Fuels 2010, 24, 2133–2145.
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Figure 15. Y-velocity contours at the freeboard of the modified reactor.
of the reactor appears to have a smoother profile, with the low velocity region to be present only at the instant of the bubble eruption. This can aid the entrainment of char particles since the gas velocities appear to be more uniform across the diameter of the freeboard. However, even in the case of particle deceleration in the low velocity zone, the improved mixing of solids that the new design offers can more easily direct the char particles to a higher gas velocity region. In this way, the sand bed hydrodynamics can also have a great impact on char entrainment. 4.6. Effect of Turbulence: Modified Reactor. Figure 18 shows the turbulent viscosity ratio in the modified reactor.
The effect of turbulence is greater at the inlet of the feeding tube with a turbulent viscosity ratio μr = 35, while it drops down to μr = 7.5 at the freeboard of the reactor and mainly above the splash zone. These are the two main zones where turbulent dissipation becomes important, while in the rest of the reactor the turbulent contribution becomes insignificant. The simulation of the modified case was not tested without a turbulence model since the turbulent intensity at the inlet of the feeding tube was close to 15% and therefore a laminar flow case would not be realistic. However, the hydrodynamics of the bed would not be expected to be affected by 5647
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Figure 16. Y-velocity components of nitrogen at different heights above the splash zone of the modified reactor.
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Figure 17. Static pressure of nitrogen at different heights above the splash zone of the modified reactor.
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Figure 18. Turbulent viscosity ratio of nitrogen along the height of the modified reactor, starting from the feeding tube.
the exclusion of the turbulence model. The only region that is expected to be affected would be the separation zone above the splash zone of the reactor. A more intense recirculation would be present as a result of the diffusion of kinetic energy being a function only of the molecular viscosity of the gas.
Acknowledgment. The authors gratefully acknowledge the financial support from the UK Engineering and Physical Sciences Research Council (Grant No. EP/G034281/1), EU FP7 SIMUSPRAY project (230715) and the LeverhulmeRoyal Society Africa Award.
Nomenclature 5. Conclusions
Cd = drag coefficient, dimensionless di = diameter, m D = diffusivity, m2/s ess = restitution coefficient, dimensionless g = gravitational acceleration, m/s2 g0,ss = radial distribution coefficient, dimensionless I = stress tensor, dimensionless I2D = second invariant of the deviatoric stress tensor, dimensionless Gk,q = production of turbulent kinetic energy, kg/ms3 hi = specific enthalpy, J/kg kq = turbulent kinetic energy of phase q, m2/s2 kΘs = diffusion coefficient for granular energy, kg/ms Kgs = gas/solid momentum exchange coefficient, dimensionless Lt,q = length scale of turbulent eddies, m m_ pg = mass transfer between the phases, kg/m3s p = pressure, Pa q = heat flux, J/m2s Qpq = intensity of heat exchange between phases, J/m3s Q0 = superficial volumetric flow, m3/s r = radial coordinate, m Re = Reynolds number, dimensionless Sq = energy source, J/m3s t = time, s U0 = superficial gas velocity, m/s Ujet = velocity of secondary gas jet, m/s Ut = particle terminal velocity, m/s vBdr = drift velocity, m/s vi = velocity, m/s VB = phase weighted velocity, m/s
The geometrical optimization of a 1 kg/h bubbling fluidized bed reactor was presented. The simulations showed that a stationary bubble was formed next to the feeding tube of the original reactor, while no fluidization took place underneath the feeder. The stationary bubble was responsible for the formation of a separation zone above the splash zone forcing the gas to flow around it and thus accelerating the gases in the freeboard. The simulations showed that the performance of the reactor can be optimized by simple geometrical alterations. The geometrical modifications performed at the reactor resulted in better quality fluidization eliminating the case of a stationary bubble. By taking advantage of the transverse motion of the sand bed and gravity, the solids were able to flow back to the bulk of the bed and thus enhance the mixing. As a result of this, the separation zone above the splash zone of the reactor was reduced by a great percentage, being only present at the instant of a bubble eruption. The effect of turbulence was insignificant on the bed hydrodynamics, however it had an important contribution on the inlet of the feeding tube and the separation zone above the splash zone. It was shown that the turbulent viscosity ratio can be up to 35 at the feeder and up to 7.5 at the separation zone, contributing significantly to the dissipation of the kinetic energy of the gas. CFD was applied to the geometrical optimization of a bubbling fluidized bed reactor that uses a secondary gas jet through the feeding tube. Several important phenomena were observed and analyzed and possible geometrical modifications were presented. CFD can really aid the design of thermochemical processing equipment by visualizing phenomena that would be otherwise left undetected by simple calculations.
Greek Letters γΘs = collision dissipation of energy, kg/ms3 5650
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ɛq = turbulent dissipation rate of phase q, m /s εi = volume fraction, dimensionless Θi = granular temperature, m2/s2 λi = bulk viscosity, kg/ms μi = shear viscosity, kg/ms μt,q = turbulent viscosity of phase q, kg/ms Fi = density, kg/m3 σlq = dispersion Prandtl number, dimensionless σk = turbulent kinetic energy Prandtl number, dimensionless σɛ = turbulent dissipation rate Prandtl number, dimensionless τ = relaxation time, s τi = stresses tensor, Pa φgs = transfer rate of kinetic energy, kg/ms3 2
3
Subscripts col = collision d = drag dr = drift velocity fr = frictional g = gas i = general index kin = kinetic l = dispersed phase index m = mixture mf = minimum fluidization p = phase index q = phase index s = solids T = stress tensor t = terminal
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