Verification of fluidized bed electrical capacitance tomography measurements with a fibre optic probe

Previous studies aimed at determining the spatial accuracy of electrical capacitance tomography (ECT) have employed phantoms placed within the ECT measurement space. No previous studies have compared ECT with a second independent measurement technique in an operating fluidized bed. In the present work, radial voidage profiles have been measured with ECT in the 0.14-m I.D. riser of a circulating fluidized bed (CFB) and in a bubbling fluidized bed with a 0.19-m I.D. The dynamic and time-averaged radial voidage profiles have been compared with measurements taken with a fibre optic probe in the same riser and in a slightly narrower (0.15-m I.D.) bubbling fluidized bed. In spite of the intrusiveness of the latter technique, the time-averaged radial profiles in the CFB riser fall within 10% of each other when the CFB is operated at high-flux conditions that lead to a very dense wall region. Iterative reconstruction of the ECT images is not needed in this case. Similar agreement is found between the two techniques in the bubbling fluidized bed, but off-line iterative image reconstruction is clearly necessary in this fluidization regime. These results suggest that ECT, which is often described as a tomographic imaging technique with low spatial resolution, can in fact provide semi-quantitative time-averaged images of the cross-section of fluidized beds of diameter comparable to or less than that used here.     

60米高循环流化床内物料浓度分布的冷态试验

利用电厂循环流化床锅炉现有的结构和设备, 搭建提升管高度60m、内径400mm的超高循环流化床冷态实验台, 重点研究了流化风速和颗粒密度对提升管内轴向和径向空隙率分布的影响。实验结果表明:空隙率分布形式与流化风速和物料密度密切相关, 对于一定的床料高度, 在底部密相区一直有床料堆积的情况下, 随着流化风速的增加, 提升管底部密相区空隙率增大, 上部稀相区的空隙率减小并且其在径向的分布变得更加不均匀;在一定的流化风速下, 密度较小的物料将更多的被带入上部稀相区, 上部稀相区的空隙率减小, 其在径向分布将变得更加不均匀。     

湍动流化床过渡段固含率分布特征的实验及数值模拟

在湍动流化床中,过渡段对于包括甲醇制烯烃在内的气固催化快反应有着重要的作用。采用PV6D反射型光纤探针对内径95 mm的湍动流化床内过渡段的固含率分布和脉动参数进行了测量,分别考察了表观气速和静床高的影响,并采用修正的基于颗粒动力学的三段曳力双流体模型进行模拟。实验表明,湍动流化床过渡段中固含率的轴向分布呈现S型和指数型两种类型,固含率轴向与径向分布都在过渡段内出现最大梯度,表明过渡段中固体浓度分布比稀相段和密相段更不均匀。表观气速和静床高的变化将导致S型和指数型分布的相互转变,并且对过渡段底部与壁面附近的固体高浓度区影响最为显著。局部固含率脉动概率密度分布表明,在静床高较小时,随着气速的增大,床层下部气含率最大值位置将从中心区移动至环隙区,呈现气含率的双峰型分布。本文提出的修正三段曳力模型考虑了颗粒团聚的影响,对过渡段中分布板影响区之外的固含率分布均能较好地模拟。     

TRANSITION FROM TURBULENT TO FAST FLUIDIZATION

Based on eight transition criteria, at least two types of transition velocities are identified for the demarcation of the transition from turbulent to fast fluidization. The "critical velocity," U se , corresponds to the significant entrainment of particles from the bed, beyond which a circulating fluidized bed operation becomes essential. The "transport velocity," U tr , defines the transition to fast fluidization based on the axial solids concentration profiles. Below this velocity, a distinct interface exists between a top dilute region and a bottom dense region. Above U tr , the variation of voidage with height becomes relatively smooth. U tr is found to be a function of measurement location and riser height, as well as gas and particle properties.     

Transition from turbulent to fast fluidization

Based on eight transition criteria, at least two types of transition velocities are identified for the demarcation of the transition from turbulent to fast fluidization. The "critical velocity," U se , corresponds to the significant entrainment of particles from the bed, beyond which a circulating fluidized bed operation becomes essential. The "transport velocity," U tr , defines the transition to fast fluidization based on the axial solids concentration profiles. Below this velocity, a distinct interface exists between a top dilute region and a bottom dense region. Above U tr , the variation of voidage with height becomes relatively smooth. U tr is found to be a function of measurement location and riser height, as well as gas and particle properties.     

Regime classification of macroscopic gas—solid flow in a circulating fluidized bed riser

A conceptual flow regime diagram for a circulating fluidized bed riser is proposed, combining existing investigations with experimental data obtained under idealized conditions in which a fully independent control of gas velocity and solid circulation rate was conducted by use of a screw feeder for solid feed into the riser. The diagram classifies the flow state into five regimes by qualitative transition lines which describe the relationship between gas velocity and solid circulation rate. These regimes are particulate fluidization, bubbling fluidization, turbulent fluidization, dense-phase transport and dilute-phase transport. The diagram suggests that S-shaped bed-density distribution or dense/dilute region interface appears only at limited conditions in the bubbling and turbulent fluidization regimes. These experimental findings were generalized by further experiments in a conventional circulation system with a ball valve between the riser and the downcomer which permits changes in the solid circulation rate and the bed height in the downcomer. The experimental results showed that the bed height in the downcomer has no particular effect on the bed density distribution or the height of the dense/dilute region interface, but an appreciable effect on the lowest gas velocity to maintain steady solid circulation at a given rate. These results are consistent with the above diagram.     

Behaviour of gas-fluidized beds of fine powders part III. Effective thermal conductivity of a homogeneously expanded bed

The effective thermal conductivity of the dense phase in expanded non-bubbling fluidized beds has been studied between minimum fluidization and minimum bubbling for 9 gas—solid systems. Except close to minimum bubbling, the effective thermal conductivity is not a function of bed voidage and is most sensitive to the gas phase thermal conductivity. The experimental results have been compared with 11 packed bed correlations and with suitable modification three of these can be used to accurately predict the effective thermal conductivity of a non-bubbling fluidized bed.     

Flow patterns in high-velocity fluidized beds and pneumatic conveying

Four flow patterns are identified for gas-solids vertical upward flows. Homogeneous dilute phase flow is characterized by the absence of both radial and axial solids segregation. Heterogeneous dilute phase flow (also called core-annulus flow) is characterized by the absence of axial solids segregation, with solids carried upward in the core and travelling downward near the outer wall due to the formation of particle streamers. Collapsed flow with a lower dense region and an upper dilute region is also referred to as the fast fluidization regime. In this case, the flow structure in the upper dilute region is similar to that in heterogeneous dilute phase flow, while the lower dense region resembles that in a turbulent fluidized bed. Dense phase flow can be reached when the riser is completely occupied by a relatively dense suspension with little axial density variation and no net solids downflow near the riser wall. The transition from fast fluidization to dense phase flow is still not clearly defined.     

A consolidated flow regime map of upward gas fluidization

A new flow regime map, resulting from more fundamental studies on the hydrodynamics and new flow regimes, is proposed in response to more practical reclassifications of the existing regimes with the development of upward gas-solids fluidization systems. The previously reported flow regime maps and flow structures of some widely used fluidized beds are carefully examined. To better reflect the industrial applications, the fast fluidization regime is reclassified as high-density and low-density circulating fluidization regimes. A consolidated flow regime map is then proposed where the flow regimes of upward fluidization expand to include new types of fluidized beds such as circulating turbulent fluidized bed and high-density circulating fluidized bed. The proposed flow regime map consists of six flow regimes: bubbling, turbulent, circulating turbulent, high-density circulating and low-density circulating fluidization, and pneumatic transport. The transitions between the regimes are discussed with new correlations proposed for fluid catalytic cracking type particles. Analysis on the dominating phase in the different types of fluidized beds reveals the dynamic changeover from solids phase continuous in conventional low-velocity batch/“fixed” fluidization operations to gas phase continuous in high-velocity continuous/“moving” fluidization operations and provides more insights to the transitions between the flow regimes for industrial design and practice.     

Distributor effects near the bottom region of turbulent fluidized beds

The distributor plate effects on the hydrodynamic characteristics of turbulent fluidized beds are investigated by obtaining measurements of pressure and radial voidage profiles in a column diameter of 0.29 m with Group A particles using bubble bubble-cap or perforated plate distributors. Distributor pressure drop measurements between the two distributors are compared with the theoretical estimations while the influence of the mass inventory is studied. Comparison is established for the transition velocity from bubbling to turbulent regime, U_c, deduced from the pressure fluctuations in the bed using gauge pressure measurements. The effect of the distributor on the flow structure near the bottom region of the bed is studied using differential and gauge pressure transducers located at different axial positions along the bed. The radial voidage profile in the bed is also measured using optical fiber probes, which provide local measurements of the voidage at different heights above the distributor. The distributor plate has a significant effect on the bed hydrodynamics. Owing to the jetting caused by the perforated plate distributor, earlier onset of the transition to the turbulent fluidization flow regime was observed. Moreover, increased carry over for the perforated plate compared with the bubble caps has been confirmed. The results have highlighted the influence of the distributor plate on the fluidized bed hydrodynamics which has consequences in terms of comparing experimental and simulation results between different distributor plates.     

Microstructural aspects of the flow behaviour in a liquid—solids circulating fluidized bed

The local instantaneous and time‐average suspension densities were determined in a 76 mm diameter by 3 m tall liquid‐solids circulating fluidized bed riser using a fibre‐optic probe. Attempts were made to qualify the microflow structure through statistical analysis of the local bed voidage fluctuations obtained under different operating conditions for the first time. The results show that local microflow structure is uniform in the axial direction but non‐uniform in the radial direction with more flow fluctuation near the wall than in the core of the column for a given axial position. The standard deviation and intermittency index tend to increase with increasing solids circulating rates. Comparing with the gas—solids CFB, the liquid—solids CFB shows much more homogeneous flow structure in both the axial and radial microscopic flow behaviours. The microflow behaviours in the conventional liquid—solids fluidization, liquid—solids circulating fluidization and dilute‐phase liquid transport regimes are also characterized by examining the probability distribution and the intermittency index of the solids holdup.     

Evaporative liquid jets in gas–liquid–solid flow system

Some aspects of the fundamental characteristics of evaporative liquid jets in gas–liquid–solid flows are studied and some pertinent literature is reviewed. Specifically, two conditions for the solids concentration in the flow are considered, including the dilute phase condition as in pneumatic convey and the dense phase condition as in bubbling or turbulent fluidized beds. Comparisons of the fundamental behavior are made of the gas–solid flow with dispersed non-evaporative as well as with evaporative liquids.For dilute phase conditions, experiments and analyses are conducted to examine the individual phase motion and boundaries of the evaporative region and the jet. Effects of the solids loading and heat capacity, system temperature, gas flow velocity and liquid injection angle on the jet behavior in gas and gas–solid flows are discussed. For dense phase conditions, experiments are conducted to examine the minimum fluidization velocity and solids distribution across the bed under various gases and liquid flow velocities. The electric capacitance tomography is developed for the first time for three-phase real time imaging of the dense gas–solid flow with evaporative liquid jets. The images reflect significantly varied bubbling phenomenon compared to those in gas–solid fluidized beds without evaporative liquid jets.     

鼓泡流化床离散模拟中的一个局部空隙率模型

气固流化床离散颗粒模拟中通常采用面积加权平均法计算局部空隙率,不能较好地反映流化床中显著的非均匀结构特性.为了考虑非均匀结构对局部空隙率的影响,文中提出一个适用鼓泡流化床的局部空隙率模型.将流场划分为稀区(气泡区)和密区(乳相区);非均匀结构对密区局部空隙率的影响通过引入局部空隙率下限和非均匀影响系数描述;将提升管流动...     

基于两相两区模型的气固流化床DEM模拟

气固流化床DEM模拟中,通常采用面积加权平均法计算局部空隙率,为了考虑网格中显著非均匀结构对局部空隙率的影响,提出一个计算局部空隙率的两相两区模型。该模型将网格中的非均匀结构虚拟划分为稀相和密相,将实际网格区域划分为稀区和密区,并采用时空关联性原理识别稀区和密区;模型还采用自适应方法计算网格中颗粒分布的非均匀度;将非均匀度作为非均匀结构的影响权重计算密区颗粒的局部空隙率。模拟了鼓泡流化床,模拟结果表明:与传统的DEM相比,基于两相两区模型的DEM能更好地模拟气泡形态,且能捕捉气泡冒出床层和气泡破裂的复杂现象。     

Effects of temperature on local two-phase flow structure in bubbling and turbulent fluidized beds of FCC particles

A novel high temperature optical fiber probe has been developed to study the effects of bed temperature on the local two-phase flow structure in a pilot scale fluidized bed of the FCC particles with bed temperatures ranging from 25°C to 420°C, covering both the bubbling and turbulent fluidization regimes. The results show that fluidization is enhanced and fluctuations of the local two-phase flow structure become more intense with increasing bed temperature. At constant superficial gas velocities, the averaged local particle concentration, the dense phase fraction and particle concentration in the dense phase decrease with increasing bed temperature, whereas both the frequency of the dilute/dense phase cycle and the ratio of the dilute phase duration to the dense phase duration increase. In addition, the effects of temperature on the dilute phase depend on superficial gas velocity. The conventional two-phase models fail to predict these changes of the local flow structure with temperature, which may be explained by the fact that the role of interparticle forces is neglected at different bed temperatures. Indeed, fluidization behaviors of the FCC particles tested increasingly shift from typical Geldart A towards B with increasing temperature due to a decrease of the interparticle attractive forces and a simultaneous increase of interparticle repulsive forces.     

Hydrodynamic study of gas and solid flow through a screen-packing

Many drawbacks of conventional gas fluidized beds can be avoided by filling the column with a fixed wire-screen Raschig ring packing of high voidage. Such systems have been studied in batch and countercurrent continuous mode. Excellent fluidization occurs with relatively coarse particles but not with smaller particles. In batch systems, the expansion is given by a Capes and McIlhinney type power law and the contribution of the packing to support the bed can be up 20 percent of the weight of solid. In countercurrent systems, for each gas velocity two operating points can occur: the dilute or the dense phase bed in the same way as for gas liquid flow. All increases in the flow of the solid or in the dilution of the bed lead the packing to support a greater part of the grains.     

EFFECTS OF PARTICLE CHARACTERISTICS ON THE PERFORMANCE OF A FAST FLUIDIZED BED REACTOR

A heterogeneous model for the fast fluidized bed reactor which carries out a gas-solid non catalytic reaction is presented. The hydrodynamics of the fast fluidized bed is characterized by the model of Kwauk et al. (1985) which assumes the existence of two phases; a dense phase and a dilute pneumatic transport phase. For a given solid flowrate, the length of the reactor occupied by each phase depends on gas velocity, particle diameter and density and average voidage within the reactor. The gas-solid reaction is assumed to follow the shrinking core model. The solids are assumed to be completely backmixed in the dense phase and move in plug How in the dilute pneumatic transport phase. The gas phase is assumed to be in plug flow in both phases

For given gas and solid flowrates, the transition from the dense phase flow to the fast fluidized bed (containing two regions) as functions of particle size and density is determined using the model of Kwauk et al. (1985). The numerical solution of the governing mass balance equations show that for given solid and gas flowrates, (and average voidage) the gas phase conversion shows an unusual behavior with respect to particle diameter and density. Such behavior is resulted from the effects of particle diameter and density on the reactor volume occupied by each phase and the effect of particle diameter on the apparent reaction rate. The numerical results show that a fast fluidized bed gives the best conversion at large particle density and for the particle diameter which results the fast fluidized bed to be operated near the pure dense phase flow.     

液固循环流化床的研究(Ⅰ)相含率及颗粒循环速率

在内径为140mm、高为3m的有机玻璃设备中对液固并流向上的循环流化床中两相流动特性进行了研究.相含率的研究表明,在循环流态化条件下,全床相含率轴向均匀分布;而径向具有显著的不均匀性,表现为中心区液相含率高,在r/R=0.7处液合率最低.颗粒循环速率的研究表明,在循环流态化区域,在一定的二次水流速下,颗粒循环速率不随总液速的改变而改变.由于在循环流态化下存在床层径向参数的不均匀性,因此床层的相合率与颗粒循环速率的关系与广义流态化方法预测的结果间存在差异,即在操作条件一定的情况下,床层内真实液合率比由广义流态化预测的低.     

Tomographic imaging of a conical fluidized bed of dry pharmaceutical granule

Hydrodynamics in a conical fluidized bed were studied using electrical capacitance tomography (ECT) for a bimodal and mono-disperse particle size distribution (PSD) of dry pharmaceutical granule. The bimodal PSD exhibited a continuous distribution with modes at 168 and 1288 μm and contained approximately 46% Geldart A, 32% Geldart B and 22% Geldart D particles by mass. The mono-disperse PSD had a mean particle size of 237 μm and contained approximately 71% Geldart A, 27% Geldart B, and 2% Geldart C particles by mass. The granule particle density was 830 kg/m³. Experiments were conducted at a static bed height of 0.16 m for gas superficial velocities ranging from 0.25 to 2.50 m/s for the mono-disperse PSD, and from 0.50 to 3.00 m/s for the bimodal PSD. These gas velocities covered both the bubbling and turbulent fluidization regimes. An ‘M’-shaped time-averaged radial voidage profile appeared upon transition from bubbling to turbulent fluidization. The ‘M’-shaped voidage profile was characterized by a dense region near the wall of the fluidized bed with decreasing solids concentration towards the centre. An increased solids concentration was observed in the middle of the bed. Frame-by-frame analysis of the images showed two predominant bubble types: spherical bubbles with particle penetration in the nose which created a core of particles that extended into, but not through, the bubble; and spherical bubbles. Penetrated bubbles, responsible for the ‘M’ profile, were a precursor to bubble splitting; which became increasingly prevalent in the turbulent regime.     

Identification of the flow structures and regime transition in gas–solid fluidized beds through moment analysis

Using statistic parameters of solids holdup signals, a moment consistency data processing method (MCDPM) was proposed. Experiments were carried out using FCC particles of 76 μm under different operating conditions, and MCDPM was used to successfully obtain solids holdups of the dense and dilute phases and the phase fractions over five fluidization regimes, bubbling (BFB), turbulent (TFB), circulating turbulent (CTFB), high‐density circulating (HDCFB), and circulating (CFB) fluidized bed systems. In BFB, TFB, and CTFB regimes, only dense phase fraction decreased with increasing air velocity, while the transition from HDCFB to CFB experienced appreciable change in the solids holdup of the dense phase. From the low‐velocity to the high‐velocity regimes, both the solids holdup and the fraction of the dense phase experienced a drastic decrease, suggesting that this transition corresponded to a profound change in flow structure and further suggesting that CTFB is in reality still a TFB. © 2012 American Institute of Chemical Engineers AIChE J, 59: 1479–1490, 2013