Cold flow model of dual fluidized bed: A review

Renewable and Sustainable Energy Reviews 53 (2016) 1529–1548

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Renewable and Sustainable Energy Reviews journal homepage: www.elsevier.com/locate/rser

Cold flow model of dual fluidized bed: A review Siddhartha Shrestha, Brahim Si Ali n, Mahar Diana Binti Hamid Department of Chemical Engineering, Faculty of Engineering, University of Malaya, 50603 Kuala Lumpur, Malaysia

art ic l e i nf o

a b s t r a c t

Article history: Received 16 September 2014 Received in revised form 24 May 2015 Accepted 17 September 2015 Available online 10 November 2015

Gasification in dual fluidized bed gasifier (DFBGs) has proven itself as a promising technology. Apart from gasification, dual fluidized bed (DFB) technology has also been utilized for CO2 capture-chemical looping combustion (CLC), calcium looping (CaL) and adsorption enhanced reforming (AER). Although pilot plants applying these technologies are available, still need improvement to be commercially viable. Fundamentally, the performance of the reactor depends upon the fluid dynamics within the reactor. So cold flow models (CFMs) are widely used in order to study the process fundamentals such as: pressure drop, solid fraction and solid circulation rate, to improve the operation and tackle the problem by troubleshooting. This paper outlines the application of scaling relationships to realize the hydrodynamic similarity among the industrial scale plant, the laboratory scale and CFMs of dual fluidized bed (DFB). Moreover, the occurring fluidization regimes and the stable operating regions in the cold model of DFB have been explored by reviewing the existing data in the literature. The pressure profiles, solid fraction profiles and the solid circulation rate obtained from the existing studies are presented and discussed together with the effect of the parameters influencing their behavior. & 2015 Elsevier Ltd. All rights reserved.

Keywords: Dual fluidized bed gasifiers Dual fluidized bed Calcium looping combustion Chemical looping combustion Adsorption enhanced reforming Cold flow models

Contents 1. 2. 3. 4. 5. 6. 7.

Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Application of scaling laws . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Fluidization regimes . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Operational map . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Pressure profile . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Solid fraction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Solid circulation rate . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.1. Effect of gas velocity in riser, inventory and particle property . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.2. Effect of air staging in riser. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.3. Effect of gas velocity in loop-seal . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7.4. Effect of gas velocity in BFB . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 8. Recommendations . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 9. Conclusion . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . Acknowledgment . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

1. Introduction The energy demand has been growing gradually due to the increase of population and the rapid development of the advanced technology. In response, contemporary world is looking for a n

sustainable alternative resources to fulfill the energy demandsupply cycle [1,2]. A few decades ago, the world was solely dependent on fossil fuels. However, rising problems like global warming and dearth of the natural energy resources have aroused scientists’ attention to realize the incorporation of some accrual

Corresponding author. Tel.: þ 60 3 79676896; fax: þ60 3 79675319. E-mail addresses: [email protected] (S. Shrestha), [email protected] (B.S. Ali), [email protected] (M.D. Binti Hamid).

http://dx.doi.org/10.1016/j.rser.2015.09.034 1364-0321/& 2015 Elsevier Ltd. All rights reserved.

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Notations A Ar Cs Cv D  dp dp De Fl Fr, FrD g G H h/l Iv Le M _ m P Q Re S/T Δt T U u* V V* w Δz

area (m2) Archimedes number (dimensionless) dimensionless solid loading (dimensionless) volume concentration of circulating solids (dimensionless) diameter (m) dimensionless particle size (dimensionless) particle diameter (m) ratios of solid to fluid densities (dimensionless) flow number (dimensionless) Froude number (dimensionless) acceleration of gravity (m/s2) solid circulation rate (kg/(m2 s)) height (m) height/base line ratio (dimensionless) total inventory, total particle mass load (kg) length ratio (dimensionless) dimensionless mass turnover (dimensionless) mass flow, (kg/s) pressure, (Pa) volumetric flow rate, (m3/h) Reynolds number (dimensionless) ratio of secondary to total air flow (dimensionless) time (s) temperature (°C) velocity (m/s) dimensionless gas velocity (dimensionless) Volume flow rate (m3/h) dimensionless volumetric flow rate (dimensionless) opening width of the riser exit (m) height of accumulated material (m)

Greek letter ΔP x

ϵ εs;core ρ μ ϕ Фd

pressure drop through component x (Pa) voidage (dimensionless) radially averaged solids fraction in the core of the riser closer to the riser exit density (kg/m3) viscosity (Pa s) sphericity (dimensionless) aerodynamic factor

Subscripts

alternative resources that could overcome the said problems [3–5]. A great deal of effort has been made to achieve this goal. These investigations introduced biomass as the most suitable substitute for fossil fuels [4,6–8]. It can also be used as additional raw material mainly generating heat and electricity [7,9–11]. Moreover, Biomass gasification can be one of the solutions for the existing pollution problems and the increasing need of energy [12]. Gasification is a thermo-chemical conversion of any solid fuels, such as, coal, petroleum, coke, plastics, biomass and solid wastes transforming into a valuable gas through partial oxidation at elevated temperatures with the help of a gasifying agent [13]. Depending on the gasification medium, the syngas calorific value ranges from 4 MJ/N m3 for air gasification to 16 MJ/N m3 for steam gasification [14–17]. On the basis of fluid dynamics, gasifiers can

AR b bot con cyc dc FR g gas ILS LLS/lls ls mf p pri t T s sc sec r, ris rc th top tr ULS v

air reactor bulk bottom connection cyclone down-comer fuel reactor gas gasifier internal loop-seal lower loop-seal loop-seal minimum fluidization particle primary gas injection terminal (applied to velocity), total solids supply chamber secondary gas injection riser recycle chamber thermal top transport (applied to velocity) upper loop-seal vertical

Abbreviations AER AR ASR BFB CaL CFB CFMs CHP CLC DCFB DFB DFBG FCC FR ICFB PSD SER

adsorption enhanced reforming air reactor air staging ratio bubbling fluidized bed calcium looping circulating fluidized bed cold flow models combined heat and power calcium looping combustion dual circulating fluidized bed dual fluidized bed dual fluidized bed gasifiers fluid catalytic cracking fuel reactor internally circulating fluidized bed particle size distribution sorbent enhanced reforming

be categorized as: fluidized bed, fixed or moving bed and entrained flow bed. Fluidized bed offers relatively high mixing and high reaction rates. Furthermore, fluidized bed is capable of being scaled up to medium and large-scale, overcoming limitations of small-scale, fixed-bed designs [18–20]. Recently, dual fluidized ded gasifiers (DFBGs) have gained interest of researchers’ due to their capability of producing high quality syngas. DFBGs, using steam as the gasifying agent, produce a syngas of high content H2 and CO with a calorific values ranging from 12 to 20 MJ/m3 nearly free from N2. In addition, the quality of syngas in DFBGs can be improved by optimized design and operation of gasifiers along with the use of catalytic bed materials, including internal reforming of tars and methane, and eventually a downstream

S. Shrestha et al. / Renewable and Sustainable Energy Reviews 53 (2016) 1529–1548

cleaning process [16]. Additional advantage of this technology is fuel flexibility [21–23]. The basic concept of the DFBGs is to divide the fluidized bed into two zones: a gasification zone and a combustion zone as shown in Fig. 1. The reactor arrangements can be any of the following types: (i) two bubbling fluidized bed (BFB), (ii) two risers, (iii) internally circulating fluidized bed (ICFB), (iv) circulating fluidized bed (CFB) (riser, combustor) with a BFB (gasifier), (v) riser (gasifier) with a BFB (combustor). Among the five types, CFB riser (combustor) with a BFB (gasifier) is the most efficient arrangement in terms of particle circulation, fuel conversion and tar production [24]. In a typical DFBGs as shown in Fig. 2 [25], gasification of biomass with superheated steam performed in BFB produces a valuable gas that flows out from the cyclone and can be used as gaseous fuel. The char and bed materials are transferred to riser via chute. The chute separates the two bed systems, ensuring only the solids (sand, un-reacted biomass and by-products of the gasification process, such as tar and char) can flow to the combustion zone restricting and preventing the mixing of gases from the gasifier to the riser and vice versa. In addition, it is suggested that the use of non-mechanical valves like loop seal or L-valve instead of chute enhances the regulation of the solid transfer and ensures better gas seal thereby preventing the mixing of BFB produce gas with the CFB flue gas [26]. The char is combusted in CFB riser in the presence of the bed materials to produce heat and flue gas, primarily N2, CO2, excess O2 and H2O. The heated bed materials from the CFB flow into the gas–solid separators like cyclones and are circulated to BFB via non-mechanical valve in order to supply the essential heat required for the endothermic gasification process. Moreover, the bed material can be used to achieve catalytic activity, CO2 capture, oxygen transportation, etc. Another promising dual fluidized bed (DFB) technology is chemical looping combustion (CLC) which offers efficient and lowcost CO2 capture [27]. The CLC is also composed of two interconnected fluidized beds where one of the reactor is fuel reactor (FR) and the other being air reactor (AR). The metal is oxidized in the air reactor (Eq. (1)) and transported to the fuel reactor. In fuel reactor, bed materials (metal oxide/oxygen carrier) oxidize the gaseous fuel to produce CO2 and water vapor (Eq. (2)). The CO2 thus formed can be readily separated by condensing water vapor thereby minimizing CO2 separation cost. Moreover, formation of NOx is also lessened. The reduced metal oxide is transferred back to the air reactor where it is re-oxidized and again circulated back to fuel reactor [28]. CLC can also be incorporated in DFBG to perform partial combustion of the fuel to reduce the tar and methane content [16]. Mex Oy  1 þ 1=2O2 -Mex Oy

ð1Þ

ð2n þ mÞMex Oy þ Cn H2m -ð2n þmÞMex Oy  1 þ mH2 O þ nCO2

ð2Þ

DFB, aiming to produce high hydrogen content in product gas as well as in situ CO2 capture known as-adsorption enhanced reforming (AER) or sorbent enhanced reforming (SER) is also under study [29–32]. In this process, the bed materials are used to capture CO2 apart from transferring heat. CO2 adsorbent bed material captures CO2 in situ in the process of gasification according to Eq. (3). This continuous removal of CO2 during gasification enhances the production of H2 (Eq. (4)) [31]. CaO þCO2 -CaCO3

ð3Þ

CO þ H2 O2CO2 þ H2

ð4Þ

Along with these pre-combustion CO2 capture process, DFB has also been employed in post-combustion CO2 capture known as calcium looping (CaL) process [33,34]. In this process, one of the reactors is carbonator and the other is regenerator. CaO is used as

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CO2-sorbent. The CO2 in the flue gas from the conventional power plant is absorbed by moles of sorbent in carbonator (Eq. (3)), thus CO2-lean flue gas exits the carbonator to the atmosphere. The carbonated sorbent is transferred to the regenerator where it is desorbed at a higher temperature to produce CO2 and steam (reverse reaction of Eq. (3)). CO2 rich stream is obtained by the condensation of water vapor which can be compressed and stored. All these DFB processes discussed above are still in demonstration stage and still need maturity to be commercially utilized. A lot of research have been done to improve the performance of such DFB reactors and are also available in the literature. Among all of them, the study of hydrodynamics of the DFB technology is one of the important and common aspects that have received huge attention of researchers. In order to design an optimized reactor, regardless of the process the DFB being operated, it is very crucial to understand the reactor geometry, bed mass and solid residence time, gas–solid contact efficiency, heat and mass transfer rate, chemical reaction performance and gas leakage in these reactors. Fundamentally, these parameters depend upon complexity behavior of the flow within the reactors. Significantly, the understanding of hydrodynamics is the key to the successful modeling, design and scale up of these reactors [35–40]. Difficulty in measurement of hydrodynamic parameters in an industrial scale plant under hot operating conditions makes them more superficial rather than actual. Hence, cold flow models (CFMs) under the ambient conditions are widely applied to study the hydrodynamics and fluid dynamics. Moreover, in the academic, a scaled CFMs is built to resemble the hydrodynamics of a larger hot rig by using scaling relationships whereas full-scale CFMs are built in the case of industrial scale to better understand the influence of the design solutions and operation on the system performance, to test process control and measurement methodologies [41]. In addition, easy handling, minimal requirements on experimental equipment and control techniques, and the possibility of visual observation in the macroscopic flow structures make CFMs cheap, comfortable and a reliable source to obtain important fluid dynamic data [42]. This paper aims to report the findings in the cold model DFB. Beginning with the application of fluidized bed scaling laws in DFB design, the paper illustrates the fluidization regimes and its operational map in the cold model of DFB. In addition, the literature is reviewed to highlight the behavior pressure profile, solid fraction profile and solid circulation rate in DFB. The information on the influencing parameters is extracted from the literature and conclusions is made in the light of previous findings. On the basis of the findings, recommendations for the future research are presented.

2. Application of scaling laws Although cold models are mainly constructed to improve the performance of industrial scale plant, data obtained from them are not applicable if relations between the two distinctive sized setups are unknown since both setups operate under different operating conditions. Therefore, the small-scale experiments based on the principle of dynamic similarity are used to obtain information for full-scale system [35]. A scaled-down cold model is designed with the solicitation of fluidized bed scaling laws for an industrial system operating at high temperatures to study the hydrodynamics and significantly improves the operation of an existing plant [42– 47]. A number of sets of dimensionless scaling parameters have been designed on the basis of the governing equations of conservations of mass and momentum of fluid and particles to ensure hydrodynamic similitude at different scales of reactors [48–55]. Typical sets of dimensionless scaling parameters are listed in Table 1 [56]. Apart from the bed geometry ratio, particle sphericity

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(φ) and particle size distribution (PSD), the scaling parameters also include the Froude number (Fr), the Reynolds number (Re), the Archimedes number (Ar), the Flow number (Fl), the ratio of solid to fluid densities (De), dimensionless solid loading  C s ¼ ρGUs g , p which all describe the following eight variables: superficial gas velocity (Ug), solid circulation rate (Gs), sauter mean particle diameter (dp), particle density (ρp), riser hydraulic diameter (D), gas density (ρg), gas viscosity (μg) and gravity (g). They are generally sufficient to control the detailed hydrodynamics of gas–solid flow in risers [55], with an assumption that the restitution among particles, the friction between particles and column wall, the electrostatic forces and the cohesion should be neglected. To obtain the hydrodynamic similarity between the cold model and the full-scale plant, these scaling parameters should be matched and verified. Foscolo et al. [57] described the visual observations of solid motion in an industrial fluidized bed and a down-scaled model, which demonstrated the dynamic similarity of the two flows, and hence suggested that the four dimensionless scaling parameters (Ar, De, Fl, Le ¼L/dp) were sufficient to achieve dynamic similarity. In their study, the scaled-up plant showed a uniform bed temperature, indicating a proper mixing of fuel as observed in the cold model. Kronberger et al. [47] estimated the solid hold-up by pressure measurements according to the different levels of the riser height and compared them to the dimensionless specific circulation rate in terms of Eq. (5). The results suggested the analysis could be performed in the laboratory prototype.

Fig. 1. Concepts in DFB (extended from [16]).

1ε ¼

M¼

Table 1 Typical sets of dimensionless scaling groups [56].

[55]

Dimensionless scaling groups Five dimensionless groups Four dimensionless groups

Gs U g ρp U g ; U t ;

Three dimensionless groups

Gs U g ρp U g ; U t ;

Two dimensionless groups

Gs U g ρp U g ; U t ;

[52,48]

Full set

[54,49]

Simplified set

[53,50]

Gs U g ρp U g ; U t ;

Viscous-limit set

Gs ρp U g ; Gs ρp U g ;

U2 ρ

FrD ¼ gDg ; ρp ; g

U 2g

ρg U g dp μg

ρ

FrD ¼

p

3

ρp ρg gdp dp ;D μ2g U2 ρ

or

Ug U mf ;

FrD ¼ gDg ; ρg

or

Ug U mf ;

FrD ¼ gDg

2

Qi et al. [56] investigated the hydrodynamic similarity in the fully developed zone of co-current upward gas–solid two phase flow systems under different operating conditions by measuring the axial profiles of the pressure gradient, the radial profiles of solid concentration and the particle velocity in two CFB risers of 15.1 m and 10.5 m high, with FCC and sand particles respectively. Under such circumstances, the scaling parameter- ρGUs g was modp

U2 ρ

ρp U g dp Gs ρp U g ; μg D

ð6Þ

ified to FrD 0:3 ρGUs g to achieve a detailed hydrodynamic similitude of

U 2g gD

FrD ¼ gDg ; ρg ; Ar ¼ μg D

ṁ D : Iv U t

p

FrD ¼ gD; ρpg

2 ρp U g dp

ð5Þ

Although these sets of dimensionless parameters have already been used to scale the fluidized beds for ages, researchers are still committed to modifying these parameters to fulfill functional and operational conditions. For the CFB riser, Kehlenbeck et al. [46] introduced the scaling parameter-the dimensionless mass turnover (M), demonstrating solid circulation rate is a function of superficial gas velocity in the riser and the total mass load in the system, rather than- C s ¼ ρGUs g as is shown in Eq. (6). They argued that Cs is not a p parameter that allows the prediction of solid circulation rate served as a function of the total mass load in the system, let alone for different bed materials. This parameter can be used for predicting satisfactory values of solid circulation rates for a wide range of particle properties
 170 r μm r dp r 860 r μm; 1480 kg=m3 r ρp r 8900 kg=m3 and particle mass load. Finally, they concluded that solid circulation depends on its particle-size distribution which is a scaling parameter that should be matched between a cold model and an industrial plant.

Fig. 2. Typical DFB configuration [25].

Ref.

Gs ðU  U t Þρp

p

U2

the gas–solid flow in the fully developed zone of the risers. Conclusions made were based on ambient conditions with air as the fluidization medium and it was emphasized that, the scaling parameter should be confirmed under both high temperature and pressure conditions with different gases. Although research have been carried out to modify the scaling parameters as discussed above, the scaling laws shown in Table 1

Table 2 Previous studies in cold model of DFB. Ref.

Design and dimensions

Kreuzeder et. al. [42]

Goo et al. [84]

ConFig.uration: CFB/BFB CFB: H¼ 2.2 m, D ¼ 0.175 m Hpri ¼ 0.175 m, Hsec ¼ 0.250 m BFB: H¼ 1.3 m, D ¼ 0.55 m Connection: Upper: loop-seal Lower: chute ConFig.uration: CFB/BFB Capacity: 30 kWth CFB: H¼ 6 m, D ¼ 0.075 m BFB: H¼ 2.1 m, D¼ 0.2 m Connection: 2 loop-seal Hls ¼ 0.2 m Dls ¼ 0.12 m

Sung et al. [78]

Karmakar and Datta [80]

ConFig.uration: CFB/BFB CFB: D ¼0.07 H¼ 5.18 BFB: 0.2 m wide  0.2 m depth  2.95 m high Typeb: 1. Below the bottom of the gasifier and distributor installed on a slat wall. 2. On the side of the gasifier which is upper than the distributor is installed on a slat wall. ConFig.uration: CFB/BFB

CFB: H¼ 5.95 m, D ¼0.050 m Hsec ¼ 0.2 BFB: H¼ 2.1 m, D¼ 0.1 m Standpipe: H ¼5 m, D¼ 0.025 Connection: 2 L-valves

ConFig.uration: CFB/BFB

Scaling criteria

Ratioa

Gas leakage

Bronze dp ¼ 180 μm ρp ¼ 8900 kg/m3

Medium: air for CFB 55:45 % He:air for BFB Iv ¼ 1–15 kg

Glicksman [52]

One-fifth scale cold model of a CFB pilot plant.

N/A

Bronze

Medium: air@ ambient conditions ρg ¼ 1.28 kg/m3 S/T ¼0.15–0.5 Qr,bot ¼ 0.0055–0.016 m3/s Qbfb ¼ 0.014 m3/s Ur,sec ¼ 2.9–4.6 m/s Iv ¼ 1-5–130 kg T ¼ 40 °C

Glicksman [52]

The dimensions of the cold model and properties of the used solid were calculated by downscaling industrial sized power plant (8 MWth)

N/A

Glicksman [52] Deviation: Rep ¼ 1 Frp ¼1 ρp/ρg ¼ 0.83 Re¼ 0.95 H/dp ¼0.99 D/dp ¼ 0.79 N/A

Gas density ¼3.62 Particle density ¼3 Particle diameter ¼0.24 Velocity ¼ 0.52 Volume flow ¼0.09 Solid flux¼ 1.76

N/A

N/A

Gas bypassing from the loop-seal connecting reactors is found to be less than 0.2 %.

Medium: dried air for gasifier/ N/A CO2 for lower loop-seal

N/A

Concenration of CO2 was measured in the gasfier to measure the gas leakage from lower loop-seal to gasifier. Gas bypassing fraction of Type 1 was found to be less than Type 2.

Medium: air Qr,pri ¼ 3–5 m3/h

N/A

N/A

N/A

N/A

N/A

N/A

dp ¼ 119 μm ρp ¼ 8750 kg/m3

Bronze Range¼ 40–280 μm dp ¼ 125 μm ρp ¼ 8730 kg/m3 ρb ¼ 5370 kg/m3 Ut ¼ 0.38 m/s ϕ ¼1 Silica sand dp ¼ 260 μm ρp ¼ 2500 kg/m3 ρb ¼ 130 kg/m3 Umf ¼0.052 m/s Ut ¼1.57 m/s Utr ¼ 3.07 m/s Silica sand Group: B dp ¼ 376.1 μm

Medium: air ρg ¼ 1.05 kg/m3 μg ¼19.9  10  6 Pa s QT ¼ 450 N m3/h U ¼6.25 m/s Gs ¼120 kg/m2 s Ur ¼ 2.75–4 m/s Ubfb ¼ 0.052–0.156 m/s (1–3 Umf) Uls ¼0.135–0.362 m/s (1.5–4 Umf)

ρb ¼ 1220 kg/m3

Umfc ¼0.12 m/s/0.1 m/s Ut ¼1.3 m/s Silica sand Range¼ 50–300/75–425, 106–500, 150–600 μm dp ¼ 147/211/334/416 μm ρp ¼ 2650 kg/m3 ρb ¼ 1696/1696/1710/1722 kg/m3 ϵmf ¼0.46/0.44/0.41/0.40 Umf ¼0.018, 0.037, 0.091, 0.138 m/s Ut ¼1.2/1.7/2.7/3.4 m/s Ar ¼ 281/835/3292/6347 ϕ ¼0.86 Silica sand dp ¼ 250 μm

Qr,sec ¼ 25–50 m3/h Qbfb ¼ 3–16 m3/h

Ur ¼ 3.5–4.25 m/s Ubfb ¼ 0–0.27 m/s

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Seo et al. [86]

Operating conditions

S. Shrestha et al. / Renewable and Sustainable Energy Reviews 53 (2016) 1529–1548

DFBGs Kehlenbeck et al. ConFig.uration: CFB/BFB [46] CFB: H¼ 0.054 m BFB: D ¼0.18 m Standpipe: D ¼ 0.024 m Connection: Upper: loop-seal Lower: chute Kaiser et. al. [44] ConFig.uration: CFB/BFB CFB: H¼ 2 m, D ¼ 0.17 m BFB: Dbot ¼ 0.25 m, Dtop ¼0.55 m Connection: Upper: loop-seal Lower: chute

Particle properties

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Table 2 (continued ) Ref.

Design and dimensions

Scaling criteria

Ratioa

Gas leakage

N/A

N/A

N/A

Farrell [105] ρg/ρp ¼ 1.08 Ar¼ 0.92 Fl¼ 1.14 Cs ¼ 0.22

Gas density ¼3.93 Particle density ¼3.576

N/A

N/A

N/A

N/A

N/A

N/A

N/A

dp ¼ 250 μm ρp ¼ 2466 kg/m3 ρb ¼ 1281 kg/m3 Umf ¼0.06 m/s Ut ¼1.403 m/s Ε ¼ 0.48

Medium Compressed air ρg ¼ 1.229 kg/m3 μg ¼1.7  10  5 kg/m s Ur ¼ 3.5/4 m/s Ubfb ¼ 0.14 m/s Uls,rc ¼ 0–0.18 m/s Uls,sc ¼ 0.06 m/s Uls,v ¼ 0.06 m/s Iv ¼ 35 kg

Glass beads

T ¼ 25 °C

Glicksman [52]

Mass ¼ 0.17

Gas tracing method using propoane to measure gas leakage. Gas leakage was found to be low and mechanisms were explored. Effective measure to reduce gas leakage were suggested.

Particle properties 3

Nguyen et al. [91]

CFB: (0.04 m  0.11 m  4.5 m high) ρp ¼ 2466 kg/m ρb ¼ 1281 kg/m3 Umf ¼0.0603 m/s BFB: (0.285 m  0.11 m  2.13 m Ut ¼1.403 m/s high) Connection: loop-seal Utr ¼ 2.5 m/s (0.178 m  0.11 m  0.41 m high) ConFig.uration: CFB/BFB Silica sand CFB: (0.04 m  0.11 m  4.5 m high)

CFB: H¼ 2.6 m, Dbot ¼ 0.18 m, Dtop ¼ 0.21 m

Manchasing et al. [71]

Wang et al. [90]

Hsec ¼ 0.15 m , Hdiffuser ¼ 0.5 m ConFig.uration: CFB/Downer CFB: H¼ 3 m, D ¼ 0.14 m Downer: H¼ 1.12, D ¼ 0.40 m Connection: Screwfeeder Gas–solid separator : seperating unit designed as a large empty emty chamber at the top of riser. ConFig.uration: CFB/BFB

Umf ¼0.0603 m/s Ut ¼1.403 m/s Copper dp ¼ 138 μm ρp ¼ 8940 kg/m3 Umf ¼0.10 m/s ϕ ¼0.60 Silica sand dp ¼ 370 μm ρp ¼ 2600 kg/m3

Silica sand

CFB 0.04 m  0.11 m  4.5 m high

BFB: 0.285 m  0.11 m  2.13 m high Connection: Loop-seal: 0.178 m  0.11 m  0.41 m high CLC Kronberger et. al. [47]

ConFig.uration: CFB(ARd þ Riser)/ BFB (FRe) CFB: DAR ¼ 0.14 m, HAR ¼ 0.53 m

Prӧll et al. [58]

Driser ¼0.072 m, Hriser ¼ 1.85 m BFB: DFR ¼ 0.25 m, HFR ¼ 0.34 m Connection: two loop-seals ConFig.uration: [DCFBf]

Range: 40–80 μm dp ¼ 67 μm ρp ¼ 2550 kg/m3 ϕ ¼1 Bronze

Uls,v ¼ 0–0.075 m/s Uls,sc ¼ 0–0.075 m/s Uls, rc ¼0–0.36 m/s h/l (height/base line)¼ 1.25– 8.0 Iv ¼ 35 kg. T ¼ 25 °C, P¼ 1 atm Ur ¼ 3.5–4 m/s Ubfb ¼ 0.12 m/s Uls,v ¼ 0.06 m/s Uls,sc ¼ 0.06 m/s Uls, rc ¼0.0–0.18 m/s h/l (height/base line)¼ 2.5 Iv ¼ 35 kg. ρg ¼ 1.18 kg/m3 μg ¼18  10  6 Pa s Qpri ¼ 32–96 m3/h Qsec ¼107–320 m3/h Qls ¼5–9 m3/h Iv ¼ 6–10 kg ρg ¼ 1.225 kg/m3 μg ¼1.7984  10  5 kg/m s Ur,pri ¼ 0.28882 m/s Ur,sec ¼ 4.81194 m/s Iv ¼ 10 kg

Medium: He/N2 Iv ¼ 1.1–2.2 kg Ur,sec ¼ 4–10 Ut Ur,pri ¼ 1.2–3 Ut Ubfb ¼ 5–15 Umf Uls ¼1.2–4 Umf ρg ¼ 1.25 kg/m3 μg ¼1.79  10  5 Pa s UAR ¼ 4.25 UFR ¼ 1.21 QAR ¼30.3 N m3/h QFR ¼ 10.1 N m3/h

Length ¼ 0.55 Area ¼ 0.552 Velocity ¼ 0.74 Volume flow ¼0.22 Solids flux¼ 0.23 Glicksman [52] Deviation (AR/FR) Rep ¼0.50/0.52 Ar¼ 0.36/0.43 Fr¼ 1/1 ρp/ρg ¼ 1.39/ 1.53 D/dp ¼ 1.01/0.99

Linear geometric scaling factor of 1:3

N/A

S. Shrestha et al. / Renewable and Sustainable Energy Reviews 53 (2016) 1529–1548

BFB: (0.285 m  0.11 m  2.13 m high) Connection: loop-seal (0.178 m  0.11 m  0.41 m high) Lim et al. [25,26] ConFig.uration: CFB/BFB

dp ¼ 250 μm ρp ¼ 2466 kg/m3 ρb ¼ 1281 kg/m3

Operating conditions

CFB (AR)/CFB (FR) DAR ¼ 0.050 m DFR ¼ 0.054 m Bischi et al. [97] ConFig.uration: CFB (AR)/CFB CFB(FR) AR: H¼ 5 m, D¼ 0.23 FR: H¼ 5 m, D ¼ 0.144 m Connection Upper: loop-seals Lower: bottom extraction lift and Shuai et al. [106] ConFig.uration: CFB/BFB

dp ¼ 54 μm QLLS ¼ 1.5 N m3/h ρp ¼ 8730 kg/m3 QULS ¼ 1.0 N m3/h ϕ ¼1 QFR,ILS ¼ 0.15 N m3/h Fe–Si alloy (rounded irregular shape) Medium: air dp ¼ 34 μm U ¼2.2 m/s ρp ¼ 7000 kg/m3 Group A

Silica sand

CFB: H¼ 1.9 m, D ¼ 0.19 m BFB: H¼ 0.5 m, D ¼0.19 m

Markström and Lyngfelt [107]

Fe–Si alloy with 80% iron

CFB(FR) AR: H¼ 5 m, D¼ 0.23 FR: H¼ 5 m, D ¼ 0.144 m Connection Upper: loop-seals Lower: bottom extraction lift and

dp ¼ 34 μm ρp ¼ 7000 kg/m3

ρg ¼ 1.188 kg/m3 μg ¼1.82  10  5 Ns/m2 UAR ¼ 2.4 m/s UFR ¼ 2.4 m/s

Concept: CLC ConFig.uration: CFB (AR)/CFB (FR) with a Circulation Riser (CR) , Carbon Stripper (CS) and 4 loop-seals

Silca sand (GA39) ρp ¼ 2650 kg/m3 dp ¼ 92 μm

T ¼ 20 °C μg ¼19.72  10  6 Pa s UAR ¼ 600 (Ln/min)

φ ¼0.75

CaL Charitos et al. [45]

Diego et al. [82]

Lisbona et al. [81]

ConFig.uration: CFB(carbonator)/ BFB(calciner) CFB: H¼ 5.3 m, D¼ 0.030 m BFB ¼D ¼ 0.049 Connection: upper ¼ double exit loop seal and cone vale Lower: loop-seal

N/A

Glicksman [54]

Scaled model of 30 Mw CLC pressurized system

Numerical simulation was formulated to predict the gas leakage.

Simulations indicated the leakage falls as the total solid inventory rises in the CLC reactor system. Glicksman scaling relationships Deviation (AR/FR) Fr¼ 0.85/0.85 De ¼1.26/1.36 Gs ¼ 1.20/1.20 Uo/Umf ¼3.13/2.77 Geometric similarity ¼ 0.30/ 0.30 Glicksman [51]

UFR ¼ 348 (Ln/min) UCR ¼50 (Ln/min) UCS ¼ 99 (Ln/min) ULS1 ¼ 12 (Ln/min) ULS2 ¼ 13 (Ln/min) ULS3 ¼ 12 (Ln/min) ULS4 ¼ 12 (Ln/min)

ZrO2

ρg ¼ 1.188 kg/m3

Glicksman [52]

dp ¼ 142/230 μm ρp ¼ 5700 kg/m3

μg ¼18  10  6 Pa s

and

dp ¼ 200 μm

U ¼5 m/s

N/A

ρp ¼ 2000 kg/m3 N/A

Ur ¼ 2.35/2.17/2.68/2.04 m/s

Glicksman [108]

ΔP x ρp gD

[48]

1:1 cold model of 150 kWth

N/A

58% size model of 100 kW Gas density ¼1 Length or diameter ¼ 0.577 Particle diameter ¼0.577 Velocity ¼ 0.759 Volume particle flux ¼0.759 Time ¼ 0.759

N/A

Downscaled by a factor of 2.33 from N/A 10 kWth CaL DFB system. Gas density ¼3.02 Particle density ¼3.16 Gas viscosity ¼ 0.46 Mass ¼ 0.23 Length ¼ 0.42 Area ¼ 0.42 Velocity ¼ 0.65 Volume flow ¼0.22 Solids flux¼ 2.05 N/A N/A

Scaled down cold model from 350 kW pilot plant

N/A

Iv ¼ 7.18/8.18/4.192/3.972 Uls ¼0.087/0.088/0.067/ 0.09 m/s

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ConFig.uration: CFB(Carbonator)/ CFB(Calciner) H¼ 35 m D ¼ 15.95 m ConFig.uration: CFB (carbonator)/ CFB(calciner) CFB: H¼ 4 m, D1¼ 0.17 m D2 ¼0.16 m Standpipe: H ¼1.447 m D¼ 0.08 m Loop-seal: H¼ 0.3 m, D ¼ 0.177 m

1:1 cold model of 150 kWth

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Bischi et al. [41]

Pot-seal: H¼ 0.15 m, D ¼ 0.14 m Downcomers: D ¼0.04 m ConFig.uration: CFB (AR)/CFB

dp ¼ 150 μm ρp ¼ 2600 kg/m3

T ¼ 20 °C ρg ¼ 1.188 kg/m3 μg ¼1.82  10  6 Pa s UAR ¼ 0.75–1.15 m/s UFR ¼ 0.09–0.31 m/s Uls ¼0.065–0.097 m/s Gs ¼26.0–62.5 kg/m2 s Iv ¼ 9 kg T ¼ 20 °C

N/A

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Table 2 (continued ) Ref.

Design and dimensions

Particle properties

Cotton et al. [83] ConFig.uraiton: Entrained flow car- Limestone bonator (EFC)/BFB calciner EFC: H¼ 4.2, D ¼0.1 m Range¼ 125–250 μm

ConFig.uration: CFB/BFB CFB: H¼ 6 m, D ¼ 0.041 m BFB: H¼ 2.0 m, D ¼0.13 m Lower loopseal, D ¼ 0.041 m

Bidwe et al. [79] ConFig.uration: CFB/BFB CFB: H¼ 4 m, D ¼ 0.069 m BFB: H¼ 2.4 m, D ¼0.14 m

a

Ratioa

Gas leakage

Uefc ¼1.91 m/s

Glicksman [54]

(Carbonator/Calciner)

Ubfb ¼ 0.23 m/s

Deviation ¼ (Carbonator/Calcinier) Re¼ 0.31/0.23 Ar¼ 0.1/0.07 FR¼ 2/2.7 De ¼2.1/3.5 D/dp ¼ 1/1 Glicksman [53]

Gas density ¼2.10/2.79 Particle density ¼1/1 Particle diameter ¼1 Velocity ¼ 0.69/0.69 Area ¼ 1/1

Gas tracing method using N2 to measure gas leakage. Greatest amount of gas bypass may occur from LLS to calciner at high loopseal velocities.

Geometric scaling ratio¼ 1:2.5

N/A

Gas tracing method using He to measure gas leakage. Gas leakage was found from the riser to the BFB reactor. The main route for the gas to leak into the BFB was via down-comer of the primary cyclone. The amount of gas leaking to the BFB was found to be function of the pressure on the BFB side. Additionally, it was found that high loop-seal aeration velocities help to minimize the gas leakage from the riser to the BFB side. N/A

Iron oxide (Fe3O4)

T ¼ 20 °C

Ragne ¼ 100–200 μm

Carbonator, medium¼ air

dp ¼ 166 μm ρp ¼ 5170 kg/m3

ρg ¼ 1.18 kg/m3 Regenerator, medium ¼Air/ CO2 ρg ¼ 1.26 kg/m3 Ur ¼ 2.5–4 m/s Gs ¼10–45 kg/m2 s Iv ¼ 5.5–9 kg

Limestone Range¼ 212–1180 μm dp ¼ 438 μm ρp ¼ 2870 kg/m3 ρb ¼ 1364 kg/m3 Umfg ¼ 0.15, 0.14, 0.11 m/s Uth ¼2.52, 1.94, 1.02 m/s

Medium : compressed air T ¼ 25/80/250 °C ρg ¼ 1.19/1/0.68 kg/m3 Ur ¼ 2.5–6.5 m/s Ubfb ¼ 0.24 m/s Uls ¼1–7 Umf Iv ¼ 12 kg Gs ¼25–139 kg/m2 s

Glicksman [52]

N/A

Steel powder/Iron oxide (Fe3O4) Range¼ 70–200 μm dp ¼ 110/119 μm Umf ¼0.03/0.02 m/s ρp ¼ 7500/5100 kg/m3

T ¼ 20 °C CFB Medium ¼ air ρg ¼ 1.18 kg/m3 U ¼5.6–4 m/s Gs ¼13–20i kg/m2 s BFB, medium¼ Air þHe ρg ¼ 0.9 kg/m3 U ¼0.26–0.52 m/s Iv ¼ 8–13 kg

Glicksman [54]

Gasifier and other components. Geometric scaling ratio¼ 1:2.5 Regeneratior Geometric scaling ratio¼ 1:3 Velocity ¼ 0.57 Solids flux¼ 1.63

Cold model: hot model. On the basis of way-out of solids from the gasifier to the CFB riser connected by loop-seal. c For type1/type2, minimum fluidization velocity can be influenced by shape and size of the reactor and distributor. d AR¼ air reactor. e FR ¼fuel reactor. f DCFB¼ dual circulating fluidized bed. g For temperatures 25, 80, 250 °C. h For temperatures 25, 80, 250 °C. i Use of iron oxide particles for extrapolation of scaled performance of the regenerator. b

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SER/AER Ramirez [85]

Scaling criteria

Uls ¼0.66 m/s Iv ¼ 4.5 kg/6 kg

BFB: H¼ 1.3, D ¼0.165 m Connection: loop-seal

Bidwe et al. [87] ConFig.uration:CFB (Carbonator)/ CFB (Calciner) Carbonator: H¼ 4 m, D ¼ 0.092 m with enlarged bottom section. Calciner: H¼4 m, D ¼0.069 Loop-seals: double exit loop-seal with one weir typ exit back into the own riser while the second exit is a cone valve mounted on the supply side of the loop-seal

Operating conditions

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are widely accepted and used to scale and design the fluidized beds. Similarly, these scaling laws are also seemingly popular and applied in the case of DFB. Table 2 lists several studies conducted in the DFB, followed by their design and dimensions, particle properties, operating conditions, scaling criteria, ratio and gas leakage between the reactors. It can be seen that the scaling relationships by Glicksman is popular and widely used. In Table 2, along with the designed setups using the scaling relationships, works with suit designs used to investigate the fluidization performance in DFB are also shown, which are also equally valuable and provide insights to the researchers. The particle and their properties used in these two different setups are not alike, mostly

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bronze, glass beads, copper, ZrO2 and Fe3O4 have been used in the scaled models whereas sand, dolomite dominates in the suit designs. This difference arises since scaled models tend to match the scaling parameters, and while fulfilling the dimensionless set, the particles are chosen to meet the scaling ratio. For example, Lim et al. [26] selected copper particles in his experiments. According to the scaling relationship, Density ratio (De) in hot and cold model should be same. As gas density of 0.3 kg/m3 is common in an industrial gasifier and sand has a density of 2500 kg/m3. Thus, Dehot ¼ 0.00012. In order to maintain the same De for the cold model operated at ambient temperature (with air density of 1.18 kg/m3), copper particles with the density of 8940 kg/m3 were selected yielding Decold ¼ 0.00013. Hence Dehot EDecold. However, in real-world it is not easy to match all the parameters in the scaling relationships [58] and deviation is always observed. The observed deviation in matching the dimensionless parameters obtained from the studies reviewed in this paper is shown in Table 2 (scaling criteria). One of the ways to minimize the deviation is to construct CFMs of similar size with hot rig and utilize the particles with same properties [41,43]. In general, proper ratio of reactor geometry, gas solid mixture and the operating conditions are selected to fulfill the scaling criteria. Briefly, the scaling laws are applied to achieve similar flow behaviors in a laboratory scale cold model that is observed in the hot plant; which is of great importance to perceive the fluidization regime encountered in the actual operating conditions.

3. Fluidization regimes

Fig. 3. Different regimes of fluidization [61].

Fluidization is an operation in which fine solids are transformed into a fluid like state through contact with gas or liquid [59]. Fluidization occurs when the supplied gas or liquid exerts sufficient force on the particles that exactly counter their weight. Fluidization process involves the suspension of a very large

Fig. 4. Regime map (a) idealized flow regime map for gas–solids transport [104] and (b) general flow regime diagram for vertical pneumatic conveying and fluidized bed systems [66].

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number of solid particles in an upward flowing fluid [60]. This interaction between the solid bed material and the fluidization gas governs the flow behavior of the bed. Correspondingly, the performance of a fluidized bed gasifier drops significantly when the operation deviates from the designed fluidization regime. Thus, a good understanding of the gas–solid motion in the furnace or reactor of a fluidized bed unit is very important [59]. Fig. 3 [61] shows different regimes of fluidization with the effect of the increasing superficial gas velocity. As the velocity increases the fluidization regimes transits from a fixed bed to bubbling, slugging, turbulent, fast fluidization and to pneumatic conveying regime, transporting the bed material out of the reactor. Characterization of these regimes along with flow regime map and methods to identify the transition velocities can be found in the literature [60,62–69]. Fig. 4 shows two different regime maps that are helpful for designers. (1) Ar as the abscissa axis and V  , where " #1=3 " # ρ2g Gs ε U V ¼ ρp ð1  εÞ g μg ðρp  ρg

ð7Þ

(2) Ar as the abscissa axis and (modified Reynolds number) Re where  Re ¼ Re 1 

ε

Cv ð1  εÞ ð1  C v Þ

 ð8Þ

where C v ¼ V̇s V̇þs V ̇f , V being the volume flow rate. In the case of the DFB composed of CFB–BFB having contrasting flow behavior in each reactor, the CFB riser operates in the fast fluidization regime whereas gasifier operates in the bubbling regime. The latter provides good solid gas mixing, uniform temperatures and high reaction rates as compared to those of fixed bed gasification. BFB has greater tolerance to the range of particle size and is safer in operation due to its good temperature control, in addition, tar cracking and reforming catalysts can also be added to the bed inventory [57]. The fast fluidization exists between the dilute and dense pneumatic transport regimes, and fluidization velocity is kept well above the transport velocity (Utr) of the largest bed material [70]. High slip velocity between the gas and solid, formation and disintegration of particle agglomerates and excellent mixing and transport mechanisms are the major characteristics of this regime [35,37,56,59]. DFBG with configuration riser-downer has also been reported [71]. Different fluidization regimes were observed in their study which showed bubbling and slugging flow patterns in the riser section with calculated solid volume fraction. Furthermore, in the case of DFB for CLC applications, BFB as FR can be problematic since there is the possibility of unconverted fuel in the bubble phase, which is expected to increase by high fluidization numbers and small particle sizes. So, CFB as FR is advantageous for CLC applications [58]. As can be seen in Table 2, almost all studies for CLC utilize two interconnected CFB as AR and FR respectively. Similarly for CaL, DFB configuration (carbonator-regenerator/calciner) as CFB–BFB, CFB–CFB and also entrained flow-BFB is available whereas for AER, CFB–BFB persists. According to the configuration and operational conditions, the fluidization regimes varies, so it is important to identify the fluidization regimes in which the reactor is to be operated. Moreover, within CFB operated in fast fluidization, different flow regimes exists. Loffler et al. [72] studied the fluid dynamics in a riser with three gas injections and a diffuser. Regime map for riser was derived according to Bi and Grace [73,74] as shown in Fig. 5, which manifests the fluidized bed is in bubbling regime below the

Fig. 5. Regime map of the riser [72].

primary air injection. Adding the primary air, a turbulent fluidized bed was formed while the solid entrainment velocity was exceeded above the secondary air. Above the diffuser, the gas velocity was reduced again and had almost the same value as the primary air injection. Similarly, Monazam et al. [65] characterized the flow regimes of CFB with respect to the time required to empty the solid out of the riser at different gas velocities. Three regions were observed when the emptying time was plotted against the gas velocity; the dense phase turbulent flow, fast fluidization flow and pneumatic conveying regime respectively. After the fluidization regimes have been identified, it is necessary for designers to learn the stability of the operational conditions. Consequently, an operational map is of practical significance.

4. Operational map Fluidization process is inherently a complex process and its operation is relatively complicated. DFB technology utilizes two fluidized bed simultaneously making it more difficult in its operation. Taking these into consideration, accordingly, numerous researches in DFB have been conducted to report the operational stability under various operating conditions. Charitos et al. [45] used a 10 kWth DFB system for CaL, with a riser carbonator and a BFB regenerator to present the operational regions of the cold model. They defined the operating behavior of the riser with respect to the gas velocity, riser pressure drop and total solid inventory. A region of stable operation bordering with the maximum and minimum velocity was identified. At a given total solid inventory, slugging occurred below the minimum riser velocity, while unstable operation was observed above the maximum velocity resulting from the variation of the riser pressure drop and solid looping rate with respect to different periods of time. Recently loop-seals have been incorporated as a non-mechanical valve and widely used in the fluidized bed systems since they tend to provide effective control of solids flow [75]. These loop-seals build necessary pressure drop in a solid circulation loop to convey particles from low pressure zone to high pressure zone without undesirable inverse gas flow [76,77]. With the application of the loop-seal to improve solid transfer, to prevent gas mixing and obtain gas seal, the operation of DFB becomes further intricate and needs better operational understanding. Sung et al. [78] presented a map of possible operation condition of the lower loop seal and riser for DFB. A boundary of stable region for a set of velocities in the riser and loop-seal was obtained. Moreover, three unstable regions were observed:1. Due to low velocity in the lower loop-seal, no smooth solid circulation and sluggish motion in gasifier. 2. Due to high riser velocity, low solid holdup exists in the riser.

S. Shrestha et al. / Renewable and Sustainable Energy Reviews 53 (2016) 1529–1548

Fig. 6. Operational map at loop-seal air flow 8 m3/h [25].

3. Due to low riser velocity, low solid is transferred from the riser to the cyclone. To minimize the unstable region 1, loop-seal velocity can be adjusted to achieve smooth solid circulation rate whereas unstable region 2 and 3 need a proper handling of riser velocity. The total gas velocity in the riser should be such that the solid would be carried out from the riser to ensure the proper gas–solid mixing but cannot be so high that the solid elutriated could be accumulated in the standpipe but not recirculated, which results in the drop of solid density in the riser. Consequently, the failure of the system either by overloading of the loop-seal or by the fluctuation of pressure drop could occur leading to unsteady solid flow. In addition, the loop-seal airflow should also not be too high to deplete solid from the loop seal resulting in the bypass of gas from the BFB. Lim et al. [25] developed an operational map on a cold model of a circulating fluidized bed based on flow visualizations using solid mass flow rate and pressure drop measurements to differentiate between fluidization regimes. A stable operating region ‘F’ where solid moved freely up and out of the riser, down through the cyclone into the standpipe and back into the riser via the loop-seal bounded by six unstable regions was developed as shown in Fig. 6. Likewise, maximum operating limits for primary air, secondary air and inventory were obtained. Furthermore, Lim [26] presented an operational boundaries for the steady state operation of the DFB as a function of the primary airflow rate in the CFB riser and the BFB airflow rate. A stable and an unstable operating region were identified as shown in Fig. 7. Moreover, he determined that chute pressure drop must be sufficiently high to prevent chute airflow bypass from CFB to the BFB. Likewise, BFB airflow and chute airflow should also be sufficiently high so that chute becomes adequately filled with solid for stable operation. Region F’ in Figs. 5 and 6 is also suggested to be a stable region, but due to experimental limitations, it was defined as hypothetical stable region. The operational maps discussed above hold true for any DFB process to be utilized with the configuration used to determine the operational map. But to improve the operation and achieve the required performance, various designs have been put forward (refer Table 2). The operational maps might not be applicable to the configurations which employ double exit loopseal, cone valves, riser with wider bottom, etc., and setups operated under different operational conditions like dissimilar particle properties, fluidization velocities, inventory, etc. Nevertheless, the results showed that it might be helpful in determining the operational map or region of the stability in any setup listed in Table 2 with some improvement mainly for the CFB which is common regardless of the process DFB is operated. Moreover operational stability was achieved when the global solid

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Fig. 7. Operational Map for dual fluidized bed (Qsec ¼324 m3/h, Qls ¼ 5 m3/h, Qch ¼ 5 m3/h) [26].

flow rate1 was steady and the total pressure drop in both the reactors remained constant over the operating-period [79]. While operating the DFB, after achieving the dynamic balance, it is essential to know its hydrodynamic characteristics like pressure drop, solid fraction and solid circulation rate and their variation under different operational conditions. 5. Pressure profile In an atmospheric DFB as shown in Fig. 2, two pressure balance loops exist, Eqs. (9) and (10) [44]. Eq. (9) shows the overall pressure balance loop of the DFB whereas Eq. (10) represents the pressure balance loop along the cyclone, down-comer and loop-seal. P Out;ris  P Out;gas ¼ ΔP gas þ ΔP con  ΔP ris  ΔP cyc P Out;ris  P Out;gas ¼ ΔP dc  ΔP ls

ð9Þ ð10Þ

Eqs. (9) and (10) are not valid for all the configurations of DFB listed in Table 2 and it depends upon the loop implemented for the process. However, in DFB mostly two pressure loops exist within the system, one representing the overall balance which includes both reactors and the lower connection between them and the other also representing the connection between the two reactors at the upper section which includes CFB (combustor) pressure drop for DFBG, AR for CLC and regenerator for CaL along with cyclone’s pressure drop. Pressure balance is also used while empirical modeling of DFB as boundary conditions [44,45,80–82]. Likewise, the obtained pressure drops are used to interpret other hydrodynamic properties such as solid holdup, inventory, etc. [45,58,79]. Furthermore, pressure profiles helps to evaluate modifications done in cold model and confirm their usability [83]. Overall pressure profiles observed in previous studies on DFB for various purposes are shown in Fig. 8.,2 In the closed loop-system, increasing the solid circulation rate in the riser pushes the solids to the BFB, consequently increases the solid inventory in BFB, thereby pressure drop and results in a decrease in riser pressure drop which can also be understood from Eq. (8) [44]. Goo et al. [84] observed pressure profile to be nearly same except in the lower region of the riser and lower seal-pot when solid circulation rate was varied in the range of 30.5–43.5 kg/m2 s by aeration rate into the lower loop-seal (Fig. 8a). Similarly, Ramirez [85] observed higher pressure drop across the riser, especially at the bottom when the solid circulation flux was increased while pressure drop across the riser declined progressively, with riser velocity 1

Global solid flow rate: solid circulation rate between the two reactors. Not all data’s has been presented and reader are urged to see the references if needed. 2

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Fig. 8. (a) Pressure profile for DFBGs. (b) Pressure profile for CLC. (c) Pressure profile for CaL. (d) Pressure profile for AER/SER.

increasing (Fig. 8d3). Similar effect of riser velocity has also been observed by Seo et al. [86] (Fig. 8a) and Charitos et al. [45]. Increase in riser velocity decreases pressure drop in the riser as solid entrainment is increased. Similarly, increase in primary air in the riser decreases the pressure drop in dense zone as solids are transported to transport zone. This leads to an increase in pressure drop in transport zone as solids get redistributed, resulting in frictional losses from the inter-particle and particle-wall interactions [25]. Apart from the solid circulation rate and riser velocity, other parameters like inventory, BFB freeboard absolute pressure drop, and double exit loop-seal aeration have been found to increase pressure drop in the riser, since the increase in all these 3 (P–Pref) The pressure at every section of the loop is plotted against the height from riser distributor, with the pressure at the outlet of secondary cyclone as the reference pressure (Pref).

parameters lead to increase in the inventory of the riser whereas the increase in the mean particle size accumulates the inventory at the lower region of riser due to the increased inertia of the coarser particles, consequently larger pressure gradients are noted in the dense and lower part of the lean core annulus region compared to that of the upper lean core annulus region and exit region [45]. Fig. 8b shows pressure profiles obtained from studies on cold model related to CLC. Since studies on CLC utilizes CFB–CFB configuration, two exponential pressure profiles of the two reactors can be observed different than pressure profiles seen in other configurations with CFB–BFB. Differences in pressure profile due to the change in configuration can also be noticed in Fig. 8c, where Cotton et al. [83] used entrained flow carbonator with BFB gasifier for CaL. In addition, Fig. 8c also shows the pressure profile of the configuration in which cone valve has been employed to control the solid looping rate [45,87]. Along with the reactor configuration, connection and the non-mechanical valves employed in

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Fig. 9. (a) Solid fraction profile for riser in case of ASR. (b) Solid fraction profile for riser with varying riser velocity. (c) Solid fraction profile for riser with varying solid circulation rate.

between them also considerably influences the pressure profile. From Fig. 8 it can be seen that the highest pressure drop observed is around the non-mechanical valves and connection applied; it is obvious since these valves are operated under higher pressure drop to acquire better gas-seal as well as to avoid inverse gas flow. Influence on pressure drop in these components due to the change in operations is vital. Only a few studies were conducted in DFB for these purposes [80,86–88].

6. Solid fraction The solid fraction is the measure of fraction of volume occupied by the solid in a gas–solid suspension and also termed as solid holdups. In fluidized bed reactors solid holdups are generally determined with the measurement of differential pressure drop in the reactors and using the properties of gas and solid phases in accordance to Eq. (11) [75]. Differential pressure drop is

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proportional to the solid concentration if the friction loss between the particles and the column wall effects are neglected, since the pressure gradient acts as a hydrostatic pressure gradient. dP ¼  g½ρp ð1  εs ðzÞ þ ρg εs ðzÞ dz

ð11Þ

For gas–solid systems with, ρp c ρg , then Eq. (11) becomes

εs ðz Þ ¼ 1 þ

1 dP g ρp dz

ð12Þ

With the knowledge of particle density and pressure gradient, the average solid holdup (εs) along the height of the riser can be determined from Eq. (12). For some cases, other available measurement techniques might be needed to be used, refer to [75]. Since the evolved heat from combustion in CFB has to be absorbed by the heat carriers, a good gas–solid interaction in the CFB-riser is needed, reflecting the importance of solid volume fraction in the riser. Moreover, solid fraction provides information on the flow regime under different operating conditions as well as defines parameters such as mixing, heat and mass transfer and reaction performance [85]. On the basis of solid fraction profile, the riser is divided into three zones: dense zone, lean core-annulus zone and sometimes an exit zone in the cases of constricted exits. Higher amount of solid fraction is present in dense zone, whereas solid fraction decays exponentially along the height in the core-annulus zone. And the solids’ density tends to increase in the exit region compared to that at the top of the lean region if the exit is abrupt due to the solid back mixing, resulting in the traditional ‘C’ shaped [89] axial solid fraction profile. Solid fraction profile in the risers of the DFB obtained in previous studies is shown in Fig. 9. Although, the profile shows similar trend that has been obtained in the CFB for combustion purposes, the results obtained in DFB differs in various ways. Löffler et al. [72] used the diffuser to enhance solids recirculation within the CFB. In their investigation, higher solid density was observed at that region where diffuser was applied. Moreover, with the decrease in B/T ratio and increase in S/T ratio, the solid’ density increased in the dense zone followed by steeper decrease in solids’ concentration as gas velocity was lower in this part of the riser (Fig. 9a). Similar results were also obtained for solid fraction profile when air was staged in CFB by Bidwe et al. [79]. Solid fraction profile with and without air staging is shown for regenerator of CaL (Fig. 9a). With the help of air staging as well as diffuser, it might be possible to adjust the solid fraction if needed for a particular area, however, in general air staging is neglected and a single riser velocity is employed. Under these circumstances, with the increase in riser velocity solid holdup in the riser decreased due to the increase in particle entrainment [84,85,90,91]. As the velocity increases, the solid fraction decreases along the height, consequently decreasing the height of the dense zone. Further increase in the velocity can cause the transition from fast fluidization to pneumatic where the dense zone disappears (Fig. 9b). At the similar riser velocity, the axial distribution in the whole riser became uniform when operated under the low Gs while at the higher Gs three distinct regions were observed leading to increment in the solid density along the height of the riser. Furthermore, profound exit effect has been observed with an increase in Gs (Fig. 9c) [78,84,85]. Similar results were also obtained when riser velocity was varied, that is solid fraction increased as Gs increased, but the fluidization regime the riser being operated might be different leading to dissimilar solid fraction [92]. Other parameters have also been found to influence solid fraction. Increase in inventory, BFB freeboard absolute pressure drop and double exit loop-seal aeration increased solid fraction, since increase in all these parameters led to increase in inventory of riser [45]. Moreover, solid fraction was also found to increase in

Fig. 10. Axial voidage profile in DFB [80].

the bottom dense region and lower part of core-annulus region whereas it decreased in the upper core-annulus region and exit region when particle diameter was increased [45]. This was also proven by Karmakar and Datta [80]. They found that solids density at the bottom zone was more for larger particles than that of the smaller particles. Both smaller and larger particles were present at the bottom zone but finer particles were embedded in the larger diameter particles which increased the solid fraction. Above secondary air injection, the solid fraction decreased to 0.002 and in the fully-developed zone, solid fraction was more for smaller particle compared that of particles with larger diameter. Data on solid fraction along with all the components of the DFB is scarce. Fig. 10 shows the axial voidage of DFB using sand particles of dp ¼ 0.334 mm and riser velocity of 5.380 m/s. And the voidage along the riser has been indicated from point-1 to point-4. The voidage in down-comer, L-valve to bubbling bed, the bubbling bed, connector and the L-valve to fast bed riser are shown from point-7 to point-16 [80].

7. Solid circulation rate In the DFB process, the bed materials adapted are circulated within both reactors and expressed as solid circulation rate which is a vital hydrodynamic parameter. In DFBGs as shown in Fig. 2, the solid circulation rate between the gasifier and combustor provides the necessary heat for the gasification reaction whereas in CLC, solid circulation rate marks the rate of supply of metal oxide between reactors and the supply of bed materials in CaL as well as in AER. In DFBGs, at higher solid circulation rate, temperature difference between the two reactors is lowered, and higher solid flux conveys more char from gasifier to combustor, thereby it reduces the amount of additional fuel [44]. Dietrich et al. [93] discussed four different methods for measuring solid circulation rate in the scaled cold flow model of CFB, among which the first method have been mostly used because of its simplicity. In this method, fluidization in non-mechanical valve is abruptly turned off. Consequently, the bed material gets amassed in the down-comer and is no longer transported and concurrently the height of the accumulated material along with the time is measured. With the known values of bulk density of the fixed bed and the cross-sectional area, a rough approximation of the mass flux can be determined in accordance with Eq. (13): ṁ ¼

U Δz ṁ  ρb  Adc and Gs ¼ Ar Δt

ð13Þ

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Fig. 11. Specific solids circulation rate design chart of conFig.uration A. Parameters: total solid inventory (TSI) and bed mass distribution, riser velocity, air/fuel ratio. Gs, min gives the required solid flow from mass and energy balance [47].

Fig. 13. Effect of inventory.

Fig. 12. Effect of riser velocity for varying particle size and inventory.

Although solid circulation rate is an important parameter, little information is available in the open literature. Researches to investigate the behavior of solid circulation rate in DFB with operating variables such as gas velocity in DFB components, inventory and particle properties would guide researches in optimizing and implementing practical control methods. 7.1. Effect of gas velocity in riser, inventory and particle property It is an utmost requirement to know the influence of gas velocity, inventory and the particle properties on solid circulation rate. Since, these are the basic parameters that are known and their impact on the solid circulation rate bears huge significance. Kronberger et al. [47] presented a solid circulation rate design chart shown in Fig. 11 for CLC with experimental setup summarized in Table 2. Fig. 11 shows that the solid inventory and the velocity in the riser affect strongly the specific solid circulation rate. Also, a wide range of solid circulation rate was obtained. In their configuration, the highest achievable solid circulation rate was 90 kg/m2 s. However, in their study fluidization velocity in the fuel reactor and particle locks did not influence solid circulation rate as “overflow type” particle return system was used in fuel reactor. Effect of gas velocity in the riser of DFB for varying particle size and inventory is shown in Fig. 12. Karmakar and Datta [80] analyzed cold model of DFB system by using L-valve incorporated between the coupled reactors. They found that solid circulation

rate increases with aeration flow and the superficial velocity increasing. For the equal superficial air flow, solid circulation was higher for lower particle size. The increasing particle size subsequently increased the aeration requirement in L-valve. Investigation regarding the effect of particle size in aeration requirement in non-mechanical valve of CFB loop has also reported the similar trends. Conclusively, the aeration required to obtain the same solid flux increases while increasing particle size [92,94,95] and the maximum obtainable solid flow rate is higher for particles with higher density [95]. Moreover, attempts to correlate the solid circulation with other operating parameters and variables of the riser have also been carried out. Although it might be of little use if applied to experimental setups other than the setups in which correlation was determined as it is highly empirical, it aids in understanding the parameters and consequently leads to a better and applicable correlation. Charitos et al. [45] implemented a double exit loop-seal and a cone valve to control the sorbent looping rate between the two beds and correlated the riser entrainment with the riser velocity for each mean particle size and also with riser exit pressure drop as shown in Eqs. (14)–(16). Significantly, in their study, negligible effect of inventory was observed on solid circulation rate. However in the open literature, it has been found that at a given Ur, Gs increases with reactor total mass load increasing. Fig. 13 shows the effect of inventory. ; R2 ¼ 0:84 for dp ¼ 142 μm Gs entrainment ¼ 0:15 U 4:96 r

ð14Þ

; R2 ¼ 0:85 for dp ¼ 230 μm Gs entrainment ¼ 0:29 U 4:21 r

ð15Þ

2 Gs entrainment ¼ 2:65ΔP 0:87 exit region ; R ¼ 0:82

ð16Þ

7.2. Effect of air staging in riser Fluidization velocity in the riser is usually divided into primary and secondary airflow as shown in Fig. 2. Primary air supplied to the CFB riser is normally above the minimum fluidization velocity and below the terminal velocity of the single particle. In DFBGs, with the reduction in primary air flow, the by-pass of primary air into the BFB via the chute is avoided [72]. Furthermore, limiting the primary airflow creates a sub-stoichiometric region below the secondary air injectors, which reduces NOx emissions [96]. The secondary air flow aids the upward transport of solids into the

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Fig. 14. Effect of primary air for different inventory and secondary air flow rates [25].

Fig. 15. Effect of primary air percentage and secondary air ratio to total fluidization air.

cyclone, it thus reduces the required primary airflow. The secondary air is placed at some height above the distributor and usually above the transport velocity. With two airflow inlet, it is obvious that solid circulation rate is influenced to a certain extent. Lim et al. [25] observed that solid circulation increases with the increase of the primary air and the secondary air flow rates and solid inventory of the plant in a cold model of a CFB as shown in Fig. 14. Based on the geometry of the outlet and solids inertia, a semi-empirical model including aerodynamic factor (Фd) was developed to estimate the ratio of solid exiting the CFB riser to solid recirculating back as shown in Eqs. (17) and (18). Moreover, they identified primary and secondary air as key operational variables.
w ð17Þ ṁ ṁs ¼ Фd π D core    πD  1 exp  aϵs;core þ 1 ð18Þ Фd ¼ w The coefficient ‘a’ took at a value of 3668 which gives R2 a value of 0.8327. Although primary air and secondary air both significantly influence the solid circulation rate, studies suggested fluidization air at bottom or primary level should be more effective [42,58,79,97]. The solid mass flux per area decreases while the ratio of secondary air in the total fluidization air is increasing; meanwhile, an increase in the bottom air flow as well as primary air would increase the circulation rate. However, primary air flow must be limited as high primary air may lead to gas leakage and pressure fluctuations leading to unstable operation. Fig. 15 shows the effect of primary air percentage on left y-axis, whereas effect of secondary air to the total fluidization air ratio for different bottom air is given on right y-axis. 7.3. Effect of gas velocity in loop-seal Accompanied by riser gas velocity, the loop-seal in DFB plays a vital role in governing the solid circulation rate. Moreover, solid circulation rate is controlled by gas velocities of the riser and non-mechanical valve [84,98,99]. The available literature data on the effect of gas velocity on the loop-seal in DFB are presented in Fig. 16 (left-axis). Goo et al. [84] designed and constructed a cold model of DFB gasifier to determine hydrodynamic properties. They reported that solid circulation rate could be controlled by gas velocities to the lower loop-seal

Fig. 16. Effect of gas velocity to the loo-seal for different riser velocities.

and the riser. Maximum Gs (90 kg/m2 s) was obtained at riser velocity of 4 m/s and loop-seal aeration at 4 Umf. The amount of circulating flux increased sharply with the increase of Uls up to 2.5–3 Umf to reach the maximum capacity of the loop-seal as a function of gas velocities to the loop-seal and the riser. Also, Gs values were correlated with dimensionless terms initially proposed by Kehlenbeck et al. [46] and Monazam et al. [99] as shown in Eq. (19).

ṁs D U 2r ¼ 8:88 X 10  5 Iv U t gD

!0:529   U ls 0:459 Ut

ð19Þ

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Fig. 18. Effect of Effect of volumetric flow rate in BFB. Fig. 17. Effect of loop-seal aeration.

The increase in solid circulation rate with the riser velocity and loop-seal aeration increasing was also reported by Ramirez [85]. Moreover, solid circulation rate depends on the amount and the location of the aeration supplied to the loop-seal. Yang et al. [100] concluded that aeration location influenced the threshold aeration rate in loop-seal and air injection to the supply chamber was more effective for adjustment of solid circulation rate as shown in Fig. 17. Moreover, solid circulation rate increases with the aeration rate, total solid inventory and fluidizing gas velocity in the riser increasing. An empirical equation was proposed to correlate solid circulating rate with the operating parameters as presented in Eq. (20). Gs ¼ 9:6877 U 0:575 Q 1:020 Iv0:543 sc r

ð20Þ

Although supply chamber aeration rate has proven to affect the solid circulation rate, it might not useful in achieving the desired value [86,101]. To start the solid flow, the aeration in the recycle chamber should pass a critical value [92]. Increasing aeration, consequently, enhances the solid flow rate up to a certain limit. After reaching this limit solid flow rate remains constant. For a stable operation, the aeration in the recycle chamber should be operated above Umf of the largest particle [86,92,102,103]. Seo et al. [86] investigated solid circulation rate and solid flow characteristics in the loop-seal of a dual circulating fluidized bed reactor. They concluded that the recycle chamber aeration rate could provide a high and broad range of solid circulation rates. The experimental setups were rectangular in shape. The fluidization air was injected at three points in the loop-seal – supply chamber, recycle chamber and vertical aeration. Their results showed that when there was no aeration in recycle chamber of the loop seal, the increasing riser superficial velocity did not affect the solid circulation rate and remained constant at 20 kg/m2 s as shown in Fig. 16 (right y-axis). The increase in solid circulation rate due to the increase in the riser gas velocity and the recycle chamber aeration rate is shown in Figs. 16 and 17. Maximum value of solid circulation was observed when vertical aeration and aeration in supply chamber were maintained at the minimum fluidization. The optimum vertical aeration position of the loop-seal was found to be a height/base line (h/l) ratio 2.5. Nguyen et al. [91] investigated the effect of the gas velocity in the riser and in the recycle

chamber of the loop-seal on the solid circulation rate. Increase in the solid circulation rate with respect to loop-seal aeration was higher for riser velocity of 4 m/s than that for riser velocity of 3.5 m/s as shown in Fig. 17.,4 7.4. Effect of gas velocity in BFB Studies discussing the effect of gas velocity in BFB on solid circulation rate are rare. The literature suggested that the solid circulation rate in DFB did not depend on the volumetric flow rate in the BFB at relative higher volumetric flow rate [46,58], which is illustrate in Fig. 18.,5 In addition, modification in configuration in the way-out of solids from gasifier was also found to influence the solid circulation rate. BFB gas velocity above 2 Umf resulted in either increase or decrease in the solid circulation rate according to the type of way-out [78]. Moreover, the stable solid circulation could be achieved with BFB fluidizing air velocity 1–1.5 Umf [86].

8. Recommendations

 Using modified scaling parameters for the CFB might help in



minimizing the deviation while matching the scaling parameter. Therefore, these parameters should be verified and matched to check their suitability while developing a scale-down cold model. The obtained conclusion can aid in obtaining better similarity in fluidization performance between the lab-scale model and an industrial scale plant. The development of a specified regime maps facilitates the selection of the fluidization regime. In DFB, although developed regimes illustrate regimes of fluidization in these reactors, further research is needed to clarify the regimes of fluidization within these reactors.

4 For, Nguyen et al. hydraulic diameter used to calculate the volumetric flow rate of recycle chamber (4A/P). 5 For, Seo et al. hydraulic diameter used to calculate the volumetric flow rate of BFB (4A/P).

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 Operational map for the system represents the region of stable



operations and assists in achieving smooth operation. It also indicates the ranges of the operating variables. Operational maps developed are limited to the ranges of only few parameters. Operational maps relating dimension parameters like height and diameter, fluidization velocities, loop-seal aeration, BFB configuration, particle properties etc. are apparently necessary. Understanding the impact of operating variables on the hydrodynamic parameters like pressure drop, solid fraction and solid circulating rate enhances the possibility of regulating the reactor according to the operational requirements. Although, this paper summarizes the published results to show the effect of fluidization velocities, inventory, particle properties on hydrodynamic parameters. Most of the experiments are conducted in batch. Experiments conducted in continuous system are necessary so that the ranges and results could be explored in order to achieve the generality in optimizing the solid circulation rate.

9. Conclusion DFB technology is developing for processes like gasification, CLC, CaL and AER which primarily involve to remove CO2. Generally, CFMs are employed to study the hydrodynamics which aid toward the optimization of the industrial scale plant. Mostly, the scaling laws consisting a set of dimensionless parameters are used to construct the CFMs to govern the dynamic similarity between CFMs and industrial scale plant. Moreover, the regime maps developed aid designers in selecting the fluidization regime for the applied conditions. In DFB, the CFB riser operates in fast fluidization regime whereas gasifier in bubbling regime. Smooth solid circulation rate is difficult to achieve. Fluidization velocities are found to be the key operational variables governing the stability of the system. To achieve the stability in the system, the loop-seal must be sufficiently fluidized. Operational map can assist in the identification of the stable and unstable regions in the system. Pressure balance, solid fraction and solid circulation rate are the three important hydrodynamic parameters. Pressure balance in the DFB helps in empirical modeling, identifying suitability and modifications needed to be done in CFMs. Increasing the solid circulation rate increases the pressure at the bottom region of the riser and then progressively declines along the height whereas increasing riser velocity subsequently decreases the pressure drop. The non-mechanical valves and connection applied are operated under higher pressure drop to acquire better gas-seal as well as to avoid inverse gas flow. Besides, pressure balance, solid fraction in riser is another vital hydrodynamic parameter. Solid fraction profile of the riser indicates a dense zone followed by a core-annulus zone and an exit zone for riser with a constricted exit. Solid fraction was found to increase with increase of solid circulation rate whereas to decrease with the riser velocity increasing. Gas velocity in the riser and the loop-seal are two primary variables influencing the solid circulation rate. Solid circulation rate increases while increasing the gas velocity and the total solid inventory in the riser. The aeration required in the loop-seal increases with the solid circulation rate and particle size increasing. Further increment in the aeration enhances the solid circulation until it reaches optimum capacity of loop-seal. Moreover, aeration in the recycle chamber of the loop-seal plays a key role in regulating the solid circulation rate. Although, the gas velocity to the BFB has negligible influence in the solid circulation rate, proper fluidization is required for smooth solid circulation.

Acknowledgment This research is financially supported by University of Malaya, Ministry of Higher Education High Impact Research (UM.C/HIR/ MOHE/ENG/30).

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