i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 Available online at www.sciencedirect.com ScienceDirect journal homepage: www.elsevier.com/locate/he CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system Hui Liu, Robert J Cattolica*, Reinhard Seiser Department of Mechanical and Aerospace Engineering, University of California, San Diego, 9500 Gilman Drive, La Jolla, CA, 92093, USA article info abstract Article history: A comprehensive CFD (Computational Fluid Dynamics) model using the MP-PIC (Multi- Received 21 December 2015 phase Particle-In-Cell) method was developed to simulate a pilot-scale (6 tons/day, MW Received in revised form th) dual fluidized-bed biomass gasification system In this model the particulate phase was April 2016 described with the blended acceleration model The momentum, mass, and energy Accepted 28 April 2016 transport equations were integrated with the kinetics of heterogeneous biomass and char Available online xxx reactions and homogeneous gas-phase reactions to predict the particle circulation, producer gas composition, and reactor temperature The simulation results were compared Keywords: with experimental data from the pilot-scale gasification system to validate the model at Biomass different operating conditions Parametric studies were conducted to investigate the Gasification impact of gasifier temperature, steam to biomass ratio (S/B), and air supply to the Fluidization combustor on the producer gas composition The studies showed that increasing gasifier Circulation rate temperature and steam to biomass ratio (S/B) promoted syngas (CO þ H2) production and CFD increased hydrogen content in producer gas The effect of air supply was minor, because Pilot-scale for the dual fluidized-bed system air was not directly involved in biomass gasification © 2016 Hydrogen Energy Publications LLC Published by Elsevier Ltd All rights reserved Introduction Currently, energy and chemical industries rely primarily on fossil fuels Hydrogen as a clean energy source with the high energy density can become an alternative to fossil fuels [1,2]; however, the current hydrogen production also depends on fossil fuels and most of hydrogen is produced from natural gas reforming and coal gasification [3,4] Biomass as a renewable energy source can be used to produce hydrogen through biomass gasification [5,6] Two types of technologies such as fixed-bed and fluidized-bed are mainly used for biomass gasification Fixed-bed biomass gasifiers are mostly preferable for small-scale syngas production with regard to the simple process and low capital investment [7]; however, due to the insufficient gas-particle contact, biomass gasification process in fixed bed reactors is slow and the tar content in the producer gas is relatively high [8e10] Therefore, fixed-bed gasifiers are not suitable for largescale syngas production and are only preferable for small size plants with the capacity of up to 1.5 MW th; comparatively, the capacities of atmospheric bubbling fluidized-bed gasifiers can be up to 25 MW th [11] In addition, fluidized-bed gasifiers demonstrate good tolerance to particle sizes The gasesolid mixing is more efficient and less tar is generated in fluidizedbed gasifiers [12,13] In a conventional single-reactor fluidized-bed gasifier, air and biomass are fed to the same reactor (gasifier) which has * Corresponding author Tel.: þ1 858 5342984 E-mail address: rjcat@ucsd.edu (R.J Cattolica) http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 0360-3199/© 2016 Hydrogen Energy Publications LLC Published by Elsevier Ltd All rights reserved Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 some disadvantages After the oxygen gas in air is consumed, the remaining nitrogen gas in air mixes with producer gas and dramatically dilutes the producer gas concentration Consequently, the heating value of producer gas generated in such a single-reactor fluidized-bed systems is low [14,15] This issue can be eliminated by the use of a dual fluidizedbed gasification system As shown in Fig 1, a dual fluidizedbed system generally consists of two reactors: a fluidizedbed gasifier and a fluidized-bed combustor In the process biomass and steam are fed to the gasifier while air is only supplied to the combustor Char is generated from biomass pyrolysis in the gasifier Char is entrained by the bed material and is delivered from the gasifier to the combustor Char reacts with air in the combustor to release heat The combustion heat is absorbed by the bed material and is returned to the gasifier through the bed material circulation within the dual fluidized-bed system Since air is only present in the combustor and the combustor is separated from the gasifier, the producer gas in the gasifier is free of nitrogen Accordingly, the producer gas in dual fluidized-bed systems will have a higher heating value than the producer gas generated in a single-reactor fluidized-bed system The dual fluidized-bed system as depicted in Fig contains a gasifier and a combustor These two reactors can be either a bubbling fluidized-bed (BFB) or a pneumatic transport bed/ riser reactor There are mainly four configurations of dual fluidized-bed gasifiers: a BFB gasifier and a BFB combustor, a BFB gasifier and a riser-combustor, a riser-gasifier and a BFB combustor, and a riser-gasifier and a riser-combustor [16] A dual fluidized-bed system with a riser-gasifier and a risercombustor was used to gasify waste and biomass in 1970s A dual fluidized-bed system with a BFB gasifier and a risercombustor was developed in 1990s and a biomass gasification plant using this technology was built and has operated in Austria since 2002 [17,18] A dual fluidized-bed system with a riser-gasifier and a BFB combustor was adopted by Ebara Co., Ltd in 2003 [19] Among these configurations, the combination of BFB gasifier and riser combustor was considered to be optimal for large-scale biomass gasification with regard to efficient particle circulation, high fuel conversion, and low tar generation [20] Fig e Schematic of dual fluidized-bed gasifier The design and scale-up of dual fluidized-bed gasifiers are challenging and have in the past depended on empirical scaling formulas, especially for dual fluidized-bed gasifiers The interaction between reactors and cyclone separators requires sophisticated analysis [15,21] In recent years, CFD (computational fluid dynamics) has proved to be a powerful tool for the simulation of gas-particle system and numerous CFD models were developed to simulate fluidized-bed reactors [22e25] Currently, three main methods have been applied for the CFD modeling of fluidized-bed gasifiers: the EulerianeEulerian (EE) approach, the EulerianeLagrangian (EL) approach, and the hybrid EulerianeLagrangian approach Compared with the EE and EL approaches, the Multiphase Particle-In-Cell (MP-PIC) method as a hybrid EulerianeLagrangian approach can provide both the required accuracy and efficiency The MP-PIC method was initially developed by Harlow et al [26] for single-phase flows and then was improved significantly by O'Rourke et al [27] for multiphase flows In the MP-PIC method an isotropic stress term is applied in the particle acceleration equation to calculate particle interactions Since this solid stress term is defined by a function of solid volume fraction, the trajectory of each particle is not needed, which saves the computation time significantly Therefore, the MP-PIC method is a computational-efficient method and can also be applied to simulate dense phase flows [28,29] For the EE approach, particle sizes in a particulate phase must be set to the same value In contrast, particles in the MPPIC method can have different diameters by the use of particle size distribution function (PSD) Additionally, the calculations of momentum, mass, and energy transfer for the particulate phase in the MP-PIC method are implemented on individual particles or numerical particle parcels Thus, if there are solidegas reactions occurring on particles, each particle size can change in accordance with solid species generation or consumption by the reactions Shi et al [30] established a hydrodynamic model using the MP-PIC method to investigate the effect of particle size Two size distribution functions, the Gaussian and Lognormal size distributions, were applied to simulate the particulate phase in a circulating fluidized-bed (CFB) riser The study showed that the PSDs had significant impact on the flow pattern in the lower region of the riser Wang et al [31] developed another hydrodynamic MP-PIC model to study a binary PSD case and a polydisperse case The predicted flow pattern and particle velocity in both of cases showed good agreement with experimental data In comparison with the EL approach that can be only applied to simulate small-scale fluidized-bed systems, the MP-PIC method is capable of simulating the full-loop of largescale fluidized-bed gasifiers Wang et al [32] built a hydrodynamic MP-PIC model to simulate the full-loop of a CFB system including a riser, a cyclone separator, and a loop-seal Their model successfully predicted the particle circulation and the pressure distribution in the full-loop of CFB system Jiang et al [33] also conducted a hydrodynamic study to simulate the fullloop of a CFB system including a riser, six cyclone separators, and six loop-seals The predicted solid circulation, pressure, and velocity profiles were validated with experimental data As shown above, the MP-PIC models are capable of presenting Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 more practical and valuable information for the design and scale up of fluidized-bed gasifiers than the EL approach Previous models were primarily developed for hydrodynamic studies without considering chemical reactions and were mainly focused on single-reactor fluidized-bed systems A complete CFD model of biomass gasification in a dual fluidized-bed system including reaction kinetics is rarely seen Additionally, previous analyses of the single-reactor systems can't be applied directly to the dual fluidized-bed systems, considering the significant difference in the configurations of two fluidized-bed systems It is necessary to develop a CFD model of dual fluidized-bed gasifier to facilitate the design and scale-up of such gasifiers In this work, a three-dimensional CFD model using the MP-PIC method is built to simulate a pilot-scale dual fluidized-bed biomass gasifier with the capacity of tons/day and MW th This simulated dual fluidized-bed system includes a BFB gasifier, a riser-combustor, a cyclone separator, and a loop-seal In this model, the gas phase is described by the Large Eddy Simulation (LES) while the particulate phase is described by the blended particle acceleration equation The momentum, mass, and energy transport equations are integrated with the kinetics of gasesolid and gasegas reactions to simulate biomass gasification in the dual fluidized-bed system The simulation results such as producer gas composition and reactor temperature are compared with experimental data to validate the model at different operating conditions The effects of important operating parameters such as gasifier temperature, steam to biomass ratio, and air supply to the combustor are also analyzed in this work ffiffiffiffi p V D¼ (7) The species transport equations is shown as follows:   v ag rg Yn vt  a m   g þ V$ ag rg ug Yn ¼ V$ VYn þ dmn; react: Sc The energy transport equation is applied to solve for the temperature of the gas phase, as shown below:   v ag rg E vt keddy ¼   À À vp þ V$ ag rg ug E ¼ ag þ ag ug $Vp À V ag kmol vt Á Á þ keddy VTg þ Sinter þ qdiff þ Q Cp meddy Prt The dual fluidized-bed gasification system contains two types of particles, biomass and bed material The solid movement and solid mixing in the binary-particle system are described by the blended particle acceleration equation The equation is shown as follows [35]: À Á Vp dup up À up ¼ Dp ug À up À þXþgþ rp dt 2tD The continuity equation and the momentum transport equation for the gas phase are as follows: h¼   v ag rg ug vt   þ V$ ag rg ug ug ¼ ÀVp þ F þ ag rg g þ V$tg (10) Particulate phase 16 as s ¼ pffiffiffiffiffiffi g0 ðas Þhð1 À hÞ tD 3p r32   þ V$ ag rg ug ¼ dmp (9) where Sinter is the energy exchange between the gas and particulate phases, qdiff is the enthalpy diffusion, and Q is the energy source by chemical reactions Gas phase vt (8) (11) where X is the modified acceleration due to the contact force between particles tD is the damping time due to inelastic particle collisions and is defined as follows: Governing equations   v ag rg (1) (2) þ ep s2 ¼ À Á2 ∭ fm up À up dmp dup dTp rp as as ¼ ∭ f (12) (13) (14) m dmp dup dTp rp (15) where dmp is the mass production from the gasesolid reactions, tg is the stress tensor of the gas phase and is calculated using the following equations [34]: rp ¼ ∭ fmup dmp dup dTp as (16)  vu  vu  vug;j 2 g;i k m þ meddy dij þ À tg ¼ mlam þ meddy lam vxj vxi vxk (3) up ¼ ∭ fmup dmp dup dTp rp as (17)   meddy ¼ Crg D2 S (4) g0 ðas Þ ¼   qffiffiffiffiffiffiffiffiffiffiffiffiffiffi S ¼ 2Sij Sij (5) r32 ¼ (6) where s is the mass-weighted particle velocity variance, r32 is the Sauter mean radius, g0(as) is the radial distribution Sij ¼   vui vuj þ vxj vxi as;cp as;cp À as (18) ∭ fr3 dmp dup dTp ∭ fr2 dmp dup dTp (19) Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 function, and ep is the particleeparticle restitution coefficient.X is defined in the following equations: Reaction kinetics " À Á À Á vtp ~p À D ug À u ~p À X¼À þ g1 ðas Þ D ug À u rp as vxi In this dual fluidized-bed system, after biomass is fed to the gasifier, moisture is released from biomass and then biomass is decomposed into char, volatile gases and ash Some char remains in the gasifier to react with gases while the rest of char is transported to the combustor to react with O2 to release heat In this model the heterogeneous reactions such as biomass drying, pyrolysis, and char gasification and combustion are included The homogeneous gas-phase reactions including the water gas shift reaction, steam reforming reaction, and gas oxidation reactions are also included The heterogeneous reactions are defined as discrete particle reactions In this model the calculations of heterogeneous reactions are implemented on each numerical particle Each numerical particle can have its own size, temperature, and solid species Note that in this model all of heterogeneous and homogeneous reactions are described with global reaction schemes, and detailed reaction mechanisms are not used It is well-known that detailed reaction kinetics can provide much more accurate predictions than global reaction kinetics; however, detailed reaction kinetics require intense computation For example, as reported by Titova et al [37], the detailed reaction kinetics of propane combustion required 599 reaction steps and 92 species For a gasification model involving both heterogeneous and homogeneous reactions, thousands of elementary reactions may be required Considering the unaffordable computation cost of detailed reaction schemes, global reaction schemes were adopted for the CFD modeling of biomass gasification [23,38e41] 1 À rp rp ! vp vxi # (20) 10Ps abp ÂÀ Á À Áà max acp À ap ; ε À ap ¼ & g1 ðas Þ ¼ D¼ (21) if as ¼ if as ¼ as;cp (22) ∭ fmDdmp dup dTp rp as (23) ∭ fmDup dmp dup dTp rp as D (24) ~p ¼ u where is the isotropic solid contact stress, Ps, b, and ε are the model constants, g1(as) is the blending function, D is the ~p is the drag-averaged average particle drag coefficient, and u particle velocity The interphase momentum transfer, F, is defined by: # ) ( " À Á Vp dmp þ up dmp dup dTp F ¼ ∭ f mp Dp ug À up À rp dt (25) The drag coefficient, Dp, is described as follows [36]: rgjug Àup j Dp ¼ Cd rp dp (26) Biomass drying > > > > > > < 24agÀ2:65 Re The biomass drying rate is described in the following Arrhenius equation [42]: ; Re < 0:5 À Á Cd ¼ 24aÀ2:65 g > þ 0:15Re0:687 ; 0:5 Re > > Re > > > : ; Re > 1000 0:44aÀ2:65 g 1000 (27) dmp dmp dup dTp dt (28) (29) dmp;n ag Mwp;n dCp;n ¼ mp dt rp ap dt (30) Á dTp kd Nu À ¼ Ap Tg À Tp mp dp dt where Cp,n is the concentration of solid species n Biomass pyrolysis The one-step global-reaction scheme is used to simulate biomass pyrolysis in which biomass is decomposed into volatile gases and char In the experiments tar content was less than 2% of mass fraction due to the high gasifier temperature of 850 C In this model tar is not included and is assumed to be fully converted to non-condensable gases for simplicity Biomass pyrolysis is modeled as follows: Biomass/a1 CO þ a2 CO2 þ a3 H2 þ a4 CH4 þ a5 C2 H4 þ a6 C2 H6 N dmp X dmp;n ¼ dt dt i¼1 CV R (1) where mbio is the mass of biomass Since the MP-PIC method is a Lagrangian-based method, mass transfer and energy transfer in the model are calculated on the basis of numerical particles Note that in the MP-PIC method a numerical particle represents a cluster of real particles to simplify the computation The mass and energy transport equations for the particulate phase are: dmp ¼ À∭ f   À10585 mbio r1 ¼ 5:13  1010 exp Tp þ a7 Char R (2) (31) Various methods can be applied to determine the values of stoichiometric coefficients One method is to assign the coefficient values of pyrolysis products based on pyrolysis experimental results [43,44] The advantage of this method is that all of the coefficient values are directly from experimental measurements However, in such a pyrolysis experiment the heating rate is generally much slower than the heating rate of Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 pyrolysis in a gasifier where both of biomass pyrolysis and combustion occur Since biomass pyrolysis is dramatically influenced by the heating rate, the product composition from slow pyrolysis may not be the same as the product composition of fast pyrolysis in a gasifier [45e47] Additionally, in Reaction-2, on the left (reactant) side the contents of C, H, and O in biomass were determined in the ultimate analysis while on the right (product) side the contents of these elements were measured in the pyrolysis experiment Since these contents were measured in two different experiment using different methods, the contents of C, H, and O in the reactant may not be the same as those in the products in Reaction-2 Therefore, the elemental balance may not be strictly followed in this pyrolysis modeling approach In this model another method is used to calculate the coefficients to achieve better elemental mass balances The values of a1 a2, a3, and a7 for the major species, CO, CO2, H2, and Char, are calculated from the proximate and ultimate analysis data, as proposed by other researchers [48e50] The values of a4, a5, and a6 for the minor species such as CH4, C2H4, and C2H6 are adjusted to fit experimental data C19.8166H24.524O11.8501 is defined as biomass, based on the ratios of C/H/O in the ultimate analysis of biomass sample The coefficient values are shown in Table It is seen that the elemental mass balances are applied The contents of C, H, and O in the products of biomass pyrolysis agree well with the elements measured in the ultimate analysis of biomass Note that due to the strict elemental mass balances, changes in any defined species of pyrolysis can cause changes in the coefficient values of other species For example, if only CH4 and C2H4 are defined as minor species, rather than three species of CH4, C2H4, and C2H6 as in the current model, the coefficient values of other species need to change to maintain the elemental mass balances [51] In the dual fluidized-bed system, some char is transported from the gasifier to the combustor and reacts with air in the combustor to release combustion heat Char combustion is defined as follows [52]: R (4) C þ H2 O4CO þ H2 R (5) C þ 2H2 4CH4 R (6) The rates of the reactions are calculated in the following equations:   À22645 ½CO2 Š r4f ¼ 1:272mc Texp Tp (33)   À2363 r4r ¼ 1:044  10À4 mc T2 exp À 20:92 ½COŠ2 Tp (34)   À22645 ½H2 OŠ r5f ¼ 1:272mc Texp Tp (35)   À6319 À 17:29 ½H2 Š½COŠ r5r ¼ 1:044  10À4 mc T2 exp Tp (36)   À8078 r6f ¼ 1:368  10À3 mc Texp À 7:087 ½H2 Š Tp (37)   À13578 À 0:372 ½CH4 Š0:5 r6r ¼ 0:151mc T0:5 exp Tp (38) The reaction kinetics were originally proposed by Syamlal and Bissett [53] and then were adjusted by Snider et al [54] to suit the purpose of particle-chemistry modeling in the MP-PIC method Water gas shift reaction, steam reforming reaction, and oxidation reactions of CO, H2, CH4, C2H4, C2H6, and C3H8 are included in this model The kinetics of homogeneous reactions are shown in Table Experiment setup and model settings (32) The pre-exponential factor of 8:68  106 is adjusted to fit experimental data Table e Pyrolysis product coefficients CO CO2 H2 CH4 C2H4 C2H6 Char C þ CO2 42CO R (3)   À29160 ½O2 Š r3 ¼ 8:68  106 ac Texp Tp Species The reactions between char, CO2, H2O, and H2 are considered to be reversible reactions and are described in the forward and reverse reactions: Homogeneous gas-phase reactions Char combustion and gasification C þ O2 /CO2 Coefficient Value a1 a2 a3 a4 a5 a6 a7 5.50765 3.1712 5.7076 2.1558 0.49665 0.4165 7.1556 The data used to validate the CFD model are from the operation of a pilot-scale dual fluidized-bed gasifier with the capacity of tons/day and MW th at Woodland Biomass Research Center (WBRC), located in Woodland, California, USA, as shown in Fig (a) The dimensions of the dual fluidized-bed system are demonstrated in Fig (b) Almond prunings were used as biomass feedstock in the experiments and the properties of almond prunings are displayed in Table As seen in Fig (c), this dual fluidized-bed system included a gasifier, a combustor, a cyclone separator, and a loop-seal In the experiments biomass was discharged constantly from a storage cart on a weight scale Biomass was transported by a bucket elevator to a couple of biomass hoppers and then was distributed continuously to the gasifier with a screw conveyor Steam was generated in a steam generator and was superheated over 330 C in a series of heat Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 Table e Homogeneous reaction kinetics [55e59] Homogeneous reactions Water gas shift reaction CO þ H2 O/CO2 þ H2 (R7) Steam reforming reaction CH4 þ H2 O/CO þ 3H2 (R8) CO oxidation CO þ 0:5O2 /CO2 (R9) H2 oxidation H2 þ 0:5O2 /H2 O (R10) CH4 oxidation CH4 þ 2O2 /CO2 þ 2H2 O (R11) C2H4 oxidation C2 H4 þ 3O2 /2CO2 þ 2H2 O (R12) C2H6 oxidation C2 H6 þ 3:5O2 /2CO2 þ 3H2 O (R13) C3H8 oxidation C3 H8 þ 5O2 /3CO2 þ 4H2 O (R14)  r7 ¼ 2:75as exp  À10079 Tg ½COŠ½H2 OŠ the pre-exponential factor, 2.75,   is adjusted to fit experimental data  ½CH4 Šþ½H2 OŠ r8 ¼ 720as exp À9057 2½H2 Šþ½COŠ Tg the pre-exponential  720, is adjusted to fit experimental data  factor, ½COŠ½O2 Š0:25 ½H2 OŠ0:5 r9 ¼ 1:28  1017 exp À34761 Tg   r10 ¼ 1:0  1014 exp À5052 ½H2 Š½O2 Š Tg   ½CH4 Š0:7 ½O2 Š0:8 r11 ¼ 5:01  1011 exp À24417 Tg   r12 ¼ 1:0  1015 exp À20808 ½C2 H4 Š½O2 Š Tg   ½C2 H6 Š0:5 ½O2 Š1:25 r13 ¼ 4:4  1011 exp À15199 Tg   r14 ¼ 8:6  1011 exp À15000 ½C3 H8 Š0:1 ½O2 Š1:65 Tg Fig e (a): Dual fluidized-bed gasifier at WBRC (b): Dimensions of dual fluidized-bed (c): Flow chart of biomass gasification system (d): Boundary conditions (BC) of CFD model exchangers Then, the superheated steam was injected at the bottom of the gasifier through six nozzles Some of char generated from biomass pyrolysis was transported from the gasifier to the combustor The air was preheated above 290 C in a series of heat exchangers and was supplied to the combustor at three different locations as the 1st, 2nd, and 3rd air supplies Char in the combustor burned with air to release the combustion heat Propane and an additional amount of air Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 Table e Proximate and ultimate analysis of biomass feedstock Proximate analysis (wt %) Moisture 5.18 Fixed carbon 20.20 Volatile matter 72.53 Ash 2.09 Ultimate analysis (wt %) C 51.3 H 5.29 O 40.9 N 0.66 S 0.01 Cl 0.04 Residual 1.80 The simulation was solved with the finite volume method As shown in Equations (4)e(7), the subgrid-scale turbulence model of large eddy simulation was applied to describe turbulent gas flows in the dual fluidized-bed system [61] The partial donor cell differencing scheme, a weighted average of central difference and upwind scheme, was applied to calculate the face value of the variables in the transport equations This scheme is shown as follows [62]: 4¼ were fed to the combustor through the startup burner to provide the additional heat to the system The bed material absorbed most of the combustion heat and then was transported from the combustor to the cyclone separator The bed material was disengaged from the flue gas in the separator and then fell down to the loop-seal After fluidized by the steam in the loop-seal, the bed material was carried back to the gasifier In the experiments a commercial ceramic bed material (CARBO HSP® 30/60) was used in the experiments Producer gas was sampled by stainless sampling lines from the gasifier Then, the gas was cleaned in an impinger filled with biodiesel at C and was analyzed with an Agilent 2000 Micro GC Tar was sampled using the tar protocol according to EU-CEN/TS 15439 The gravimetric tar was between and 17 g/ Nm3, and benzene was 1500 ppm These compounds had a mass fraction of less than 2% and they were not included in the model for simplicity The error of producer gas analysis was mainly from GC calibration process, GC measurement, and gas sampling In this study the overall experimental uncertainties were estimated at 10% for the major contents including H2, CO, CO2, and CH4 and 15% for the minor contents such as C2H4 and C2H6, as reported by Billaud et al [60] A 3D CFD model was built in the CFD software, Barracuda Virtual Reactor® The full-loop of the dual fluidized-bed system including a gasifier, a combustor, a cyclone separator, and a loop-seal was simulated in this model As displayed in Fig (d), a particle injection boundary condition (BC) with 25 injection points was defined as the biomass inlet A fluid injection BC with 48 injection points was applied to simulate steam nozzles in the gasifier Note that the arrow direction of injection BC indicated the flow direction In the model the steam flow was set in the downward direction, which simulated the effect of the cap of the steam nozzle In the experiments each of steam nozzles was equipped with a cap When steam was introduced to the gasifier through the nozzles, steam initially flew upwards and then was diverted and flew downwards after encountering the caps A fluid injection BC with 36 injection points was used to simulate the nozzles of the 1st air supply in the combustor A fluid injection BC with injection points was used to simulate air feeding pipes as the 2nd air supply Another injection BC with injection points was applied to simulate air feeding pipes as the 3rd air supply Two fluid injection BCs with injection points each were defined for two steam feeding pipes in the loop-seal The propane supply to the combustor was defined by a mass flow rate BC The outlets of the gasifier and cyclone separator were defined by two pressure outlet BCs u1 a1 A1 r1 þ u2 a2 A2 r2 (39) the donor cell property: & Qd ¼ Q1 if > Q2 if < (40) the acceptor cell property: & Qd ¼ Q1 if > Q2 if < (41) the face property: 1 Q 12 ¼ Q d ð1 þ JÞ þ Qa ð1 À JÞ 2 (42) J ¼ a þ bC (43) C¼ 2Dtj4j a1 V1 þ a2 V2 (44) where a and b are model constants and are defined as 0.2 and 1.0, respectively The no-slip boundary condition was applied at walls for the gas phase while the partial-slip boundary condition was implemented for the particulate phase The time-step size was between  10À3 and  10À5 s and was automatically controlled by the CFL value (Courant-Friedrichs-Lewy Scheme) and the maximum temperature change in a cell: if the CFL value is lower than 0.8, the time-step size is increased; when the CFL value is higher than 1.5, the timestep size is then decreased Additionally, whenever the temperature change in a cell at a time-step exceeds 300 K, the time-step size also decreases Orthogonal structured grids were generated in Barracuda Virtual Reactor® for this CFD model Three computational grids with 216,972, 243,423, and 348,768 cells were compared and the grid of 243,423 cells was finally selected due to its acceptable accuracy and affordable computational cost [51] The residuals of the equations of volume fraction, pressure, velocity, energy were set to 10À7, 10À8, 10À7, and 10À6 as the simulation convergence criterion The homogeneous gas-phase reactions were calculated with the cell-volume averaged chemistry method and the heterogeneous gasesolid reactions were modeled by the discrete particle chemistry method The simulation was performed with the accelerated GPU (graphics processing unit) computing in a computer workstation with a GTX TITAN black graphics card The simulation time for each run was set as 100 s and the final solutions were averaged from 80 to 100 s, which took about days to be completed The detailed model settings are shown in Table Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 Table e Model settings and operating conditions Biomass density (kg/m3) Biomass mean diameter (m) Bed material density (kg/m3) Mean diameter of bed material particles (mm) Particle size distribution of bed material Standard deviation of the normal distribution Solid volume fraction at close pack Initial bed height (m) Pressure at the gasifier and cyclone outlets (atm, abs.) Case settings The biomass feed rate (kg/h) The steam supply to the gasifier (kg/h) The 1st air supply to the combustor (kg/h) The 2nd air supply to the combustor (kg/h) The 3rd air supply to the combustor (kg/h) The additional air supply to the combustor (kg/h) The propane supply to the combustor (kg/h) The steam supply to the cyclone separator (kg/h) Results and discussion Particle flow pattern and gas distributions in the dual fluidized-bed system The particle circulation in the system between and 60 s is presented in Fig In the figure, particles are fully fluidized 550 0.0057 3560 488 Normal distribution 0.146 dp 0.56 2.50 1.0 Case 228.04 85.32 28.89 240.68 305.63 561 14.63 78.21 Case 219.55 78.21 40.85 283.37 312.62 561 21.92 85.32 Case 243.04 64.71 26.52 238.01 322.76 561 64.71 and are entrained by the gases from the combustor to the cyclone separator The particles fall down to the loop-seal and are fluidized by steam Finally, particles are delivered back to the gasifier It is also seen that the particles accumulate in the loop-seal at s The accumulation of particles continues to grow in the loop-seal and even reaches the bottom of the cyclone separator at 20 s After that, the accumulated particles are Fig e Particle circulation and solid build-up in the dual fluidized-bed system (Case 1) Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 gradually removed from the loop-seal between 30 and 60 s and are transported to the gasifier The predicted accumulation of particles in the loop-seal matches the observation during the plant startup It is mainly caused by the high rate of the initial solid mass flow from the combustor to the cyclone separator and loop-seal during startup Initially, most of the particles in the combustor are quickly carried outside and delivered to the cyclone separator and loop-seal from to s due to the high gas velocity in the combustor The rate of the initial solid mass flow is so high that the particles entrained to the loop-seal can't be removed quickly and accumulate from to 20 s Meanwhile, particles are continuously transported from the gasifier to the combustor through the connection pipe; due to the small size of the connection pipe, the solid flow rate from the gasifier to the combustor is much smaller than the initial solid flow rate from the combustor to the cyclone separator As a result of the slow incoming solid flow to the combustor, the solid mass flow rate from the combustor to the cyclone separator is then decreased Accordingly, the particle accumulation in the loop-seal stops growing and the accumulated particles are gradually removed from the loop-seal and are delivered to the gasifier The sectional views of H2, CO, CO2, CH4, H2O, and O2 on the xz (vertical) and xy (horizontal) planes in Case are shown in Fig H2, CO, and CH4 are generated in the lower left region of the gasifier The gases are accumulated in the bed of the gasifier and then penetrate through the bed to reach the freeboard region, and finally leave the gasifier at the top During the process, none of the volatile gases such as H2, CO, and CH4 leak to the combustor CO2 as a product of biomass pyrolysis and char combustion is found in both the gasifier and combustor Meanwhile, steam is injected to the bottom region of the gasifier through nozzles Most of steam rises to the freeboard region to react with other gases after permeating through the right side of the bed A small amount of steam escapes from the gasifier to the combustor through the connection pipe In comparison, O2 only appears in the combustor to react with char and propane, and there is no O2 escaping to the gasifier Since no air is present in the gasifier, the producer gas in the dual fluidized-bed system is free of N2 and can have a high heating value The low content of N2 in the producer gas can also be beneficial for the downstream processing by saving the cost of nitrogen gas removal in the process of syngas purification The sectional views on the xy plane at the heights of 1.6, 2.0, 2.6, and 4.0 m show that the gases are unevenly distributed in the bed of the gasifier The volatile gases such as H2, CO, CO2, and CH4 are accumulated in the left region of the gasifier where biomass enters; steam mainly stays in the right region of the gasifier The non-uniform patterns of the gas distributions may be caused by the feeding locations of biomass and steam in the gasifier Biomass is fed at the left side of the gasifier while steam is injected at the bottom through the nozzles The asymmetrical shape of the gasifier makes the steam injection location closer to the right side of the gasifier The advantage of the right-side steam injection is that steam can act as a sealing gas to avoid the volatile gases escaping from the gasifier to the combustor and prevent O2 leaking from the combustor to the gasifier For a single-reactor fluidized-bed systems, part of the producer gas is burned so Fig e H2, CH4, CO2, CO, H2O, and O2 distributions in the dual fluidized-bed system (Case 1) The horizontal section views show the distributions of H2, CH4, CO2, CO, and H2O (at the heights of 1.6, 2.0, 2.6, and 4.0 m), and O2 (at the heights of 0.7, 2.2, 3.0, and 5.0 m) Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 10 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 not all of the producer gas is available to be delivered to the downstream unit for further processing which is the case for the dual fluidized-bed system Simulation results The simulation time for each test was set as 100 s and the predicted simulation results were averaged between 80 and 100 s Fig demonstrates the producer gas composition at the outlet of the gasifier in 100 s for Case It is seen that the concentrations of H2, CO, CO2, CH4, C2H4, and C2H6 increase rapidly between and 40 s After that, the process reaches the steady-state and all of gas concentrations become constant The similar trends are also observed in Cases and In Fig (aec), Fig (aec), and Fig (aec) , the predicted gas compositions and reactor temperatures for Cases 1e3 are compared with three sets of experimental data Gasifier temperatures in the bottom region, the bed, and the upper region were chosen for the gasifier comparison at the heights of 0.66, 1.12, and 3.05 m Three combustor temperatures as the bottom, lower, and upper temperatures at the heights of 0.55, 1.83, and 6.4 m were selected for the combustor comparison The differences between the predicted and measured concentrations of major contents such as H2, CO, CO2, and CH4 are 1.40%, 3.65%, 11.80%, and 2.70% in Case 1, 2.78%, 1.58%, 8.97%, and 1.83% in Case 2, and 10.63%, 3.86%, 16.67%, and 14.35% in Case The differences in temperature prediction at locations of the gasifier and combustor are 2.11%, 3.24%, 1.41%, 1.19%, 0.37%, and 2.86% in Case 1, 6.04%, 6.53%, 2.57%, 1.64%, 2.56%, and 4.87% in Case 2, and 5.35%, 1.54%, 2.23%,1.49%, 0.12%, and 3.20% in Case As shown above, the predicted gas compositions and the gasifier and combustor temperatures are close to the measured values in the experiments Large discrepancies are observed in the predictions of C2H4, and C2H6 It is mainly because in the model only two species, C2H4 and C2H6, are considered as minor contents, and other content such as tar (less than 2% mass fraction) is not included for simplicity Fig e (a): Gas composition comparison (Case 1) (b): Gasifier temperature comparison (Case 1) (c): Combustor temperature comparison (Case 1) Effect of gasifier temperature To understand the effect of the gasifier temperature, Case as a base case and three additional cases were investigated The propane flow rates were set as 0, 21.9, and 43.2 kg/s for the 0.4 0.35 Mole Fraction 0.3 H2 CO CO2 CH4 C2H4 C2H6 0.25 0.2 0.15 0.1 0.05 0 20 40 60 Time (s) 80 100 Fig e Producer gas composition in 100 s (Case 1) additional cases to generate different gasifier temperatures In this study the gasifier temperature is represented by the “gasifier bed” temperature as described in the previous section The influence of gasifier temperature on the producer gas composition is demonstrated in Fig (a) The concentrations of H2 and CO increase and the concentrations of CO2 and CH4 decrease as the gasifier temperature increases The predicted trends are consistent with the experimental data The increase in temperature generally promotes the endothermic reactions such as char and carbon dioxide reaction (R4), char and steam reaction (R5), and steam reforming reaction (R8) [63,64] Therefore, when the gasifier temperature increases, the concentrations of H2 and CO as products of Reactions-4, and 8, increase and the concentrations of CO2 and CH4 as reactants of Reactions-4 and 8, decrease Compared with the single-reactor fluidized-bed systems, the gas composition in this dual fluidized-bed system varies in a relatively narrow range with gasifier temperature [65,66], Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 Fig e (a): Gas composition comparison (Case 2) (b): Gasifier temperature comparison (Case 2) (c): Combustor temperature comparison (Case 2) indicating that the effect of gasifier temperature on the producer gas composition in the dual fluidized-bed gasification system is not as significant as those of the single-reactor fluidized-bed systems In biomass gasification producer gas is mainly generated from biomass pyrolysis, charegas reactions, and other gas-phase reactions As described previously, in the current dual fluidized-bed system char is primarily transported from the gasifier to the combustor and only a portion of char remains in the gasifier Due to the reduced amount of char in the gasifier, the gas production and consumption from the charegas reactions including Reaction-4 and are significantly reduced Therefore, the impact of the gasifier temperature in this dual fluidized-bed system is less significant than that in the single-reactor fluidized-bed systems As indicated previously, the main purpose of biomass gasification is to produce hydrogen as an alternative to fossil fuels Therefore, the effect of gasifier temperature on CO and H2, two important contents for hydrogen production, is also investigated in this section Fig (bec) show the impact of gasifier temperature on the net heating value (NHV) of syngas (CO þ H2) and the molar ratio of H2/CO It is observed that the NHV of syngas (CO þ H2) 11 Fig e (a): Gas composition comparison (Case 3) (b): Gasifier temperature comparison (Case 3) (c): Combustor temperature comparison (Case 3) increases from 6570 to 7045 kJ/kg (biomass) as the gasifier temperature increases from 802 to 929 C; meanwhile, the molar ratio of H2/CO also increases from 1.13 to 1.20 As shown Fig.4e9 (dee), 4.04%, 6.92%, and 15.74% of increases in gasifier temperature result in 3.10%, 3.87%, and 7.22% of increases in the NHV of syngas (CO þ H2) The increases in gasifier temperature also achieve 1.96%, 3.36%, and 6.28% of increases in H2/CO It indicates that high gasifier temperature promotes syngas production and increases hydrogen concentration in producer gas Effect of steam to biomass ratio In this section Case and three more cases are compared to examine the effect of steam to biomass ratio (S/B) on the producer gas composition The predicted gas compositions from these four cases are presented in Fig 10 (a) It is observed that when the steam to biomass ratio increases, H2 concentration increases and CH4 concentration decreases The effects of steam to biomass ratio on the NHV of syngas (CO þ H2) and H2/CO are presented in Fig 10 (bec) Similar to Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 12 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 (d) Fig e (a): Effect of gasifier temperature on producer gas composition (b): Effect of gasifier temperature on NHV of CO þ H2 (c): Effect of gasifier temperature on H2/CO (d): NHV increase of CO þ H2 (e): Increase of H2/CO the effect of gasifier temperature, high ratio of steam to biomass promotes syngas (CO þ H2) production and increases hydrogen content in producer gas, which was also observed by Pinto et al [65], Franco et al [66], and Gungor et al [67] In Fig 10 (dee), it is seen that 19.40%, 79.10%, and 138.80% of increases in steam to biomass ratio result in 1.33%, 3.49%, and 6.36% of increases in the NHV of syngas and 0.94%, 3.48%, and 5.08% of increases in H2/CO On the other hand, compared with the impact of gasifier temperature, the effect of steam to biomass ratio on syngas production and hydrogen content is much smaller The similar trend was also observed by other researchers [68e71] Generally, increasing steam can promote the reactions such as char and steam reaction (R5), water gas shift reaction (R7), and steam reforming reaction (R8) However, as mentioned previously, in the dual fluidized-bed system char is mostly transported from the gasifier to the combustor through the solid circulation As a result of less char remaining in the gasifier, the gas production from the reaction of char and steam (R5) is smaller In addition, as stated in Section 4.1, in the current system steam is injected through a few nozzles at the bottom of the gasifier A major disadvantage of steam jet is that it can cause channeling in fluidized-beds As shown in Fig 11, steam forms a channel in the fluidized-bed of the gasifier right after it is injected to the gasifier Through the channel, steam can quickly escape from the bed to the freeboard region and steam may not have enough time to have full contact with the solids and other gases Consequently, the gas productions of R7 and R8 also become smaller Due to the lower amount of char remaining in the gasifier and shorter residence time of steam in the bed, the impact of steam to biomass ratio for the dual fluidized-bed system is less significant than expected Effect of air supply To examine the effect of the air supply to the combustor, two additional cases are compared to Case The primary air supplies for Case and other two cases are 28.89 kg/h, 43.34 kg/h, and 57.79 kg/h, respectively In Fig 12, the concentrations of H2, CO, CO2, and CH4 from three cases are compared The gas compositions of three cases are only slightly different from each other, indicating that the air supply to the combustor has insignificant impact on the gas composition in the gasifier This finding is different from the conclusion for the single-reactor fluidized-bed systems that the producer gas composition changes significantly with air supply [72] As indicated previously, in the dual fluidized-bed system air is only introduced to the combustor and the combustor is Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 Mole Fraction (a) 13 0.42 0.37 H2 CO CO2 CH4 0.32 0.27 0.22 0.17 0.12 0.07 0.3 NHV of CO+H2 (kJ/kg bio.) (b) 0.5 0.7 Steam to Biomass Ratio 0.9 7250 7200 7150 7100 7050 7000 6950 6900 6850 6800 6750 0.3 H2/CO (Molar Ratio) (c) 0.5 0.7 Steam to Biomass Ratio (S/B) 0.9 1.22 1.21 1.2 1.19 1.18 1.17 1.16 1.15 0.3 0.4 0.5 0.6 0.7 Steam to Biomass Ratio (S/B) 0.8 0.9 Fig 10 e (a): Effect of steam to biomass ratio on producer gas composition (b): Effect of steam to biomass ratio on NHV of CO þ H2 (c): Effect of steam to biomass ratio on H2/CO (d): NHV increase of synthesis gas (CO þ H2) (e): Increase of H2/CO isolated from the gasifier Fig also confirms that O2 only appears in the combustor and is separated from other gases in the gasifier So, O2 is not directly involved in biomass gasification Increasing more air supply in the dual fluidized-bed system will not dramatically affect producer gas composition as long as the air supply is adequate for char and propane combustion in the combustor and the bed material circulation within the dual fluidized-bed system The effect of the air supply to the combustor on syngas and hydrogen production is also insignificant, because air is not directly used in the gasifier Conclusions Fig 11 e Time-averaged solid volume fraction In this study, a three-dimensional CFD model was established to simulate a pilot-scale dual fluidized-bed biomass gasifier using the MP-PIC method The CFD model was validated by experimental data at different operating conditions and good agreement was achieved As predicted by the CFD model, in the current dual fluidized-bed system no producer gas escaped to the combustor and no air leaked to the gasifier in the presence of steam as a sealing gas Consequently, all of the producer gas was free of N2 and was preserved in the gasifier for the downstream processing The effect of gasifier temperature was investigated As the gasifier temperature increased, H2 and CO concentrations increased, and CO2 and CH4 concentrations decreased The Please cite this article in press as: Liu H, et al., CFD studies on biomass gasification in a pilot-scale dual fluidized-bed system, International Journal of Hydrogen Energy (2016), http://dx.doi.org/10.1016/j.ijhydene.2016.04.205 14 i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y x x x ( ) e1 u V velocity, (m/s) computational cell volume, (m3) Greek symbols a volume fraction r density, (kg/m3) laminar viscosity, (m2/s) mlam meddy turbulent viscosity, (m2/s) Fig 12 e Effect of air supply study also showed that high gasifier temperature promoted syngas production and increased hydrogen content in producer gas The similar trends were also observed in the study of steam to biomass (S/B) ratio However, the effect of steam to biomass ratio was much smaller than gasifier temperature The effect of the air supply on producer gas composition, syngas, and hydrogen production was minor due to the fact that in the dual fluidized-bed system air was only supplied to the combustor and was not directly involved in biomass gasification Acknowledgments The authors acknowledge the financial support from the California Energy Commission Grant (CEC-PIR-14-023), West Biofuels (CEC-AVR-15-017), and the University of California Discovery Pilot Research and Training Program (Award 211974) Nomenclature Ap Cp CV Dp E f F g kd kmol keddy Mw Nu p Prt Re Sij Sc T particle surface area, (m2) specific heat at constant pressure, (kJ/(kg K)) specific heat at constant volume, (kJ/(kg K)) aerodynamic drag function Enthalpy, (kJ/kg) particle size distribution function interphase force between the gas and particle phases gravity, (m/s2) the thermal conductivity of the particle phase, (W/ (m K)) the molecular conductivity of the gas phase, (W/(m K)) the turbulent conductivity of the gas phase, (W/(m K)) molecular weight, (kg/mole) Nusselt number pressure, (Pa) turbulent Prandtl number Reynolds number Strain rate tensor turbulent Schmidt number Temperature, (K) Subscripts c char cp close packing g gas phase i,j coordinate index react reaction p particle 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