Flow regime dense

The most notable difference in behaviour as Gj is increased beyond 200-300 kg/m s at relatively high /g is the gradual disappearance of net solids downflow at the wall (e.g., Issangya et al., 2000 Karri and Knowlton, 1998, 1999 Liu, 2001). This leads to a different flow regime (dense suspension upflow, as discussed in Sec. 3) and a more homogeneous flow structure. Typical axial solids holdup profiles appear in Fig. 19. Note that at these high values of Gj, the profiles tend to reach relatively constant holdups of the order of 15%, indicating an approach to fully developed flow. Parssinen et al. (2001) found that there were four zones in axial profiles for a much taller riser. 

Figure 20 Predicted flow regime diagram of the industrial MIP reactor, with solids flux as a function of the imposed total pressure drop at fixed gas flow rate. The snapshots of voidage profile refer to the transition, from left to right, the dilute transport, choking transition in between with different solids inventory, to the dense fluidization (Lu et al., 2007).

Figure 9.3. Various flow regimes or patterns in dense-phase fluidization (a) Particulate fluidization (b) Bubbling fluidization (c) Turbulent fluidization (d) Slugging (e) Spouting (f) Channeling.

Cai, P. (1989). Flow Regime Transition in Dense-Phase Fluidized Beds. Ph.D. Dissertation. 

Pneumatic conveying systems can be classified on the basis of the angle of inclination of pipelines, operational modes (i.e., negative- or positive-pressure operation), and flow characteristics (i.e., dilute or dense phase transport steady or unsteady transport). A practical pneumatic conveying system is often composed of several vertical, horizontal, and inclined pipelines. Multiple flow regimes may coexist in a given operational system. 

Industrial reactors generally operate at very high velocities (of order 1 m/s) much in excess of terminal falling velocity for at least the finest powder fractions. Powder is continually elutriated and returned to the bed via cyclones. Under these conditions there is disagreement as to whether or not bubbles retain their identity and such beds have been described as "turbulent" or "fast fluidised". What little evidence there is supports the continued existence of bubbles but now in a much disturbed or heterogeneous dense phase and with a less definite shape. Until more is known about this physical situation it is not easy to see how the bubbling bed reactor models should be modified correctly to describe this flow regime. 

Cai, P. The transition of flow regime in dense phase gas-solid fluidized bed, Ph. D. Thesis, Tsinghua University (China) and Ohio State University (U.S.A.) (1989). 

Especially for multiphase systems flow visualization (Wen-Jei Yang, 1989 Merzkirch, 1987) can provide valuable initial information on the prevailing flow patterns and should at least always be considered as a first step. Of course, in applications that involve extreme conditions such as high temperature and/or pressure it is very difficult if not impossible to apply flow visualization and other techniques should be considered. Here the use of cold flow models which permit visual observation might be considered as an alternative as an important first step to obtain (qualitative) information on the flow regime and associated flow pattern. Of course, multiphase flows exist such as dense gas-solid flows that do not permit visual observation and in such cases the application of idealized flow geometries should be considered. A well-known example in this respect is the application of so-called 2D gas fluidized beds to study gas bubble behavior (Rowe, 1971). 

Fig. 2 Various flow regimes or patterns in dense-phase fluidization (A) particulate fluidization (B) bubbling fluidization (C) turbulent fluidization (D) slugging (E) spouting (F) channeling (G) fast fluidization. (From Ref...

The first term on the RHS in the viscosity closure denotes the kinetic contribution and dominates in the dilute regime. The second term on the RHS denotes the collisional contribution and dominates in the dense flow regime. 

Most gas-solid systems experience a range of flow regimes as the gas velocity is increased. Several important gas-solid fluidization regimes for the chemical process industry are sketched in Fig 10.1. In dense fluidized beds regions of low solid density may be created. These gas pockets or voids are frequently referred to as bubbles. 

Lean phase fluidization As the gas flow rate increases beyond the point corresponding to the disappearance of bubbles, a drastic increase in the entrainment rate of the particles occurs such that a continuous feeding of particles into the fluidized bed is required to maintain a steady solid flow. Fluidization at this state, in contrast to dense-phase fluidization, is generally denoted lean phase fluidization. Lean phase fluidization encompasses two flow regimes, these are the fast fluidization and dilute transport regimes. 

The fluidized bed reactors can roughly be divided into two main groups in accordance with the operating flow regimes employed. These two categories are named the dense phase and lean phase fluidized beds. 

However, it is not always easy to distinguish between the flow behavior encountered in the fast fluidization and the transport bed reactors [56]. The transport reactors are sometimes called dilute riser (transport) reactors because they are operated at very low solids mass fluxes. The dense riser transport reactors are operated in the fast fluidization regime with higher solids mass fluxes. The transition between these two flow regimes appears to be gradual rather than abrupt. However, fast fluidization generally applies to a higher overall suspension density (typically 2 to 15% by volume solids) and to a situation wherein the flow continues to develop over virtually the entire... 

Figures 14 and 15 show the normalized pressure drop factor for a densely packed bed of monosized spherical particles. For Rem < 7,fv is fairly independent of Rern, and at high Rem values, it increases fairly linearly with Rem. The data points are the experimental results taken from Fand et al. (110), where the bed diameter is D = 86.6 mm and the particle diameter is ds = 3.072 mm. One can observe that the 2-dimen-sional model of Liu et al. (32), referred to as equation 107, agrees with the experimental data fairly well in the whole range of the modified Reynolds number. From Figure 14, one observes a smooth transition from the Darcy s flow to Forchheirner flow regime. The one-dimensional model of Liu et al. (32) (i.e., equation 106) showed only slightly smaller fv value. Hence, the no-slip effect or two-dimensional effect for this bed is small. As shown in Figures 14 and 15, the Ergun equation consistently underpredicts the pressure drop. The deviation becomes larger when flow rate is increased.

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