Chemrcal Engineerirrg Science, Vol. 45, No. 8, pp. 1953-1966, 1990. Printed in Great Britain.

ooo9~2509/90 $3.00+0.00 0 1990 Pergamon Press plc

There is Increasing interest in high-velocity fluidized bed reactors operated in the turbulent and fast fluidization regimes. Understanding of the hydrodynamics of these fluldization regimes has improved greatly in recent years, and there are prospects for applications beyond those- practiced in industry.at this time. Reactors operated in these regimes offer some unique features for gas-solid contacting. However, considerable work is required to achieve a better understanding of these high-velocity systems and to allow them to be optimized with respect to reactor geometry and operating conditions.

HIGH-VELOCITY FLUIDIZED BED REACTORS

J.R. GRACE

Department of Chemical Engineering, University of British Columbia, Vancouver, Canada V6T LWS

ABSTRACT

KEYWORDS

Fluidization; turbulent fluidized beds; circulating fluidized beds; combustion; catalytic cracking.

INTRODUCTION

Pluidieed beds have found application in a number of processes involving gas-solid and solid-catalysed gas-phase reactions. Successful processes have included catalytic cracking. acrylonitrile manufacture, ore roasting, polyethylene production, calcination operations and combustion of a wide variety of fuels (Yerushalmi, 1982). Advantages of fluidized beds for these operations have included such factors as the capability of operating with small catalyst particles and hence high effectiveness factors, favourable bed-to-immersed-surface heat transfer coefficients, ability to withdraw and add particulate solids continuously, and the possibility of operation on a very large scale.

Many of the earliest (1940's) operations of fluidized beds tended to be at relatively high gas velocities, of the order of 1.5 m/s or more. (See Jahnig et al, 1980; Squires et al, 1985; Avidan et al. 1989 for discussions of these early units-) However, technical difficulties led to a decrease in superficial gas velocities in the early years. At the same time, the growing enterprise of academic fluidization research which began to flourish in the 1950's and 1960's became intrigued by the study of bubbles and slugs in fluidized beds. While this led to significant advances in the understanding of processes carried out at relatively low gas velocities, high-velocity processes and hydrodynamics were all but ignored. A number of rather simple two-phase models based on the concepts of bubbling beds were proposed and tested for low-velocity beds. (For reviews, see Grace (1971), Yates (1983) and Grace (1986a)). These models have had moderate success in predicting the principal features of experimental small-scale fluidized bed reactors operated in the bubbling and slugging fluidization regimes.

Recent years have seen a resurgence of interest in fluidized beds operated at high gas velocities in hydrodynamic regimes beyond the slugging and slugging regimes. In part, this is because industrial fluid bed reactors for sc~me processes (e.g. ore roasting, acrylonitrile manufacture) have always tended to operate at such high velocities. In part also, the interest has been kindled by the development of new "circulating bed" processes involving gas superficial velocities of 5 to 10 m/s and substantial external recycle of entrained solids. Moreover, academic institutions have gained the experience and courage to construct and operate larger units. with the requisite efficient particle collection devices, capable of exploring the high velocity regimes. While substantial progress has been achieved, there remains considerable uncertainty regarding the behaviour, modelling and control of high-velocity beds, as we will see. Fluidized beds, while they have had some notable successes, continue to be plagued by the idea in many quarters that they are extremely difficult to scale up, subject to severe operating problems, and unreliable in practice. The object of this paper Is to indicate areas where progress in the understanding of high-velocity fluid bed reactors has been made, and to delineate other aspects where considerable work remains to be performed.

1953

1954 J. R. GRACE

INCREASING U, e b

P2

FIXED BE -_ OR DELAYED REGIME BUBBLING

JRBULENT REGIME FLUIDIZATION

Fig. 1. Schematic showing key features of the various hydrodynamic regimes of fluidization.

TURBULENT FLUIDIZATION: HYDRODYNAMICS

The various hydrodynamic regimes of fluidization are shown schematically in Fig. 1. The recognition of the hydrodynamic regime now called "turbulent fluidization" is commonly attributed to Lanneau (1960) or to Kehoe and Davidson (1971). although Zens (1949) clearly viewed this regime many years earlier. There has been considerable controversy over the turbulent regime, Rhodes and Geldart (1986) even arguing that it is not a regime at all, but an artifact of differential pressure fluctuation measurements obtained where the inventory of solid particles is insufficient to cover the top pressure tap as the gas velocity is increased. Leaving this contention aside, the most frequent observations regarding the turbulent regime are as follows:

Differential pressure fluctuations tend to be of higher frequency and significantly lower amplitude than for the bubbling and slugging regimes (e.g. see Yerushalmi and Avidan. 1985). A similar decrease in heterogeneity can also be detected with other techniques such as capacitance probes.

There tend to be two fundamentally different descriptions of the basic structure of the two-phase flow. In one (e.g. Kehoe and Davidson), 1971). the flow is described visually in terms of transitory voids darting obliquely upwards, with the "turbulence" akin to random fluctuations in single-phase flows. These voids are smaller and much more regular than the bubbles and slugs viewed in the lower-velocity regimes of fluidieation. An alternate description (e.g. Rowe et al. 1980; Brereton. 1987) suggests that the turbulent bed Is composed of periods when there are slug-like structures, interspersed with periods when there are clusters and dilute zones typical of the higher velocity fast-fluidization regime. This description involves intermittency, with the fast fluidization characteristics becoming more and more prevalent as the gas velocity is increased. Further research is clearly required to Identify whether or not there are truly two different modes of turbulent bed operation, and whether traditional methods of characterizing turbulence in single-phase flows can be useful in describing turbulent f luidized beds.

The transition to the regime, even when the first of the above descriptions appears to apply, is usually gradual rather than sharp.

There appears to be a significant influence of the scale of the reactor on the onset of turbulent regime, with experimental evidence (e.g. Thiel and Potter, 1977; Sun and Chen. 1989) suggesting that the transition occurs at lower gas velocities as the vessel diameter is increased.

P2 High-velocity fluidized bed reactors 1955

The turbulent regime is transitional between bubbling or slugging on the one hand, where there is a dense "continuous" phase composed of a gas-solid emulsion, and the fast fluidization regime on the other hand, where the continuous phase is now the more dilute phase (Grace and Tuot, 1979). In this respect, as well as in general appearance, the turbulent regime shows analogous appearance to the churn-turbulent regime commonly identified for two-phase (gas-liquid) flows (Grace, 1986b).

There are pronounced radial gradients for beds operated in the turbulent regime (Abed. 1983), with a marked tendency for solids to be present fn much greater concentrations in the wall region, while the core of the column has a significantly smaller volume fraction of particles.

There is a significant influence of particle size distribution. with wide size distributions appearing to trigger both earlier transition to the turbulent regime and giving an increased concentration of particles in the dilute phase (Sun and Grace, 1990).

The presence of horizontal heat transfer tubes or other internal surfaces, by causing increases in local gas velocity, can trigger regions of local turbulent fluidisation in beds that would otherwise operate solely in the bubbling or slugging flow regime (Staub and Canada, 1978).

Data on the onset of turbulent fluidization have been summarized by Sun and Chen (1989). For most cases, transition appears to occur when the overall void fraction is of the order of 0.7 to 0.8 and for a value of U-Umf of 0.4 to 0.6 m/s; the regime then extends to an overall voidage of the order of 0.9, where fast fluidiaation takes over. The practical operating range corresponding to the turbulent regime is shown in the regime map which appears in Fig. 2. Note that beds can operate in this regime over very broad ranges of particle and gas properties and for superficial gas velocities above or below that corresponding to the terminal settling velocity of the particles.

t’ I “I I , ll,111, I I I ,,,,,

DILUTE CONVEYING

1 d; = Ar”3 = d;&,pg,p*]‘”

lo*

Fig. 2. FLuidization regime map adopted from Grace (1986b), with dimensionless superficial gas velocity plotted vs. dimensionless particle diameter. Practical operations of the various hydrodynamic regimes are indicated by the various wedge-shaped regions. Approximate boundaries between the different powder groups proposed by Geldart (1973) are shown at the bottom.

1956 J.R.GRAcE

5 k’

10

P2

5

Fig. 3. Comparison of ozone decomposition conversion in three hydrodynamic regimes with identical FCC catalysts and 102 mm dia. column. Open symbols: wide size distribution catalyst; closed symbols: narrow size distribution catalyst. Circles: bubbling regime; triangles: turbulent regime; hexagons: fast fluidieation. PF and PM denote single-phase plug flow and perfect mixing respectively.

DLSPERSION, CONTACTING AND REACTION IN TURBULENT BEDS

A number of commercial fluid bed reactors, including ones used for such reactions as sulphide ore roasting, Fischer-Tropsch synthesis and acrylonitrile manufacture, routinely operate in the turbulent regime of fluidization. This regime offers a number of advantages over the better-known lower velocity regimes: heat transfer coefficients between the bed and immersed or containing surfaces reach their maximum values within this regime; rapid exchange of gas between the fine and transient voids also appears to minimize gas bypassing and promote intimate gas-solid contacting; while some axial dispersion of both gas and solids does. occur, backmixing appears to be Less extensive and more random than at somewhat lower gas velocities. At the same time. the turbulent regime retains the advantages of temperature uniformity and solids handling capability which are critical to applications at lower gas velocities.

Most academic studies have stopped short of the turbulent regime, i.e. the gas velocities have only been sufficient for bubbling and/or slugging fluidization. However, early experimental work by Massimilla (1973) demonstrates clearly that turbulent fluidization is propitious for gas-solid contacting. Some recent data from our laboratory (Sun and Grace, 1990) plotted in Fig. 3 show the importance of the hydrodynamic regime and the benefits of operating in the high-velocity regimes. Direct comparison of the regimes was facilitated in this study by activating the same catalysts (several different size distributions) to different extents, and then adjusting the gas velocity and bed depth to ensure equal value of the dimensionless first order (ozone decomposition) rate constant or Damkohler number, k'-kH,*/U. All results were also obtained in the same 102 mm diameter reactor and with the same analysis system to enable direct comparisons. As shown in Fig. 3, the turbulent regime gives a significantly higher conversion than the bubbling or slugging regime, especially for the wide particle size distribution, which is more favourable than the narrow size distribution. The improvement is almost certainly due to improved gas-solid contacting within the turbulent regime, since this is the factor to which conversfon is most sensitive for reasonably fast reactions.

Modelling of the turbulent regime of fluidization is made difficult by the lack of suitable descriptors (bubbles, slugs, velocity regimes.

clusters) which have been helpful in modelling of lower and higher The simplest approach. suggested by Van Swaaij (1978) and Wen (1979). has been

to treat the reactor as a single-phase (homogenous) plug-flow reactor. However, it is clear from Fig. 3 that while turbulent beds give improved contacting, plug-flow limit.

they do not attain the single-phase At least part of the discrepancy may be due to non-ideal mixing behaviour.

Avidan (1982) and Wen (1984) suggested that a single-phase axially dispersed plug-flow model might be appropriate. dispersion coefficients

Experimental result? (Avidan, 1982) suggest that effective gas phase axial are of order 0.5 m /s for turbulent fluidization of FCC-like materials.

In later work, Avidan and Edwards (1986) suggested that the Peclet number, Pe=UH/D,x, is of order 5 to 10 for a 0.6 m diameter demonstration plant operating in the turbulent regime. Results summarized by Van Deemter (1980) suggest that the particle size distribution may have a major influence on the axial mixing of both gas and solids.

The radial gradients cited earlier, the occurrence of voids (albeit small in scale and transitory in nature), and the intermittent passage of slug-like structures suggest that more complex two-phase models may be required to describe turbulent beds. Bolthrunis (1989) recently dismissed bubbling bed models for beds that might be operating in the turbulent regime. In doing so, he failed to recognize that these models provide generic two-phase models. whose parameters can be readily adjusted to test regimes other than those for which they were originally intended.

P2 High-velocity fluidized bed reactors 1957

OR REGION

t .= 0 h @=fraction of I’” total solids Y

’ I

Fig. 4. Schematic showing a generic two-phase or two-region model which reduces to various literature models for specific values of the four parameters @, D2, Kc and $.

It is not difficult to revise two-phase fluid bed reactor models, originally intended for the bubbling bed regime, so that they are now intended to apply to other regimes of fluidization (Grace, 1986c). In common with most models for catalytic processes in fluid bed reactors, it will usually be satisfactory to ignore gas-particle mass and heat transfer resistances within each of the individual phases or regions. The particles will also usually be small enough that intraparticle concentratfon and temperature gradients can be ignored, and the temperature is commonly sufficiently uniform that temperature gradients can be neglected both inside and between the phases or regions. As with other chemical reactors, the sensitivity to the modelling approach will increase as the order of reaction fncreases. as the required conversion increases, and when dealfng with selectivity for mxe complex reactions.

Figure 4 shows a schematic representation of a generalized, but simple, two-phase (or two-region) fluidized bed reactor model where each phase or region is treated as a one-dimensional zone. The model requires specification of four parameters:

(a) the fraction, p, of gas which flows through the more dilute of the two phases or regions; (b) an axial dispersion coefficient, D

Z' for gas within the denser of the two phases or

regions. plug flow being assumed w thin the more dilute phase or region; (c) an interphase mass transfer (or crossflow) coefEicient, Kc, which regulates interchange

between the two phases or regions; (d) the mass fraction of solid particles, ,p. which are inside the more dilute phase or region.

As shown in Table 1, this simple generalized model not only encompasses, as special cases. many of the best-known bubbling bed reactor models. but it also covers. again as limiting cases, a well-known slugging bed model and the simple turbulent bed models that have been cited. Although the hydrodynamics of turbulent beds certainly differ in important respects from those of bubbling beds, at the stage of abstraction commonly adopted by the modeller, it is not clear that fundamentally different approaches are required.

Stringer (1989) has recently suggested that fluidized beds may behave as chaotic dynamic systems. Application oE chaos theory would appear to be most promising in the turbulent regime of fluidization. It is clear that considerable work remains to be done both to characterize and to model fluidized beds operated in the turbulent regime.

Table 1: Parameters required for the generic two-phase or two-region model shown in Fig. 3 to make it reduce to models Ear various hydrodynamic flow regimes of fluidization.

Model Regime B D2 K - - c Q

-

May 1959 Bubbling Determined Parameter 0 Orcutt et al 1962 Bubbling

l-U/Umf Predicted 0

Orcutt et al 1962 Bubbling l-U/Umf m l-U/Umf 0 Predicted 0

Hovmand & Davidson Slugging 0 Predicted 0 1968

L-U/Umf

van Swaaij 1978 Turbulent 1 0 1 Avidan 1982 Turbulent 1 Determined m 1 Grace 1984 Bubbling 1 0 Correlation Parameter Brereton et al 1988 Fast Fl'n 1 0 Parameter Parameter

P2 1958 J. R. GRACE

PAST F'LUIDLZATION: OCCURRENCE. CONFIGURATION AND APPLICATIONS

When the superficial gas velocity is increased within the turbulent fluidization regime, the overall bed voidage increases and the top surface between the "dense bed" and the freeboard becomes less and less distinct. At the same time, the rate of entrainment from the top of the vessel increases so that it becomes essential to feed fresh solids (or, more usually, to capture and return the entrained solids) if the inventory of particles in the system is to be maintained. The fast fluidization regime can be considered to be initiated when there is no longer a clear interface between a dense bed and a more dilute freeboard region. Instead, there is a continuous, usually gradual, decrease in solids content over most or all of the height of the column.

It is important to distinguish between "fast fluidization" and the flow behaviour encountered in riser or transported bed reactors. The former 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 height of the reactor, whereas the flow usually associated with transported bed reactors tends to be more dilute (typically 1 to 5% by volume solids) and more nearly fully developed. By virtue of the greater reflux and density of the suspension within the fast fluidized regime, there are also likely to be smaller temperature gradients than within more dilute pneumatic conveying. Nevertheless, these differences are often small in practice, so that there is significant overlap between the two types of reactor. This overlap is also apparent in the flow regime map shown in Fig. 2.

There have been relatively few studies which have unambiguously reported the transition to the fast fluidization regime, and there appear to be no general approaches capable of predicting this transition accurately for a wide range of conditions and systems. As with the transition to turbulent fluidization discussed above, the transition to fast fluidization appears to be generally gradual rather than abrupt. As shown in Fig. 2, practical operating systems correspond principally to Geldart type A and B powders. For Geldart group A powders, the superficial gas velocity for transition is of the order of 1.5 to 2 m/s, with somewhat higher velocities required for group B solids. The useful range of gas velocities for this regime extends from about 4 to 12 m/s for many systems.

The fast fluidization regime is most often encountered in systems called "circulating fluidized beds" where provision for continuous return of a significant flow of entrained solids is an integral part of the equipment. In combustion systems, the return is accomplished by capturing the entrained solids in one or more external cyclones or in impingement separators. The captured particles are then sent back to the base of the reactor or “riser” through a vertical standpipe, and then through a non-mechanical seal (e-g. a "loop seal" as indicated in Fig. 5) or a non-mechanical valve (e.g. an L-valve as indicated in Fig. 7 below). Mechanical valves (e.g. slide valves) are in common usage in FCC installations. In some cases. the solids pass through a separate low-velocity fluidized bed reactor (e.g. an FCC regenerator or low velocity fluid bed heat exchanger in some CFBC installations) during their journey from cyclone capture to reinjection. When there is sufficient capacity in the return loop and a flow control valve (e.g. an L-valve or slide valve), the concentration or density of solids in the reactor can be controlled as a separate variable, something which is considerably more difFicult in low-velocity systems.

Cooling watlsrwall

1 Standpipe

Fig. 5. Schematic showing principal features of circulating fluidiced bed combustion systems.

P2 High-velocity fluidized bed reactors

Table 2: Comparison of typical operating conditions for the two principal applications of fast fluldization:- fluid catalytic cracking and circulating fluidized bed combustion.

Particle density, kg/m3 Mean particle diameter, pm Particle size distribution Geldart powder group Superficial gas velocity, m/s Exit temperature, 'C Temperature uniformity Pressure, kPa Net solid flux, kg/m*= Appare t suspension density,

kg/m tt Exit geometry Outer wall geometry

FCC Riser Reactor

1100-1500 60-70 Quite broad A 8-18 increasing with height 500-550 Significant gradient 150-300 400-1400 160-650 at bottom, 50-80 at top Various Smooth (hot wall design) or hex mesh refractory (cold wall design)

CFB Combuator

1800-2600 150-250 Broad B 5-9 850-900 Uniform llO-L20 10-100 150-600 at bottom, lo-40 at top Abrupt Membrane wall with vertical tubes of dia. 0.04 to 0.1 m

Key features of circulating fluidized bed combustion systems are also shown schematically in Fig. 5. The bottom section of the riser is commonly tapered to prevent solids from sitting and agglomerating in the bottom section. In addition to the primary gas entry points at the bottom of the reactor, gas may be injected through secondary ports (and sometimes tertiary ports) at higher levels, this being especially useful for NO, control in CFB combustors. Instead of being smooth, the walls of the riser may be composed of membrane waterwall surfaces, similar to those employed in conventional pulverized coal and recovery boilers. For large units, wing walls may also project into the interior of the reactor chamber. These features can profoundly influence solids flow patterns in their vicinity (Wu et al, 1990). hence influencing mixing and gas/solid contacting.

Circulating fluidized beds utilizing the fast fluidization regime have been used for a number of gas-solid reactions including calcination, combustion of a wide variety of fuels, gasiffcation, and dry scrubbing of gas streams (Reh. 1986; Hirsch et al, 1986). Applications for catalytic reactions can be taken to include the transport reactors employed in modern catalytic cracking operations in the petroleum industry and certain Fischer-Tropsch synthesis reactors, with other applications. such as CFB production of maleic anhydride (Contractor, 1988) also under development. Some of the principal features of the two principal applications of the fast fluidization regime, fluid catalytic cracking and solid fuel combustion, are compared in Table 2. Conditions in circulating bed calcination, the third major application, tend to be similar to those in combustion.

Advantages of the fast fluidization regime, relative to the bubbling, slugging and turbulent regimes, include higher gas throughput per unit area, adjustable retention time of solids, limited axial dispersion of gas coupled with near uniformity of temperature and solids composition, reduced tendency for particles to undergo agglomeration, and possibility of staged addition of gaseous reactants at different levels. Separate reactions can also be carried out in the return path of the principal circulation loop. Gas-solid contacting also tends to be very favourable as shown in Fig. 3. However, CFB systems tend to have higher capital costs than Low-velocity systems, so that one or more of the above advantages must be very significant for this option to be viable.

FAST FLUIDIZATION: HYDRUDYNAMICS AND HEAT TRANSFER

Fast fluidized or dense riser transport reactors are subject to strong radial gradients, with significantly higher concentrations of solids near the outer wall than in the interior of the reactor (e-g- Bartholomew and Casagcande, 1957; Hunt et al. L957; Saxton and Worley, 1970; Bier1 et al, 1980: Schuurmans. 1980; Weinstein et al, 1986; Bader et al, 1988; Rhodes, 1990). This has led a number of workers (e.g. Brereton et al, 1988; Berruti and Kalogerakis, 1989; Rhodes, 1990) to propose simple core/annulus two-zone models for circulating fluidized beds, with one- dimensional upflow of gas and entrained solids in a dilute central core and downflow of dense strands or clusters assumed in a thin annular region at the wall. Gas is then assumed to travel upwards in the core. while either moving upwards much more slowly or moving downwards in the outer annular region.

While this simple picture captures some of the principal features of the flow, the true situation is much more complex. In reality, the interface between the two regions is diffuse and varies with time; the dilute core even reaches the wall at times, especially for dilute suspensions. There is a net outwards flux of solids to the wall layer over most of the riser height. In addition. the wall strands have a wide range of local voidages and velocities of descent. Figure 6 shows the variation of wall voLdages eeas~lred with a capacitance probe in recent cold model

1960 J. IX. GRACE P2

Void Fraction Fig. 6. Cumulative time fraction of wall coverage by strands of different voidage measured by Wu et al (1990) in 152 mm dia. x 9.3 m tall cold model column glth U-7 m/a. d 1 1171 pm, p -2650 kg/m .

8 Numbers on the curves ar

a ea-average apparent suspension densities in kg/m fj cross-sectional .

work in our laboratory. The mean downwards particle velocity has typically been found to be of order 0.5 to 1.5 m/s. As far as the dilute core is concerned, the time-mean particle velocity tends to have an approximately parabolic profile (Bader et al, 1988), with the centre-line gas velocity typically being 2 to 3 times the superficial gas velocity. Turbulent fluctuations are also prominent, especially in the lower region where turbulence likely originates in a lower aona of turbulent fluidization or due to secondary gas entry nozzles, solids feed nozzles or recycle ports.

Because of the extensive internal recirculation inside most CFB systems, as well as external recirculation, the temperature tends to be relatively uniform within CFBC systems (e.g. see Stromberg, 1982). Figure 7 plots temperature profiles for an experimental CFB combustor. This shows less than a 20 degree C overall variation of temperature for the entire principal recirculation loop. Heat transfer coefffcients depend strongly on hold-up of solids within the unit, and are typically of order 150 to 200 W/m2 K for large scale units (Grace, 1990). Radiation plays a much more prominent role than for conventional fluidized beds.

8 x 58

4 Qr----

E \l r: 4 .P

: 3- 'I 3

9 I 9

2

- 10 I

j--'* 0 1

850 ! 30' 0 r emperature, “c

Fig. 7. Steady state temperature profiles in a pilot scale (152 mm square x 7.3 m tall) CFB combustion system with highvale coal as fuel. IP8.1 mls; set/primary air ratio-1.1; oxygen In flue gas=3.2%; apparent overall suspension density-121 kg/m3.

P2 High-velocity fluid&d bed reactors 1961

ibruit exit a.

*0 200 400 -I

APPARENT DENSITY. kg/m3

Fig. 8. Effect of riser exit geometry on apparent density profiles tn a 152 mm dia. x 9.3 mm tall cold model CFB unit from Brereton (1987) with G -73 kg/m2s for both exits. U=7.lm/s; d 1177 pm; p -2650 kg/m 3;Ssolids return 1.98 m above air distribueor in bothPcases.

The geometric configuration of the riser employed in fast fluidized bed applications has a profound, and commonly unrecognized, influence on the hydrodynamic behaviour in such systems.

Exit Effects

Riser exits may be divided into two types: (a) "Once-through exits": These exits are smoothly curved or tapered as shown in Fig. 8, as in Synthol reactors. There is minimal internal separation of solids at the rap of the reactor. These exits allow large net circulation fluxes and are likely to be optimal when short uniform residence times of particles are required, as with quickly decaying catalysts. (b) "Internal reflux exits": These are abrupt exits. e.g. with a capped 90 degree top and side outlet as shown in Fig. 8. The abrupt exit causes substantial internal separation of entrained solids from the gas reaching the top of the reactor, so that there is a considerable internal reflux of particles down the walls of the reactor. This geometry is common in circulating fiuidiaed bed combustors. Experimental results (Brereton. 1987; Ambler, 1988; Bolton and Davidson, 1988; Brereton et al, 1988; Jin et al, 1988) show that the abrupt exit leads to substantially Increased solids hold-up at a given net solids flux, reduced pressure fluctuations and axial dispersion of gas at a given overall pressure drop. and an increase in solids density in the top portion of the riser. Typical profiles of apparent solids density obtained with two different exits appear in Fig. 8. These show that IIhe influence of the exit is felt right down to the bottom of the reactor, and not simply in the vicinity of the exit. It is clear that the designer can usa the exit geometry to promote or minimize reflection at the top of the riser, depending on the process.

Wall Geometry

There is also evidence that the wall geometry can profoundly influence the behaviour of fast fluidized systems. Wu et al (1990) showed that there were major differences in hydrodynamic behaviour depending on whether the outer wall is smooth or s membrane surface. For smooth- walled vessels, sheets of particles stream downwards along the outer wall, and particles are stripped off relatively quickly by the upflowing dilute core flow. On the other hand, for membrane surfaces, strands tended to run downwards as "rivers" in the channels or valleys between the tubes. The tubes protect the channels, with the result that particles travel downward much further before being arrested and stripped off into the interior.

In work carried out in New Zealand (Davies and Graham, 1988; Davies and O'Hagan, 1988), wall protrusions have been found to cause particles to be stripped off the wall of vertical risers, leading to reduced downflow of solids and lower pressure drops. Proprietary projecting surfaces on the wall have been used in industrial fast fluidized beds for the same purpose. Reactor internals, bends or projecting surfaces can have a significant influence on flow patterns and solids density profiles (Bartholomew and Casagrande. L957; Saxton and Morley, 1970; Grace et al, 1990). It is clear that the behaviour of reactors operated in the fast fluidizad mode may be very sensitive to the roughness of the wall or to any protrusions from the surface. Again this is a factor which has not been adequately recognized or exploited In reactor design.

1962 J. R. GRACE

'0 I I

300 200 3 NO, Concentration, ppm

P2

Pig. 9_ Oxygen and NO concentrations at different vertical and horizontal positions &side a 152 mm square x 7.3 m tall CFB combustion pilot plant. Ux7.6 m/s; T=1163K; set/prim air ratiowl.1; Ca:S molar ratio-3.0; oxygen in flue gas=3.0%; fuel: Minto coal. A, M. and W refer to axis, midpoint and wall, as shown on insert. Open symbols and broken lines - oxygen; closed symbols and solid lines: %.

DISPERSION, CONTACTING AND REACTION IN FAST FLUIDLZED BEDS

Experimental results (Ambler, 1988; Bader et al, 1988) show that some solids are carried up very rapidly In the core of the riser while others take much longer to reach the exit when they reach the outer annular region. Nevertheless, except for very fine particles, which may fail to be captured by the cyclone(e) on the first pass through the reactor loop, the extensive internal and external reflux of particles in circulating fluidixed systems ensures that the residence time distribution for all but the finest particles approaches closely to that of perfect mixing. Typically, particles make tens or even hundreds of circuits before leaving the system with each circuit requiring - 2 to 20 s in the riser (depending on the exit geometry, reactor height and superficial velocity) and minutes in the return system. Experimental findings (e.g. Grace et al. 1990) also show that the solids compositions at different locations within the primary circulation loop of CFBC systems tend to be nearly identical, suggesting that there is relatively little particle segregation within the riser or primary recirculation loop.

There is relatively little axial mixing of gas in CFB systems, except near the outer wall. There the downflowing strands drag gas downward with them, causing significant deviations from axially dispersed plug flow (Bader et al. 1988; Brereton et al, 1988). The axial dispersion model is not truly appropriate for this type of mixing pattern.

Lateral gas mixing in fast fluidised systems is one area of concern in fast fluidized systems. For example, this influences the number of solid fuel feed points required In CPB combustion systems. Experimental measurements obtained in a pilot scale circulating fluidized bed combustor and shown in Fig. 9 indicate that significant radial concentration gradients can persist over the entire height of fast fluidizsd beds. Here the concentration of oxygen is lowest near the wall where the carbon concentration is highest, due to the sheets of solids descending in the vicinity OP the outer wall. On the other hand, the NOx concentration is initially highest near the wall in the lower part of the reactor, due to increased release of fuel nitrogen in the outer region where there are more particles, but by the top of the riser, the NOx concentration has become lowest near the wall because of the carbon content of the descending sheets of solids there. The most reliable radial gas mixing studies (Adams, 1988; Bader et al, 1988) have suggested that effective radial gas dispersion coefficients are of the order of 30 to LOO cm2/s. with some indication that these decrease as the net solids circulation flux increases.

There are very few published experimental results allowing gas-solid contacting to be evaluated in circulating or fast fluidised beds having well-characterized hydrodynamics. It is therefore difficult to make definitive statements regarding the effectiveness of gas-solid contacting and the types of reactor model which may be appropriate for describfng CFB reactors. The earliest

P2 High-velocity ffuidized bed reactors 1963

: Two-ZoneModel

air

Fig. LO. Schematic diagram indicating flow patterns of solids (solid arrows) and gas (dashed arrows) and showing elements which need to be included in CFB reactor models.

published attempt to derive a model explicitly for a circulating bed reactor appears to be that of Weiss and Fett (1986) who treated a CFB reactor for sodium bicarbonate decomposition in terms of a series of homogeneous cells in series, with external recyle of solids. This model was later extended to solfd fuel combustion in an effort to predict transient behaviour (Weiss et al, 1988). In a very different approach, Hastaoglu et al (1988), while recognizing the wall-annulus separation within fast fluidiaed risers, adopted a reactor model which assumed that the gas and wood particles within a CFB wood gasifier moved through the bed in single phase plug flow.

Clearly, the core-annulus hydrodynamic structure of flow in the riser must be accounted for in any comprehensive model. Experience with circulating bed combustion also makes it clear that the cyclone can play an important role in burning out volatiles, meaning that the cyclone must be modelled as an integral part of the CFB system, at least for some reactions. Moreover, it may well be necessary to make separate allowance for the region of different hydrodynamics (commonly turbulent fluidization) at the bottom of CFB risers, and heat removal or losses within the reactor and return system ran also be significant. Some of the elements which need to be incorporated within CFB reactor models are shown schematically in Figure 10. It may again be possible to adapt bubbling bed reactor models for at least some parts of the system. For example, Table 1 indicates a possible set of parameters which could be adopted in a first approximatFon for that part of the riser (typically all but the bottom 10%) which is operating in the fast fluidization regime). Alternatively, one could utilize the two-zone counter-current flow model of Van Deemter (1961), for the major part of the riser. Either approach would have to be coupled with separate allowance, perhaps simply CSTR behaviour, for the entry region. As already noted, it is essential <also to model reactions occurring in the cyclone(s), and allowance is also required for heat release and exchange in the entire primary circulation loop, including the loop seal, L-valve or other solids return device. This composite model could account for the principal features of most CFB reactor systems. Naturally, when separate reactions occur in the recycLe loop, as with FCC regenerators, allowance for these will have to be included. If such models prove to be unable to predict the behaviour. it may be necessary to consider more complex models which make allowance for such features as the flow development with height, radial gradients and reflection of particles from the top of the reactor. It is likely that a number of simple and more complex models will be written and tested in the coming decade.

DIRECTIONS FOR FUTURE WORK

It is clear that turbulent and fast fluidized beds are having major impacts as chemical reactors both Ear some gas phase solid-catalysed reactions and for some gas-solids reactions. However, as has so often been the case in the field of fluidfzation, applications have tended to precede real

1964 J.R.GRAcE P2

understanding, and this creates the dangers of misapplication and under-utilization of the technology. Some areas in which considerable work is required in the coming decade are as follo"s:

-While the high-velocity regimes have been identified and studied, we still have a rather poor idea of how to predict the transitions from one hydrodynamic regime to another.

-As discussed above, even a rudimentary picture of the turbulent regime is lacking. It is not even clear whether we are talking about true turbulence, in the sense of random fluctuations of such variables as local voidage and velocity, or whether the regime consists of a transitional regime of intermittency between lower and higher velocity occurrences.

-While a more coherent picture of the fast fluidization regime has emerged, there are still unknown factors regarding the mechanisms which govern the development of a core-annulus structure within the riser. It is also now evident that the exit and wall geometries have a very important influence on the flow and mixing patterns within the risers. There is a pressing need to optimize the geometric factors.

-It is also probable that particles size distribution plays an important role in turbulent and fast fluidization. as in the lower velocity regimes. More work is required both to understand the role of particle size distribution, as well as to determine the effects of other variables like particle shape, particle density and gas properties.

-It is important, both for catalytic operations and for gas-solid reactions, that models be established which allow simulation of unsteady state conditions, as well as steady state conditions. Energy considerations will be critical factors in these models.

-Erosion has turned out to be a critical problem in many of the circulating fluidized bed combustors which have been put into operation in the past decade. It is essential to develop an understanding of the factors which promote wall erosion and to develop strategies for alleviating this problem.

ACKNOWLEDGEMENT

I am indebted to Clive Brereton, Jim Lim. Richard Senior and Amos Avidan for useful discussion of a number of the issues treated in this paper.

NOMENCLATURE

A AK D z?

dp* C", g H H mf % k k' Fe U u* u mf "T

cross-sectional area, m* Archimedes number effective overall axial dispersion coefficient 1

or gas. m*/s dense phase gas axial dispersion coefficient, m /s mean particle diameter, m dimensionless particle diayeter net circulation flux, kg/m s acceleration of gravity, m/s 2

expanded bed depth, m bed depth corresponding to minimum fluidization. m interphase mass transfer coefficient, m/s first order rate constant, based on particle volume, s-l dimensionless rate constant Peclet number for gas axial dispersion superficial gas velocity, m/s dimensionless gas velocity superficial gas velocity at minimum fluidization, m/s particle terminal settling velocity, m/s

fracrion of gas passing through the more dilute phase of difference in density between the particles and gas. kg/m

se&on

fraction of solid !I

which are present in the more dilute phase gas density, kg/m particle density, kq/m3 gas viscosity, Ns/m bed voidage

P2 High-velocity fluidized bed reactors 1965

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