- Email: [email protected]
An innovative non-mechanical solids feeder for high solids mass fluxes in circulating fluidized bed risers
IBIIItimI]~ ELSEVIER
Powder
Technology88 (1996) 123-131
An innovative non-mechanical solids feeder for high solids mass fluxes in
circulating fluidized bed risers T.S. Pugsley
~, B . J . M i l n e 2 F . B e r r u t i
Deparlm~nt of Chemical Engineering. The University of Calgary, Calga~%Alia T2N IN4, Cunuda
Received 19 May 1994;revised~ November 19~5
Abstract
A non-mechanical solids feeder for circulating fluidized bed risers has been developed to achieve very high and carefully controllable solids mass fluxes in a laboratory scale experimental unit. The design involves the replacement of the traditional standpipe/L-valve configuration with an aerated annular bed of solids and an innovative radial gas distributor surrounding the base of the riser. Solids enter the riser threugh orifices drilled at the riser base. Solids mass fluxes of up to 700 kg/ra2 s for sand have been achieved in a 0.05 in diameter, 5 m tall riser at riser superficial gas velocities between 5.5 and 8.5 m/s. This paper reports results of the experimental investigation of the circulating fluidized bed pressure loop in order to understand better the operation of the solids feeder and to gain insight into proper design and scale.~p of the device. Keywords: Circulatingbeds; Non-mechanicalvalve;Highsolidsfluxes;Pressure loop
1. Intreduetlen
Circulating fluidized beds (CFBs) have been successfully used for the combustion of low-grade fossil fuels and mineral processing since the 1970s [ I ]. More recently, the CFB has also gained attention in the petrochemical industry as a catalytic reactor for the preriuction of certain specialty chemicals [2] and in the petroleum industry as the riser nnd/ur regenerator in fluid catalytic cracking (FCC) units [3,4]. While CFB combusturs operate at low solids mass fluxes ( ~ 20 k g / m 2 s), FCC risers operate in the range of 500 kg/ ilia a, and CFB catalytic reactors operate at |000 k g / m 2 s or greater. The hydredynamicsin CFB risers at these high solids mass fluxes may be radically different. While a core-annular flow structure with solids downflow next to the wall has been consistently observed in risers of CFB combustors, gas residence time distribution (RTD) studies on the du Pont butane oxidation demonstration plant reported by Contractor et al. [5 ] indicate that the predominant flow structure at high solids mass fluxes is no longer core-annulus with solids reflux. The majority of hydrodynamic studies reported in the literatuse for laboratory scale CFBs have been conducted at t Depanment of ChemicalEngineering,The Universityof New Brunswick, Fmdericton.NB E3B 5A3. Canada. 2 Departmentof ChetmcalEngineering,The Universityof Saskatchewan, S~katoon, Sask. STN 5C9, Canada, flO32-5910/9615tg.00© 1996ElsevierS¢~er.ceS.A. All rightsreserved PIIS0032-5910(96)031 Ig-X
relatively low solids mass fluxes. The standpipe height required to build up the pressure head necessary to achieve very high solids fluxes is prohibitively large in a laboratory setting, The objective of the present work was to develop a non-mechanical solids feeder device for achieving high solids mass fluxes in a laboratory scale riser that eliminates the traditional standpipe/L-valve configuration. The result is a new technology for a solids feeder applicable to both pilot and industrial scale applications requiring high and carefully controllable solids mass fluxes.
2. Experimental 2.1. Apparatus
A schematic of the experimental CFB equipped with the non-mechanical solids feeder is given in Fig. I. The riser section is constructed of a series of 5 cm i.d. flangedplexiglass sections. Any number of available sections may be connected so that the riser height can vary from a total o r s to 10.5 m . The first 1.8 m of riser length is encased in a 20 cm i.d. plexiglass vessel. This vessel may be filled with solids to a height o f 1.2 m. At the bose o f the riser, which sits inside the annular bed of solids, eigh~ equally-spaced 1 cm orifices are drilled around the riser circumference. Fluidizing gas to the
124
TS. Pugsley el aL / Powder Teclmology ~8 (1996) 123-131
,clone
leg
lq
Fig.2. Detailedscaledrawingof the radialdistributor( 1:3scale).
Fig. I. Schematicof the expedmcnlalCFBequippedwiththe non-mechanical solid~feeder {notto scale). system is split between two streams, the riser gas and the auxiliary, gas streams, both individually controlled by rotameters. The auxiliary gas stream flows through the innovative radial gas distributor, and into the annular bed of solids surrounding the riser. A fraction of the gas may tend to flow upward through the annular bed, but the majority of thc gas follows the path of least resistance through the orifices drilled at the riser base. The drag of the gas on the solids as it flows through the orifices carries the solids from the annular bed and into the riser. The solids are then transported vortically in the riser by the high velocity riser gas stream which is simply the sum of the riser gas plus the fraction of the auxiliary gas flowing through the orifices. The gas-solids suspension leaves the riser via an abrupt exit configuration. Two cyclones attached in series separate the gas and solids, and the solids arc transferred by a return leg to the top of the annular bed of solids. A scale drawing of the radial distributor cross-section is shown in Fig. 2. The distributor consists of a 20 cm i.d., 5 cm long plexiglass section with 1000 regularly spaced holes of 0.5 mm diameter drilled along its circumference. The 1000 holes are divided among 5 rows of 200 holes each, and each
row is staggered, resulting in a regular triangular pitch between any three holes. A 300 mesh (75 p.m aperture) stainless steel wire mesh is spot welded to the inside and outside walls of the distributor to prevent loss of fines into the windbo×. The entire distributor arrangement is connected by flanges to the annular bed of solids to facilitate removal for cleaning or design modifications, if necessary. The auxiliary ga~ is introduced at four, equally spaced locations around the circumference of the windbox to further ensure proper gas distribution. These locations are connected by low pressure polyethylene tubing to the rotameter controlling the auxiliary gas Now. In order to prevent jet penetration through the windbox and hence maldistribution of the gas, small plexiglass deflectors are attached to the outer wall of the windbox at the four locations where the auxiliary gas is introduced. 2.2. I n s t r u m e n t a t i o n
In this work, the solids circulation rate is measured with a 3.2 mm o.d. stainless steel rod attached to a small piece of stainless steel mesh (see Fig. I ). Solids flow downward in the annular bed in a moving packed bed flow. The stainless steel rod is dropped onto the upper surface of the annular bed, well away from the wall. Hence the rod moves in the bulk solids flow so that the effect of the vessel wall, which was found by Patience et al. [6] to result in underestimation o1" the solids mass flux, is eliminated. As the original upper surface of the bed descends, the rod descends with it, and its rate of descent is measured by tracking the movement of a mark etched on the rod over a known distance. Th~. corresponds to the downward bulk velocity of the solids in the annular bed. Happel [ 7] showed that the voidage of a moving packed bed is approximately equal to that of a loosely packed bed of the same solids. Assuming the loosely packed voidage in the annular bed, and knowing the annular bed cross-sectional area along with the downward solids velocity, the total mass flowrate of solids into the riser through all of tho orifices
~LS. Pugsley et aL /Powder Technology 88 f1996) I23-131
Table I Physicalcharacterizationof LaneMountainsilicasand PropeRv Particledensity(kg/m3) Bulkdensity(kg/m3) Meandianmlerterm) Packedbed voidage MinimumfluidizaUonvelocity(m/s) Minimumfluidizationvoidage Terminalsettlingvelocity(m/s) Sphericity
2.5.80 1650 208 0.36
0.038 0.45 1.5 0.90
may be calculated. Dividingthis mass flow rate by the crosssectional area of the riser gives the riser solids mass flux. The sulids used throughout this study were Lane Mountain silica. The physical properties of the sand are summarized in Table 1. In order to ascertain the minimum fluidization velocity of this material, a fluidization test was performed in an available 10.2 cm bed. Daring these experiments, it was observed that bed expansion was negligible prior to incipient fluidization conditions. Thus it is concluded that the voidage remains constant and equal to the loosely packed bed voidage until incipient fluidization, and the calculation of solids mass flux described above is accurate. Beyond incipient fluidization, bubbles will pass through the annular bed and the descent of the stainless sleet ~od is no longer smooth, making the measurement inaccurate. Results of these experiments and those reported by Milne et al. [8 ] have proven the descending rod measurement technique to be very reliable and entirely reproducible when sand is used as the circulating solids. When using fluid catalytic cracking catalyst or other Geldart A type particles, the descending rod method will require more careful calibration, As conditions of incipient fluidization are approached, a bed of FCC catalyst expands significantly. Therefore, it is not valid to assume the voidage equal to that of a loosely packed bed of FCC catalyst. Furthermore, as the fine FCC catalyst circulates through the system, air may become trapped in the interstices, and the annular bed can he significantly aerated by this trapped gas. The circulating fluidized bed apparatus is well tapped to facilitate cxamiantion of the entire pressure loop. The riser section is equipped with 21 pressure taps when operating at a height of 5 m, and no less than 37 taps when operating at a height of 10.5 m. The taps are fabricated from 6.35 mm threaded brass tubing inserted flush to the tuner wall of the riser, and 200 mesh stainless steel wire mash is spot welded to the taps to prevent backflow of solids and clogging of the lines. Tygnn tubing connects the pressure taps to a bank of U-tube water manometers. The first tap is located 5 cm above the orifices drilled at the riser base, and is denoted as point A in Fig. 1. The subsequent 11 taps are equally spaced at 15 era and the balance are spaced at 30 era. The pressure tap at the riser exit is indicated as point B in Fig. I. The pressure drop in the annular bed is measured by three 6.35 mm o.d. brass taps inserted at axial locations of 86, 40,
125
and 10 cm above the riser orifices ( points D, E, and F, respectively, in Fig. I ). A fourth tap is immersed in the annular bed at the level of the orifices (tap G). A final 6.35 mm brass pressure tap (tap C) is Ineated 2,5 cm belqw the primary, cyclone outlet. This tap, in conjunction with tap B, moni#.o,-x the pressure drop over the exit region of the riser and the primary cyclone. The point labelled C ' in Fig. 1 is the pressure at the upper surface of the annular bed of solids. This value was not actually measured, but was instead obtained by extrapolating the pressure gradient in the annular bed between points D, E, and F to the upper bed surface.
3, Critical assessment of the non-mechanical guilds feeder design Replacing the traditional standpipe/L-valvedesign of circulating fluidized beds with the annular bed of solids and radial distributor results in a unique solids feeder device, hut the concept of an annular solids feeder is not an entirely new idea. The most recent example of an annular fluidizod bed surrounding the base of the riser is seen in the work of Harris and Davidson [9]. However, :be circulating fluidized bed of Harris and Davidson [91 still has a standpipe and slide valve to control solids flow. The annular bed behaves similarly to the fluidized bed in an overflow standpipe design. Solidsentor the bed from the standpipe and s,mply overflow in order to establish a more uniform solids distribution to the riser. The annular solids feeder designed in this work is more closely related to the systems devised by Hirama et al. [ !0] and Fusey [ 11]. The circulating fluidized bed of the latter study essentially consisted of an insert capped by a baffle and placed in an existing fluidized bed of solids. Experiments were performed with FCC catalyst and PVC resin in which the annular bed was maintained at, or just below, minimum fluidization conditions. The maximum level of solids in the annular bed was 80 cm and the maximum solids mass flux aehievod in this rig was 71 kg/m2 s. This design of the annular feeder appeared to lead to operational problems. Fusey [ ! 1] reported that the pressure at the base of the insert was the same as that at the base of the annular bed. This would seem to suggest a loss of conUol of the solids flow. Fusoy [ 1I] noted that one drawback of this design was less direct and precise control of solids mass flux. The solids feeder described in this paper is different due to the "arlial gas dis~ibutor. This design provides superior solids flow coo|xol and higher solids mass fluxes than that of Fusey [ t 1]. Furthermore, the annular bed in the work of Fusey [ 11] as well as that of Hirama e t at. [ 10] was maintained at or just below conditions of incipient fluidization. Superior flow control is also realized with the solids feeder developed in this work by operation below incipient fluidization in the annular bed. Valuable design experience for the annular solids feeder was gained from the work of Milne [ 12] which studied the inteqrnally circulating fluidized bed (ICFB). As the name
126
Z& Pugsleyet aL /Powder Technology88 (t 996) 123-131
suggests, the gas-solids suspension flowing in the riser of this circulating fluidized bed does not exit to an external cyclone, but, instead, the solids impinge on a tee-shaped ballistic separator and drop back down to the surface of a moving packed bed of annular solids. This design enables the solids to circulate within a single vessel. The solids feeder system in the ICFB consists of the combination of auxiliary gas, distributor, and orifices drilled at the riser base, However, the auxiliary gas is supplied to the bottom oftbe windbox, so that the gas has asignificant vertical component upon entering the annular bed. The innovative radial distributor design described in Fig. 2 is an advancement over the ICFB design because the auxiliary gas entering the distributor has no vertical velocity component. This results in a reduction of the fraction of gas passing upward through the annular bed for a given auxiliary gas flow, and hence a higher flow rate of solids through the orifices.
4. Experimental results After installation of the non-mechanical solids feeder on the laboratory scale circulating fluidized bed, systematic experiments were undertaken to investigate the feeder performance. It was desired to establish the ability of the feeder to achieve very high and carefully controllable solids mass fluxes and solids-to-gas mass ratios in the riser. A complete understanding of the workin~ ef the feeder required analysis of the entire CFB pressure loop under various operating conditions. 4.1. Control of solids massflux
The ability of the solids feeder to achieve very high and carefully controllable solids mass fluxes is depicted in Fig. 3. In this figure, the solids mass flux of sand and the solids flow per orifice is plotted as a limction oftbe auxiliary gas flow at 800
700
0.16 ~Uol=S n.S/o=t : 0.12 0.08 ~L 200
0
0.04 i
i
i
i
1
2
3
4
AuxililUy Gas Flow (sL/s)
Fig. 3. Solidsflowperorificeas a functionof auxiliarygas flowfor varyiag risersuperficialvelocities.Riserheight~ 5 m.
two different riser gas superficial velocities. The solids flow per orifice is simply obtained by multiplying the solids mass flux by the riser cross-sectional area, and dividing by the total number of orifices drilled at the base of the riser. The data points are represented by the symbols, while the solid lines are the result of linear regression of the data at e~ch riser gas superficial velocity. As indicated by this figure, the solids feeder is able to supply a maximum solids mass flux of 700 kg/m 2 s in the 5 m riser. This represents a solids-to-gas mass ratio of 68.6 at a gas velocity of 8.5 m/s. Comparing these values with the solids mass fluxes and solids-to-gas mass ratios listed in Table 2, it is seen that the magnitude of tbese quantities in this work surpasses nearly all of the previously reported results of other authors. While Table 2 is not a complete list of references, it is representative of the magnitude of the solids mass fluxes and solids-to-gas mass ratios found in the open literature. In this work, high solids mass fluxes and solids-to-gas mass raties in the range of the du Pont pilot scale CFB catalytic reactor ( Contractor et al., [ .5] ) have been achieved in a laboratory scale riser of only 5 m. Furthermore, the du Pont rig employed a slide valve to control the solids circulation as opposed to the non-mechanical solids feeder of this study. Contractor etal. I5] did report a solids mass flux 685 kg/m z s and a gas velocity of 5.7 m/s in their work, giving a solids-to-gas mass ratio of 100. Such conditions could not be achieved in the laboratory scale CFB due to the small riser diameter. Feeding solids at such a rate at that gas velocity led to unstable operation as solids tended to accumulate at the inlet and form slugs. The results of the regression analysis in Fig. 3 show that the data is a well-behaved linear function of the auxiliary gas flow rate. This corroborates the findings ofMilne [ 12] in the ICFB. MiMe [ 12] also concluded that the solids flow per orifice, or solids mass flux, in the ICFB was independent of riser gas velocity. Fig. 3 suggests a slightly more eomplex behaviour in the CFB of this work. At the lower solids mass fluxes, the two lines coincide proving that the solids mass flux is indeed independent of the riser gas velocity in this region. However, the slope of the regressed line is steeper at the higher gas velocity of 8.5 m/s, indicating a sharper response to the auxiliary gas flow, or equivalently, the solids mass flux. Similar behaviour was found by Fusey [ 11 ]. Therefore, the curves diverge at higher auxiliary gas flows so that the solids mass flux eventually becomes a function oftbe riser gas velocity. The seemingly contradictory findings of Milne [ 12] may be attributed to the short riser and high solids mass fluxes in the 1CFB resulting in the solids being in accelerating flow for the entire riser length. As a result, the solids mass flux was relatively insensitive to the riser gas velocity in that work. 4.'2. Pressure loop study
Results of the pressure drop across the riser in a circulating fluidized bed are now abundant in the literature. However. the cyclone, solids re-injection valve, and standpipe pressure
"~S. Pugsley et aL /Powder Technology 88 (1996] 123-131
127
Table 2 Sample of published expcriroentalwork indicatingsolids mass fluxes and solids-to-gasmass ratios achieved Height
Author
(m]
Bader etal. [23] Har~geet aL [24] Horio el al. [25] Arenaet al. [26] Brereton and Grace [27] Bodelein etal. [28] Schnitzleinand Weinstein [29] Rhodes etal. [ 30] Bal el al. [ 31] Lcuge et al. [ 32] Miller and Gidaspow [33[ Wong [34] Mori etal. [ 35] Azzi etal, [36] This work Contractor et al. [5] Viitanen [37]
12.00 8.40 2.80 6.40
9.30 10.Q0 6.50 62.0 8 DO 7.00 550 3.00 4.55 ] 1.70 5.00 27.00 39.00
Diameter (m)
G, (kg/m 2 s)
Valve
0.3050 0A000 0.0500 0.0410 0.1520 0.1440 0.1520 0.1520 0.1860 0.1970 0.0750 0.0500 0.0500 0.1900 0.0500 0.1500 1.0000
147.0 30.0 11.7 500.0 116.0 52.9 250.0 I I ! .0 140.0 125.0 32.8 126.0 150.0 150.0 700.0 590.0 485D
slide L-valve beaerfiy/aearauon butterfly/aeration L-valve L-valve butzerfiy/aeration L-valve ?? gate valve slk[e valve b~ll valve U-heed/aeration slide valve solids feeder slide valve slide valve
drop are often not considered. Along with the riser, these components comprise the CFB pressure loop and merit equal consideration, especially the standpipe, or in the case of this work, the annular bed of solids. The standpipe has been referred to in the genre of fluidizatiun as the engine that drives a circulating system (see Reddy Karri and Knowiton [ 13] ). A typical t-'ressure loop for the CFB is presented in Fig. 4 for a solids mass flux of 360 kg/lll 2 s and riser gas superficial velocity of 6 m/s. The solid symbols represent the pressure measured by the riser taps, and the open symbols the pressure measured on the solids return side of the loop (i.e. the cyclone and annular bed).
B
e-
~
•
e
5-
is2. 1-
¢
Pntma'etPt ~ m )
Fig. 4, Typical plessme loop at a solids ~ flux, G,~ 3~Okg/m2 s, and riser gas superficialvelocity, fie= 6 m/s. Riser height= 6.4 m.
Uo (m/s)
9.1 3.7 1.2 7.0 7.1 4.33 4.0 5.0 7.0 5.0 2.89 8.0 5.0 6.2 8.5 7.8 13
w,l• 13.46 6.76 8.13 59.52 13.62 10.18 52.08 1&50 16.67 20.83 9.46 13.13 25.00 20.16 68.63 63.03 31.09
Several points may be gleaned from this plot. Firstly, the highest pressure point occurs at point G, immediately upstream of tl:~ orifices and at the bottom of the annular bed of solids. At this location, the auxiliary gas is introduced through the radial distributor. The gas will then either flow through the orifices, or pass upward through the annular bed. In either case, the flow of gas results in z pressure chop, so that the pressure at point G must be the maximum in the system. The lowest pressure point in the system is seen at point C at the prir.~2.,'y cyclnne sofids exit. The pressuse at this point is only slightly above atmospheric. Fig. 4 also illustrates the pressure drop associated with flow of solids and gas through the orifices from point G to point A. For these particular operating conditions, this pressure chop is app[oximately 850 Pa. The pressure drop through the riser is also clear from this figure. At the riser base, the prassure drop is very shazp as the solids entering the riser from the annular bed arc accelerated. At approximately 2 m, the pressure gradient becomes constant until the vicinity o f the riser exit. where the abrupt exit configuration causes some donsifica.lion. A slight pressure build-up is seen in the return leg as the solids flow from the primary cyclone exit to the annular bed. This is consistent with the pressure increase associated with lean phase downflow in CFB dowocomers as described by Wirth [ 14]. Fiaally, in the annular bed o f solids, the pressure increases from point C ' to G. If the loop is considered to begin at point G, the high wessut~ point in the system, then there exists a series o f pressure chops downstream o f that point through the orifices, riser, and cyclone. A small fraction of this pressure drop is recovered in the return leg, but from a practical standpoint, most o f the pressure drop is recovered in the annular bed. Hence the annular bed serves the same purpose as a conventional standpipe by transporting solids from a low pressure point ( C ) to a high pressure point ( G ) . The pressure increase from C ' to G mast balance the sum of
T.S. Pugsley¢t al. /Powder Technology 88 (1996) 123-131
t28
a
: ~
I - o - ~ = 24o k~mZS:Uo. . . . . /s
s=
; o = 5 8.m / s
s
21" o+
i
3000
~
6O00
9000
-
- 7
12O00
15000
[hcssu.+'e(Pl. glUBC) Fig. 5. Comparison of pressure loops at a solids mass flux, G~=240 kg/m a s, and two riser g~ superficial velocities. Riser height =6.4 m the pressure drops and close the pressure loop. Using the
nomenclature of Knowlton [ 15], the pressure drop through the annular bed is the dependent portion of the pressure loop, while the pressure drops from G to A (orifices), A to B (riser), and B to C (cyclone) ma~:eup the independent portion of the loop. Therefore: A P c c , + A P c , o = APGA + A P ~ + ABc
(1)
Based on the data of Fig. 4, the orifice pressure drop is 850 Pa, the riser pressure drop is 8036 Pa, and the cyclone pressure drop is 2243 Pa, for a total independent pressure drop of I 1129 Pa. The measured pressure gain in the annular bed is 10903 Pa which closes the loop to within 2.1%. This excellent closure indicates that the discrepancy associated with neglecting the pressure increase in the return leg is only slight. Hence, in subsequent analyses of the pressure loop, the pressure increase in the return leg is ignored. The pressure drop in the riser due to solids acceleration and hydrostatic head constitutes over 70% of the total pressure drop in the system, while the pressure drop through the orifices is only approximately 7.5% of the total. Fig. 5 compares the pressure loops for a solids mass flux of 240 kg/re 2 s at two riser gas superficial velocities. For the perpose of this comparison, the intermediate pressure readings in the riser have been eliminated, and points A and B joined by a straight line. The result of decke~iag ~ e ga~ velocity at a fixed solids mass flux is an increase in the total riser pressure drop due to increased solids suspension density, and an increase in the cyclone pressure drop. In order to compensate for the increases, the pressure drop through the annular bed of solids must also increase, and this is reflected by the larger pressure gradient in the bed at the riser gas superficial velocity ofS.0 m/s. The pressure gradient diatates
the pressure at point G, and results in a pressure drop through the orifices that is less for the case of lower riser gas superficial velocity. Milne [ 12] presented a correlation for predicting the solids flow per orifice in the ICFB as a unique function of orifice pressure drop. In Fig. 5, the solids flow per orifice is the same for both curves, but the pressure drop across the orifices changes w~th riser gas superficial velocity. This suggests that the flow of solids through the orifices is not a unique function of the orifice pressure drop in this system. This is attributed to the inject ion of the solids into a closed circulation loop which creates a certain back pressure at the level of the orifices. The expression of Milne [ 12] was derived by modifying equalions for the discharge of solids from aerated bins and hoppers to atmosphere. The open literature abounds with such studies (for example [16-19]) presenting experimental results ~nd correlations for the gas and solids flow as a function of the orifice Jpressure ¢~op. A slightly different study was performed by Chin nt at. [20] where solids were discharged through multiple orifices from the inner surface of an aerated annular bed of solids. Their experimental set-up was very similarto the annular bed solids feeder described in this paper. However, as tha solids flowed from the annular bed, through the orifices, and into a central tube, they did not enter a circulation loop as is the case in our work. Instead, they were , ,oply jetted to atmosphere and collected in a receiver at the bottom of'he tube. Based on the data collected in the present study, it appears that the work of Chin et at. [ 20], as well as all other aforementioned references dealing with solids and gas discharge from orifices [ 16--19], cannot be extended in a straightforward manner to analyze the performance of the new solids feeder. This is because these studies jetted solids and aeration gas through orifices against a constant back pressure ( aUnospheric ).The success of Milne [ 12] in eorral~ing the solids flow per orifice as a function of orifice pressure drop is attributed to the solids acceleration in the ICFB riser. This effectively created a constant back pressure against which solids were injected by removing the dependence on the riser gas velocity. The influence of the back-pressure introduced by the closed circulation loop in the CFB is not yet clear. It seems likely that the pressure field around the orifices in the annular bed is affected by the back pressrun. Experiments arc planned to investigate this hypothesis by traversing the tap at point G to establish the pressure field for various riser operating conditions. Table 3 summarizes the operating conditions of solids mass flux and riser gas superficial velocity investigated, along with the corresponding auxiliary gas flow-rate, Qa=~,the flowrate of gas throagh the orifices, Qo, and the orifice pressure drop, APo. "fins table does not indicate any clear trend of solids mass flux as a function of orifice pressure drop. What is important to note from this table, however, is the relatively low orifice pressure drop at high solids flows. This pressure drop is much less than that reported for flow of solids through L-valves. For example, Rhodes and Lausmann [21] measured L-valve pressure drops between 4 and 10 kPa, for riser
129
T S. Pugsley et al. I Powd~r Techno[ogy g~ (1996) 123-131
Table3 Summaryof orificepressuredropdataand orificegas flowat venoussolids massfluxesand riser supedicialgas velocities Gs (kg/m2s)
Uo (m/s)
Q,, (sL/s)
Qo (sL/s)
Aeo
8.1 10.5 19.8 25.0 150
2.6 2.7 3.4 3.6 5.5
240
5.8
5.0 6.0
036 0,39 0,39 0.69 0.93 1.42 1.69 1.78
186.8 31.1 112.1 623.0 485.0 996.7
240 360
0.50 0,45 oA5 0.98 0,98 1,45 1.79 1.79
(Pa)
6"12.8
847.2
solids mass fluxes of only 5 to 50 kg/m 2 s. Ruddy Karri and Knowlton [ 13] reported an L-valve prossure drop of 20 to 30 kPa in their test rig. The standpipe height was 8.5 m. The low pressure drop across the orifices is important for achieving high solids Ioadings in the riser. The requirements for pressure recovery in the annular bed due to the orifices is low, which allows for a larger pressure drop or, equivalently, higher solids ~aass flux in the riser. 4.3. Annular b e d pressure drop a n d auxiliary gas bypassing
The pressure drop through the moving packed bed is a function of the relative velocity between the downward flowing solids and the fraction of auxiliary gas flowing in the bed. This may be described by the Ergun equation using the relative velocity for moving packed bed flow, v; AP _ ( i - e,~) o'.v~ r- z p,.v I T 2 -b 1.75i - ¢.~, r - .~.f=.~. --=iSU
(2)
where the relative gas velocity is calculated from: v,=vp-v 8
(3)
with velocity positive in the downward direction. Therefore, the fraction of auxiliary gas that passes through the annular bed of solids, vs, is determined from the pressure drop in the dependent portion of the loop. F_,q.(2) bulds until the point of incipient fluidization, where the pressure drop through the bed becomes constant with
changes the riser solids mass flux. Thus the feeder design ig consistent with the valve mode of opuratiou of non-mechaw ical valves described by Knowlton [ 15]. Eq. (2) also indicates that for a movingpacked bed, incipient flaldizarion conditions are reached when the relative velocity is equal m the incipient fluidizatiou velocity, as opposed to the superficial gas velocity in a stationary packed bed. Therefore, based on Eq. (3), if the downward solids velocity is much larger than the upward gas velocity, the annular bed may become fluidized even though the gas velocity in the annular bed is very small. A large downward solids velocity is indicative of a high riser solids mass flux, which consequently requires a larger pressure drop in the annular bed in order to close the pressure loop. In order to maximize the solids mass flux attainable in the system, it is desirable to achieve a high solids mass flux with as little auxiliary gas flow as possible. This highlightsthe importaxce of the radial distributor design in this work. The auxiliary gas enters with a zero vertical velocity component, and has a greater tendency to pass through the orifices. Therefore, more solids axe dragged through the orifices at a lower au~liasy gas flow, and less gas is available potentially to pass upward through the annular bed. With velocity defined as negative in the upward direction, it is seen from Eq. (3) that if mum gas flows upward for a given downward solids velocity ( or given solids mass flux), the relative velocity is higher, and the bed fluidizas earlier to give a lower value of maximum solids mass flux. The success of the radial distributor in reducing the flow of gas through the annular bed is illustrated in Fig. 6, which plots the flow of gas through the orifices versus the auxiliary gas flow. With the exception of the data points represented by the open circles, better than 95,% of the auxilimy gas flows through the orifices in all cases. Thus the fraction of gas 4.0
S.5 3.0
:~ zs
increasing velocity. This sets the upper l i m i t on the riser solids
mass flux. Further increases in the auxiliary gas flow can no longer raise the solids mass flux because the maximum pressure drop has been realizod in the annular bed. As the pressure loop must always close, the condition of incipient fluidization corresponds to the maximum solids mass flux that can he achieved in the system. The previous studies using fluidized bed solids feeders [10,11] maintained the bed at incipient fluidizatiou conditions. Thus, the bed could not be used to control the solids flow rate and the solids mass flux was simply controlled by the riser gas velocity. Since die solids feeder described in this work is operated below minimum fluidizatiou conditions, varying the auxilimy gas flow
].o o.s o,o 0.0
i
i'
i
,
,
i
i
0.5
1JO
1.5
ZO
2.5
3.O
3.S
Fig.6. Frncdonof auxi]ia~gas passing0=roughII~ odfic~s.
q,O
130
Z S. Pussley et aL /Powder Technology88 (1996) 123-131
diverted through the orifices with the radial distributor design is greater than that in the ICFB of Milne [ 12 ], where orifice gas was approximately 85% of the auxiliary gas flaw. The points specially designated by open circles indicate that only roughly 70 to 80% of the auxiliary gas flows through the orifices. These points correspond to a range of solids mass fluxes of 8 to 25 kg/m 2 s and riser gas superficial velocities of 2.6 to 3.6 m/s. These operating conditions are typical of the turbulent fluid~zation regime. Even though the net solids mass flux is low, the dense bed characteristic of the turbulent regime forms at the riser base and results in a large total riser pressure drop. As a result, the pressure gradient in the annular bed of solids must be large in order to close the pressure loop. As the solids mass flux is low, the downward particle velocity in the annular bed in Eq, (3) is low. but, in order to generate the required pressure drop, the relative velocity between solids and gas must be high. Therefore, the upward gas flow through the annular bed in Eq. (3) is necessarily very high. In summary, the fraction of auxiliary gas that passes through the orifices (or upward through the annular bed) is a result of the dependent pressure drop in the annular bed of solids required to close the pressure loop. The radial distributor provides a design that is able to direct 95% or more of the auxiliary gas flow through the orifices. Such a design maximizes the solids mass flux. A lower upward gas velocity in the annular bed (Vg in Eq. (3)) allows for a very high downward solids velocity without fluidizing the annular bed, and hence, a very high solids mass flux in the riser. 4.4. Design and scale-up o f the solids feeder
Only one feeder of the dimensions given in Fig. 2 was constructed for this work. However, the study of the pressure loop provides valuable insight into the proper design and scale-up of this device for CFB risers of different sizes. In order to increase the maximum solids flux for a given diameter of the annular bed of solids, the height of the bed may be increased. The larger head of solids provides a higher pressure drop in the annular bed, and hence a higher solids mass flux can be achieved before incipient fluidization conditions are reached. The influence of the diameter of the annular bed is illustrated in the fraction of auxiliary gas which bypasses through the annular bed. Consider a certain solids mass flux and gas superficial velocity in the riser. These conditions create a certain pressure drop in the independent side of the circulation loop, and hence a pressure drop is established in the annular bed in order to close the loop. This fixes the relative velocity according to Eq. ( 1). If the annular bed diameter is increased, the downward solids velocity necessary to deliver this solids mass flu x to the riser will decrease, and, as a result, the upward gas velooity required to maintain the relative gas-solids velocity in the annular bed will increase. Therefore, it may be concluded that the annular bed gas bypassing will increase when the annular bed diameter is increased.
From this consideration of the annular bed height and diameter, the rule of thumb for the design of the feeder is that the height of solids controls the maximum solids mass flux attainable, and the bed diameter dictates the amount of gas bypassing. A scenario may be envisioned in which a certain amount of gas bypassing in the annular bed is desired. This may be the case if the auxiliary gas is to be used as a heat cartier or stripping gas for the solids. Knowledge of the expected pressure drop in the annular bed would allow the determination of the required diameter. Thus, proper design of the solids feeder requires a knowledge of the pressure drop expected in the riser, cyclone, and across the orifices for a given value of riser solids mass flux and superficial gas velocity. Results of experiments in the laboratory scale riser indicate that the riser pressure drop dominates at high solids loadings, with the cyclone and orifices contributing 20 to 30% of the total pressure drop in the independent portion of the loop (for example, Fig. 4). Reliable hydrodynamic models such as that developed by our group [22] are becoming available to predict the pressure drop in the riser for design purposes. Such models when combined with the experimental work presented here, provide design tools for a promising new type of solids feeder device for CFB risers.
5. List of symbols AP dp G~ L Q~u~ Qo Uo v
pressure drop (Pa) average particle diameter (p.m) solids circulation rate (kg/m 2 s) length (m) auxiliary gas flow rate (sL/s) orifice gas flow rate (sL/s) riser gas superficial velocity ( m / s ) velocity (m/s)
Greek letters
e tz p d~
voidage viscosity (Pas) density (kg/m ~) sphericity
Subscripts
p pb g r
particles packed bed gas relative
6. Acknowledgement The Authors wish to acknowledge Mr Dan Fantini for his technical support in the design and construction of the apparatus. This work has been funded by the Natural Sciences and
T.X Pugsley er al. / Powder Technology 88 (1996) 123-131 E n g i n e e r i n g R e s e a r c h C o u n c i l o f C a n a d a operating grant and the U n i v e r s i t y o f C a l g a r y .
References [ 1] R.J. Dry and RD. La Nauze, ChorL Eng. Prog.. 86 (1990) 3 I. [2] R M. Comractorm~d J. Chaouki, in P. Base, M. Horio. and M. H~atani ( eds.), Circulating Fluidized Bed Technology IlL Pergamon, Toronto, 1991, p, 39. [3] A, Avidan and R, Shinnar, Ind. Eng. Chcm, Bes,. 29 (|990) 931, [4l D. King, in A. Avidan (ed.), CircRlating Flaidlzed Bed Technology IV, American Institute of Chemical Engineers. New York. 1994, p. 500. [5] R.M. Contractor. G.S. Patience, D.L Gameu, H.S. Horowite. G.M Sis[er and H.E, Bcrgna. in A, Avidaa (ed.)• Circulating Fla[dizedBed Technology IV. American Institute of Chemical Engineers, New york. 1994, p. 387, [6 ] G.S. Patiene.e, J, Chaouki and B.P.A, Gram[jean. Powder Technol., 61 (1990) 95. [7] J. Happel, Ind. Eng. Chem..41 (1949) Ileal. [ 8] BJ. Milnc, F. Bermti and L.A. Behle, in O.E. potter a~d D.J. NicHin (eds.). Fluldization VII, E~gineering Fom'~lation. New "~ork. 1992. p. 235. {9] B J, Harris and J,F, Davidson, in O.E, Pot~'.r aad DJ, Nicklin (eds,), Fluidization VII, E ngianering Foundation. New ~ ~;,'k, 1992, p. 219. [10] T. Hirama, H. Takeuchi, S Chiba, and T. !'idba, in p. l ~ l :~ d J.F. Large (eds.). Circulating Fluidized Bed ",¢che~lagy li, :'.~,~anon, Tol~anto, 1998, p. 511. [11] L Fusey, M.A.$c. Thesis, The University of British Columbia, Vancouver, BC, Canada, 1985. [ 121 BJ. Milne, Ph,D. Dissertation, The University of Calgaff. Calgary, Alta, Canada, 1993. [ 13] S.B. Reddy KalTi and T.M, Knowlton, in O.E. Porter and D,L Nicklin ( eds. ). Flaidizmion VII, Engineering Foundation, New York, 1992. p, 345, [ 14] K,-E, Wit|h, in Flu[dization VIII Preprints, Bnginee~ng Foundalion, New York. 1995, p. 105. [ 15] T.M. Knowlton, in P. Bssu and J.E Large (eds.). Circulating Flaidized Bed Technology II, Pergamon, Toronto, 1988, p. 31. [ 16] L. Massimilla, V. Betta and C, Della Rocca, AICh£ 1., 7 [ 1961) 502.
131
[17] D.R.M. Jones amd I.F. Davidso~ RheoI. Acta, 4 (1965 ) 180. [ 18] M.R. Judd end P.D. Dixon. AIChE Syrup. Set. 74, (176) (1978) 38. [ 19] T.M. Take,shita, M, Hamda~ H. Yosbim~ R, Ym~azaki and G. )imbo, Int. Chem. Eng.. 30 (1990) 268. [20] ~.J. Chin, RJ. Mmz. and J.R. Grace, Powder TechnoL. 25 (19'30) 191. [ 21 ] M.J. Rhodes and p. Lausma~n, Can. J. Chem, Eng., 70 (1992) 625. [22] T.S. Pugsley and F. Befruri, Powder Technol.. (1996) m:c~ed for peblicalion [23] R, Bader, J, Findlay and T,M. Knowllon, in P. Basu ~ J.F. [Ah']~g [eds, ), Grcu[ating Fluid[zeal Bed Technology II, Pergamon. Toronto, 1988, p. 123. [24] E.-U. Hinge, D, Ransaor tnd J, Wet|her, ia P. Basu and J.F. Larl~ ( cds. ), Circulating Flaidlznd Bed Technology II, Pexg~aoa, Teeon~o, 1988, p. 165. [25] M, Horio, K. Mo~ishita. O. Tachibe~a and M. Murata. in P. Baste and J,F, Large (eds.), Cb~ularing Flaid~zed Bed Technology II, ~ , Toronto, 1988, p. 14"/. [26] U, A~ena, A. Canallaml& L. Massin~lla and D. Piroz~,in P. Barn ami J.F. Large ( eds. ) •Circ idaflng Fla idized Bed Technology II, Pe~amon, Toronto, 1988, p, 223. [27] C.M.H. B ~ t o n and J.R. Glace, ChemL En&. SoL, 48 (1993) 2565. [28] P. Bodeleio, Y. Molodtsof and A. Delebarle, in A.A. Avidem ( ed, ) ,Circulating Flaidized Bed Technology/V. American Insril~.~:of Chetm,.,~alEngineer, New Ymk, 1994, p. 118. [29] M. Schnitzlem and H. Weinstein. in P. Base and |.1~. Large (eds.), Circular[rig Fluidized Bed Technology II, PeqpJnon, Torr~o, 1988, p. 205. [30] M J. Rhodes, X.S, Waag, H. Chen8 ami T. Hirama, Chem- Eng. Sc/., 47 (1992) 1635. [ 3 [ ] D.-R. Bai, Y.Jin, Z.-Q. Yu and J.-X. ~ Powder Technot, 71 (1992) 51. [32] M. Louge, DJ. IAscher and H. Chang, Powder Technot. 62 [ 1990) 269. [33] A. Miller and D. Gidespow.AIChEJ., 38 [ 1992) 1801. [34] R.Wong, M.Sc. Thesis, TheUnivemityofCalgevJ, Alta.C~ada, 1991. [35] S. Mort. y. Yah, K. Kato, K. Matsuhexa ami D. LiQ. in P. Ba~u, M. Horio and M. Hasatani (nds.), Circulating Finldlzed Bed Technology III. Pergamon. Toronto, 1991, p. 113. [36] M. A.-zi, p. Tudier, J,F. Lar~ am[ J.R. Bernaxd. in P. Basu, M. Horio amd M. Hasatani (eds.), Circulating Flafitiznd Bed Tectmok~ HI, Pergamon, Tofomo, 1991, p. 189. [ 3"/] P.I. Viitanen, In& Eng. Cher~ Re~., 32 (1993) 577.
Powder
Technology88 (1996) 123-131
An innovative non-mechanical solids feeder for high solids mass fluxes in
circulating fluidized bed risers T.S. Pugsley
~, B . J . M i l n e 2 F . B e r r u t i
Deparlm~nt of Chemical Engineering. The University of Calgary, Calga~%Alia T2N IN4, Cunuda
Received 19 May 1994;revised~ November 19~5
Abstract
A non-mechanical solids feeder for circulating fluidized bed risers has been developed to achieve very high and carefully controllable solids mass fluxes in a laboratory scale experimental unit. The design involves the replacement of the traditional standpipe/L-valve configuration with an aerated annular bed of solids and an innovative radial gas distributor surrounding the base of the riser. Solids enter the riser threugh orifices drilled at the riser base. Solids mass fluxes of up to 700 kg/ra2 s for sand have been achieved in a 0.05 in diameter, 5 m tall riser at riser superficial gas velocities between 5.5 and 8.5 m/s. This paper reports results of the experimental investigation of the circulating fluidized bed pressure loop in order to understand better the operation of the solids feeder and to gain insight into proper design and scale.~p of the device. Keywords: Circulatingbeds; Non-mechanicalvalve;Highsolidsfluxes;Pressure loop
1. Intreduetlen
Circulating fluidized beds (CFBs) have been successfully used for the combustion of low-grade fossil fuels and mineral processing since the 1970s [ I ]. More recently, the CFB has also gained attention in the petrochemical industry as a catalytic reactor for the preriuction of certain specialty chemicals [2] and in the petroleum industry as the riser nnd/ur regenerator in fluid catalytic cracking (FCC) units [3,4]. While CFB combusturs operate at low solids mass fluxes ( ~ 20 k g / m 2 s), FCC risers operate in the range of 500 kg/ ilia a, and CFB catalytic reactors operate at |000 k g / m 2 s or greater. The hydredynamicsin CFB risers at these high solids mass fluxes may be radically different. While a core-annular flow structure with solids downflow next to the wall has been consistently observed in risers of CFB combustors, gas residence time distribution (RTD) studies on the du Pont butane oxidation demonstration plant reported by Contractor et al. [5 ] indicate that the predominant flow structure at high solids mass fluxes is no longer core-annulus with solids reflux. The majority of hydrodynamic studies reported in the literatuse for laboratory scale CFBs have been conducted at t Depanment of ChemicalEngineering,The Universityof New Brunswick, Fmdericton.NB E3B 5A3. Canada. 2 Departmentof ChetmcalEngineering,The Universityof Saskatchewan, S~katoon, Sask. STN 5C9, Canada, flO32-5910/9615tg.00© 1996ElsevierS¢~er.ceS.A. All rightsreserved PIIS0032-5910(96)031 Ig-X
relatively low solids mass fluxes. The standpipe height required to build up the pressure head necessary to achieve very high solids fluxes is prohibitively large in a laboratory setting, The objective of the present work was to develop a non-mechanical solids feeder device for achieving high solids mass fluxes in a laboratory scale riser that eliminates the traditional standpipe/L-valve configuration. The result is a new technology for a solids feeder applicable to both pilot and industrial scale applications requiring high and carefully controllable solids mass fluxes.
2. Experimental 2.1. Apparatus
A schematic of the experimental CFB equipped with the non-mechanical solids feeder is given in Fig. I. The riser section is constructed of a series of 5 cm i.d. flangedplexiglass sections. Any number of available sections may be connected so that the riser height can vary from a total o r s to 10.5 m . The first 1.8 m of riser length is encased in a 20 cm i.d. plexiglass vessel. This vessel may be filled with solids to a height o f 1.2 m. At the bose o f the riser, which sits inside the annular bed of solids, eigh~ equally-spaced 1 cm orifices are drilled around the riser circumference. Fluidizing gas to the
124
TS. Pugsley el aL / Powder Teclmology ~8 (1996) 123-131
,clone
leg
lq
Fig.2. Detailedscaledrawingof the radialdistributor( 1:3scale).
Fig. I. Schematicof the expedmcnlalCFBequippedwiththe non-mechanical solid~feeder {notto scale). system is split between two streams, the riser gas and the auxiliary, gas streams, both individually controlled by rotameters. The auxiliary gas stream flows through the innovative radial gas distributor, and into the annular bed of solids surrounding the riser. A fraction of the gas may tend to flow upward through the annular bed, but the majority of thc gas follows the path of least resistance through the orifices drilled at the riser base. The drag of the gas on the solids as it flows through the orifices carries the solids from the annular bed and into the riser. The solids are then transported vortically in the riser by the high velocity riser gas stream which is simply the sum of the riser gas plus the fraction of the auxiliary gas flowing through the orifices. The gas-solids suspension leaves the riser via an abrupt exit configuration. Two cyclones attached in series separate the gas and solids, and the solids arc transferred by a return leg to the top of the annular bed of solids. A scale drawing of the radial distributor cross-section is shown in Fig. 2. The distributor consists of a 20 cm i.d., 5 cm long plexiglass section with 1000 regularly spaced holes of 0.5 mm diameter drilled along its circumference. The 1000 holes are divided among 5 rows of 200 holes each, and each
row is staggered, resulting in a regular triangular pitch between any three holes. A 300 mesh (75 p.m aperture) stainless steel wire mesh is spot welded to the inside and outside walls of the distributor to prevent loss of fines into the windbo×. The entire distributor arrangement is connected by flanges to the annular bed of solids to facilitate removal for cleaning or design modifications, if necessary. The auxiliary ga~ is introduced at four, equally spaced locations around the circumference of the windbox to further ensure proper gas distribution. These locations are connected by low pressure polyethylene tubing to the rotameter controlling the auxiliary gas Now. In order to prevent jet penetration through the windbox and hence maldistribution of the gas, small plexiglass deflectors are attached to the outer wall of the windbox at the four locations where the auxiliary gas is introduced. 2.2. I n s t r u m e n t a t i o n
In this work, the solids circulation rate is measured with a 3.2 mm o.d. stainless steel rod attached to a small piece of stainless steel mesh (see Fig. I ). Solids flow downward in the annular bed in a moving packed bed flow. The stainless steel rod is dropped onto the upper surface of the annular bed, well away from the wall. Hence the rod moves in the bulk solids flow so that the effect of the vessel wall, which was found by Patience et al. [6] to result in underestimation o1" the solids mass flux, is eliminated. As the original upper surface of the bed descends, the rod descends with it, and its rate of descent is measured by tracking the movement of a mark etched on the rod over a known distance. Th~. corresponds to the downward bulk velocity of the solids in the annular bed. Happel [ 7] showed that the voidage of a moving packed bed is approximately equal to that of a loosely packed bed of the same solids. Assuming the loosely packed voidage in the annular bed, and knowing the annular bed cross-sectional area along with the downward solids velocity, the total mass flowrate of solids into the riser through all of tho orifices
~LS. Pugsley et aL /Powder Technology 88 f1996) I23-131
Table I Physicalcharacterizationof LaneMountainsilicasand PropeRv Particledensity(kg/m3) Bulkdensity(kg/m3) Meandianmlerterm) Packedbed voidage MinimumfluidizaUonvelocity(m/s) Minimumfluidizationvoidage Terminalsettlingvelocity(m/s) Sphericity
2.5.80 1650 208 0.36
0.038 0.45 1.5 0.90
may be calculated. Dividingthis mass flow rate by the crosssectional area of the riser gives the riser solids mass flux. The sulids used throughout this study were Lane Mountain silica. The physical properties of the sand are summarized in Table 1. In order to ascertain the minimum fluidization velocity of this material, a fluidization test was performed in an available 10.2 cm bed. Daring these experiments, it was observed that bed expansion was negligible prior to incipient fluidization conditions. Thus it is concluded that the voidage remains constant and equal to the loosely packed bed voidage until incipient fluidization, and the calculation of solids mass flux described above is accurate. Beyond incipient fluidization, bubbles will pass through the annular bed and the descent of the stainless sleet ~od is no longer smooth, making the measurement inaccurate. Results of these experiments and those reported by Milne et al. [8 ] have proven the descending rod measurement technique to be very reliable and entirely reproducible when sand is used as the circulating solids. When using fluid catalytic cracking catalyst or other Geldart A type particles, the descending rod method will require more careful calibration, As conditions of incipient fluidization are approached, a bed of FCC catalyst expands significantly. Therefore, it is not valid to assume the voidage equal to that of a loosely packed bed of FCC catalyst. Furthermore, as the fine FCC catalyst circulates through the system, air may become trapped in the interstices, and the annular bed can he significantly aerated by this trapped gas. The circulating fluidized bed apparatus is well tapped to facilitate cxamiantion of the entire pressure loop. The riser section is equipped with 21 pressure taps when operating at a height of 5 m, and no less than 37 taps when operating at a height of 10.5 m. The taps are fabricated from 6.35 mm threaded brass tubing inserted flush to the tuner wall of the riser, and 200 mesh stainless steel wire mash is spot welded to the taps to prevent backflow of solids and clogging of the lines. Tygnn tubing connects the pressure taps to a bank of U-tube water manometers. The first tap is located 5 cm above the orifices drilled at the riser base, and is denoted as point A in Fig. 1. The subsequent 11 taps are equally spaced at 15 era and the balance are spaced at 30 era. The pressure tap at the riser exit is indicated as point B in Fig. I. The pressure drop in the annular bed is measured by three 6.35 mm o.d. brass taps inserted at axial locations of 86, 40,
125
and 10 cm above the riser orifices ( points D, E, and F, respectively, in Fig. I ). A fourth tap is immersed in the annular bed at the level of the orifices (tap G). A final 6.35 mm brass pressure tap (tap C) is Ineated 2,5 cm belqw the primary, cyclone outlet. This tap, in conjunction with tap B, moni#.o,-x the pressure drop over the exit region of the riser and the primary cyclone. The point labelled C ' in Fig. 1 is the pressure at the upper surface of the annular bed of solids. This value was not actually measured, but was instead obtained by extrapolating the pressure gradient in the annular bed between points D, E, and F to the upper bed surface.
3, Critical assessment of the non-mechanical guilds feeder design Replacing the traditional standpipe/L-valvedesign of circulating fluidized beds with the annular bed of solids and radial distributor results in a unique solids feeder device, hut the concept of an annular solids feeder is not an entirely new idea. The most recent example of an annular fluidizod bed surrounding the base of the riser is seen in the work of Harris and Davidson [9]. However, :be circulating fluidized bed of Harris and Davidson [91 still has a standpipe and slide valve to control solids flow. The annular bed behaves similarly to the fluidized bed in an overflow standpipe design. Solidsentor the bed from the standpipe and s,mply overflow in order to establish a more uniform solids distribution to the riser. The annular solids feeder designed in this work is more closely related to the systems devised by Hirama et al. [ !0] and Fusey [ 11]. The circulating fluidized bed of the latter study essentially consisted of an insert capped by a baffle and placed in an existing fluidized bed of solids. Experiments were performed with FCC catalyst and PVC resin in which the annular bed was maintained at, or just below, minimum fluidization conditions. The maximum level of solids in the annular bed was 80 cm and the maximum solids mass flux aehievod in this rig was 71 kg/m2 s. This design of the annular feeder appeared to lead to operational problems. Fusey [ ! 1] reported that the pressure at the base of the insert was the same as that at the base of the annular bed. This would seem to suggest a loss of conUol of the solids flow. Fusoy [ 1I] noted that one drawback of this design was less direct and precise control of solids mass flux. The solids feeder described in this paper is different due to the "arlial gas dis~ibutor. This design provides superior solids flow coo|xol and higher solids mass fluxes than that of Fusey [ t 1]. Furthermore, the annular bed in the work of Fusey [ 11] as well as that of Hirama e t at. [ 10] was maintained at or just below conditions of incipient fluidization. Superior flow control is also realized with the solids feeder developed in this work by operation below incipient fluidization in the annular bed. Valuable design experience for the annular solids feeder was gained from the work of Milne [ 12] which studied the inteqrnally circulating fluidized bed (ICFB). As the name
126
Z& Pugsleyet aL /Powder Technology88 (t 996) 123-131
suggests, the gas-solids suspension flowing in the riser of this circulating fluidized bed does not exit to an external cyclone, but, instead, the solids impinge on a tee-shaped ballistic separator and drop back down to the surface of a moving packed bed of annular solids. This design enables the solids to circulate within a single vessel. The solids feeder system in the ICFB consists of the combination of auxiliary gas, distributor, and orifices drilled at the riser base, However, the auxiliary gas is supplied to the bottom oftbe windbox, so that the gas has asignificant vertical component upon entering the annular bed. The innovative radial distributor design described in Fig. 2 is an advancement over the ICFB design because the auxiliary gas entering the distributor has no vertical velocity component. This results in a reduction of the fraction of gas passing upward through the annular bed for a given auxiliary gas flow, and hence a higher flow rate of solids through the orifices.
4. Experimental results After installation of the non-mechanical solids feeder on the laboratory scale circulating fluidized bed, systematic experiments were undertaken to investigate the feeder performance. It was desired to establish the ability of the feeder to achieve very high and carefully controllable solids mass fluxes and solids-to-gas mass ratios in the riser. A complete understanding of the workin~ ef the feeder required analysis of the entire CFB pressure loop under various operating conditions. 4.1. Control of solids massflux
The ability of the solids feeder to achieve very high and carefully controllable solids mass fluxes is depicted in Fig. 3. In this figure, the solids mass flux of sand and the solids flow per orifice is plotted as a limction oftbe auxiliary gas flow at 800
700
0.16 ~Uol=S n.S/o=t : 0.12 0.08 ~L 200
0
0.04 i
i
i
i
1
2
3
4
AuxililUy Gas Flow (sL/s)
Fig. 3. Solidsflowperorificeas a functionof auxiliarygas flowfor varyiag risersuperficialvelocities.Riserheight~ 5 m.
two different riser gas superficial velocities. The solids flow per orifice is simply obtained by multiplying the solids mass flux by the riser cross-sectional area, and dividing by the total number of orifices drilled at the base of the riser. The data points are represented by the symbols, while the solid lines are the result of linear regression of the data at e~ch riser gas superficial velocity. As indicated by this figure, the solids feeder is able to supply a maximum solids mass flux of 700 kg/m 2 s in the 5 m riser. This represents a solids-to-gas mass ratio of 68.6 at a gas velocity of 8.5 m/s. Comparing these values with the solids mass fluxes and solids-to-gas mass ratios listed in Table 2, it is seen that the magnitude of tbese quantities in this work surpasses nearly all of the previously reported results of other authors. While Table 2 is not a complete list of references, it is representative of the magnitude of the solids mass fluxes and solids-to-gas mass ratios found in the open literature. In this work, high solids mass fluxes and solids-to-gas mass raties in the range of the du Pont pilot scale CFB catalytic reactor ( Contractor et al., [ .5] ) have been achieved in a laboratory scale riser of only 5 m. Furthermore, the du Pont rig employed a slide valve to control the solids circulation as opposed to the non-mechanical solids feeder of this study. Contractor etal. I5] did report a solids mass flux 685 kg/m z s and a gas velocity of 5.7 m/s in their work, giving a solids-to-gas mass ratio of 100. Such conditions could not be achieved in the laboratory scale CFB due to the small riser diameter. Feeding solids at such a rate at that gas velocity led to unstable operation as solids tended to accumulate at the inlet and form slugs. The results of the regression analysis in Fig. 3 show that the data is a well-behaved linear function of the auxiliary gas flow rate. This corroborates the findings ofMilne [ 12] in the ICFB. MiMe [ 12] also concluded that the solids flow per orifice, or solids mass flux, in the ICFB was independent of riser gas velocity. Fig. 3 suggests a slightly more eomplex behaviour in the CFB of this work. At the lower solids mass fluxes, the two lines coincide proving that the solids mass flux is indeed independent of the riser gas velocity in this region. However, the slope of the regressed line is steeper at the higher gas velocity of 8.5 m/s, indicating a sharper response to the auxiliary gas flow, or equivalently, the solids mass flux. Similar behaviour was found by Fusey [ 11 ]. Therefore, the curves diverge at higher auxiliary gas flows so that the solids mass flux eventually becomes a function oftbe riser gas velocity. The seemingly contradictory findings of Milne [ 12] may be attributed to the short riser and high solids mass fluxes in the 1CFB resulting in the solids being in accelerating flow for the entire riser length. As a result, the solids mass flux was relatively insensitive to the riser gas velocity in that work. 4.'2. Pressure loop study
Results of the pressure drop across the riser in a circulating fluidized bed are now abundant in the literature. However. the cyclone, solids re-injection valve, and standpipe pressure
"~S. Pugsley et aL /Powder Technology 88 (1996] 123-131
127
Table 2 Sample of published expcriroentalwork indicatingsolids mass fluxes and solids-to-gasmass ratios achieved Height
Author
(m]
Bader etal. [23] Har~geet aL [24] Horio el al. [25] Arenaet al. [26] Brereton and Grace [27] Bodelein etal. [28] Schnitzleinand Weinstein [29] Rhodes etal. [ 30] Bal el al. [ 31] Lcuge et al. [ 32] Miller and Gidaspow [33[ Wong [34] Mori etal. [ 35] Azzi etal, [36] This work Contractor et al. [5] Viitanen [37]
12.00 8.40 2.80 6.40
9.30 10.Q0 6.50 62.0 8 DO 7.00 550 3.00 4.55 ] 1.70 5.00 27.00 39.00
Diameter (m)
G, (kg/m 2 s)
Valve
0.3050 0A000 0.0500 0.0410 0.1520 0.1440 0.1520 0.1520 0.1860 0.1970 0.0750 0.0500 0.0500 0.1900 0.0500 0.1500 1.0000
147.0 30.0 11.7 500.0 116.0 52.9 250.0 I I ! .0 140.0 125.0 32.8 126.0 150.0 150.0 700.0 590.0 485D
slide L-valve beaerfiy/aearauon butterfly/aeration L-valve L-valve butzerfiy/aeration L-valve ?? gate valve slk[e valve b~ll valve U-heed/aeration slide valve solids feeder slide valve slide valve
drop are often not considered. Along with the riser, these components comprise the CFB pressure loop and merit equal consideration, especially the standpipe, or in the case of this work, the annular bed of solids. The standpipe has been referred to in the genre of fluidizatiun as the engine that drives a circulating system (see Reddy Karri and Knowiton [ 13] ). A typical t-'ressure loop for the CFB is presented in Fig. 4 for a solids mass flux of 360 kg/lll 2 s and riser gas superficial velocity of 6 m/s. The solid symbols represent the pressure measured by the riser taps, and the open symbols the pressure measured on the solids return side of the loop (i.e. the cyclone and annular bed).
B
e-
~
•
e
5-
is2. 1-
¢
Pntma'etPt ~ m )
Fig. 4, Typical plessme loop at a solids ~ flux, G,~ 3~Okg/m2 s, and riser gas superficialvelocity, fie= 6 m/s. Riser height= 6.4 m.
Uo (m/s)
9.1 3.7 1.2 7.0 7.1 4.33 4.0 5.0 7.0 5.0 2.89 8.0 5.0 6.2 8.5 7.8 13
w,l• 13.46 6.76 8.13 59.52 13.62 10.18 52.08 1&50 16.67 20.83 9.46 13.13 25.00 20.16 68.63 63.03 31.09
Several points may be gleaned from this plot. Firstly, the highest pressure point occurs at point G, immediately upstream of tl:~ orifices and at the bottom of the annular bed of solids. At this location, the auxiliary gas is introduced through the radial distributor. The gas will then either flow through the orifices, or pass upward through the annular bed. In either case, the flow of gas results in z pressure chop, so that the pressure at point G must be the maximum in the system. The lowest pressure point in the system is seen at point C at the prir.~2.,'y cyclnne sofids exit. The pressuse at this point is only slightly above atmospheric. Fig. 4 also illustrates the pressure drop associated with flow of solids and gas through the orifices from point G to point A. For these particular operating conditions, this pressure chop is app[oximately 850 Pa. The pressure drop through the riser is also clear from this figure. At the riser base, the prassure drop is very shazp as the solids entering the riser from the annular bed arc accelerated. At approximately 2 m, the pressure gradient becomes constant until the vicinity o f the riser exit. where the abrupt exit configuration causes some donsifica.lion. A slight pressure build-up is seen in the return leg as the solids flow from the primary cyclone exit to the annular bed. This is consistent with the pressure increase associated with lean phase downflow in CFB dowocomers as described by Wirth [ 14]. Fiaally, in the annular bed o f solids, the pressure increases from point C ' to G. If the loop is considered to begin at point G, the high wessut~ point in the system, then there exists a series o f pressure chops downstream o f that point through the orifices, riser, and cyclone. A small fraction of this pressure drop is recovered in the return leg, but from a practical standpoint, most o f the pressure drop is recovered in the annular bed. Hence the annular bed serves the same purpose as a conventional standpipe by transporting solids from a low pressure point ( C ) to a high pressure point ( G ) . The pressure increase from C ' to G mast balance the sum of
T.S. Pugsley¢t al. /Powder Technology 88 (1996) 123-131
t28
a
: ~
I - o - ~ = 24o k~mZS:Uo. . . . . /s
s=
; o = 5 8.m / s
s
21" o+
i
3000
~
6O00
9000
-
- 7
12O00
15000
[hcssu.+'e(Pl. glUBC) Fig. 5. Comparison of pressure loops at a solids mass flux, G~=240 kg/m a s, and two riser g~ superficial velocities. Riser height =6.4 m the pressure drops and close the pressure loop. Using the
nomenclature of Knowlton [ 15], the pressure drop through the annular bed is the dependent portion of the pressure loop, while the pressure drops from G to A (orifices), A to B (riser), and B to C (cyclone) ma~:eup the independent portion of the loop. Therefore: A P c c , + A P c , o = APGA + A P ~ + ABc
(1)
Based on the data of Fig. 4, the orifice pressure drop is 850 Pa, the riser pressure drop is 8036 Pa, and the cyclone pressure drop is 2243 Pa, for a total independent pressure drop of I 1129 Pa. The measured pressure gain in the annular bed is 10903 Pa which closes the loop to within 2.1%. This excellent closure indicates that the discrepancy associated with neglecting the pressure increase in the return leg is only slight. Hence, in subsequent analyses of the pressure loop, the pressure increase in the return leg is ignored. The pressure drop in the riser due to solids acceleration and hydrostatic head constitutes over 70% of the total pressure drop in the system, while the pressure drop through the orifices is only approximately 7.5% of the total. Fig. 5 compares the pressure loops for a solids mass flux of 240 kg/re 2 s at two riser gas superficial velocities. For the perpose of this comparison, the intermediate pressure readings in the riser have been eliminated, and points A and B joined by a straight line. The result of decke~iag ~ e ga~ velocity at a fixed solids mass flux is an increase in the total riser pressure drop due to increased solids suspension density, and an increase in the cyclone pressure drop. In order to compensate for the increases, the pressure drop through the annular bed of solids must also increase, and this is reflected by the larger pressure gradient in the bed at the riser gas superficial velocity ofS.0 m/s. The pressure gradient diatates
the pressure at point G, and results in a pressure drop through the orifices that is less for the case of lower riser gas superficial velocity. Milne [ 12] presented a correlation for predicting the solids flow per orifice in the ICFB as a unique function of orifice pressure drop. In Fig. 5, the solids flow per orifice is the same for both curves, but the pressure drop across the orifices changes w~th riser gas superficial velocity. This suggests that the flow of solids through the orifices is not a unique function of the orifice pressure drop in this system. This is attributed to the inject ion of the solids into a closed circulation loop which creates a certain back pressure at the level of the orifices. The expression of Milne [ 12] was derived by modifying equalions for the discharge of solids from aerated bins and hoppers to atmosphere. The open literature abounds with such studies (for example [16-19]) presenting experimental results ~nd correlations for the gas and solids flow as a function of the orifice Jpressure ¢~op. A slightly different study was performed by Chin nt at. [20] where solids were discharged through multiple orifices from the inner surface of an aerated annular bed of solids. Their experimental set-up was very similarto the annular bed solids feeder described in this paper. However, as tha solids flowed from the annular bed, through the orifices, and into a central tube, they did not enter a circulation loop as is the case in our work. Instead, they were , ,oply jetted to atmosphere and collected in a receiver at the bottom of'he tube. Based on the data collected in the present study, it appears that the work of Chin et at. [ 20], as well as all other aforementioned references dealing with solids and gas discharge from orifices [ 16--19], cannot be extended in a straightforward manner to analyze the performance of the new solids feeder. This is because these studies jetted solids and aeration gas through orifices against a constant back pressure ( aUnospheric ).The success of Milne [ 12] in eorral~ing the solids flow per orifice as a function of orifice pressure drop is attributed to the solids acceleration in the ICFB riser. This effectively created a constant back pressure against which solids were injected by removing the dependence on the riser gas velocity. The influence of the back-pressure introduced by the closed circulation loop in the CFB is not yet clear. It seems likely that the pressure field around the orifices in the annular bed is affected by the back pressrun. Experiments arc planned to investigate this hypothesis by traversing the tap at point G to establish the pressure field for various riser operating conditions. Table 3 summarizes the operating conditions of solids mass flux and riser gas superficial velocity investigated, along with the corresponding auxiliary gas flow-rate, Qa=~,the flowrate of gas throagh the orifices, Qo, and the orifice pressure drop, APo. "fins table does not indicate any clear trend of solids mass flux as a function of orifice pressure drop. What is important to note from this table, however, is the relatively low orifice pressure drop at high solids flows. This pressure drop is much less than that reported for flow of solids through L-valves. For example, Rhodes and Lausmann [21] measured L-valve pressure drops between 4 and 10 kPa, for riser
129
T S. Pugsley et al. I Powd~r Techno[ogy g~ (1996) 123-131
Table3 Summaryof orificepressuredropdataand orificegas flowat venoussolids massfluxesand riser supedicialgas velocities Gs (kg/m2s)
Uo (m/s)
Q,, (sL/s)
Qo (sL/s)
Aeo
8.1 10.5 19.8 25.0 150
2.6 2.7 3.4 3.6 5.5
240
5.8
5.0 6.0
036 0,39 0,39 0.69 0.93 1.42 1.69 1.78
186.8 31.1 112.1 623.0 485.0 996.7
240 360
0.50 0,45 oA5 0.98 0,98 1,45 1.79 1.79
(Pa)
6"12.8
847.2
solids mass fluxes of only 5 to 50 kg/m 2 s. Ruddy Karri and Knowlton [ 13] reported an L-valve prossure drop of 20 to 30 kPa in their test rig. The standpipe height was 8.5 m. The low pressure drop across the orifices is important for achieving high solids Ioadings in the riser. The requirements for pressure recovery in the annular bed due to the orifices is low, which allows for a larger pressure drop or, equivalently, higher solids ~aass flux in the riser. 4.3. Annular b e d pressure drop a n d auxiliary gas bypassing
The pressure drop through the moving packed bed is a function of the relative velocity between the downward flowing solids and the fraction of auxiliary gas flowing in the bed. This may be described by the Ergun equation using the relative velocity for moving packed bed flow, v; AP _ ( i - e,~) o'.v~ r- z p,.v I T 2 -b 1.75i - ¢.~, r - .~.f=.~. --=iSU
(2)
where the relative gas velocity is calculated from: v,=vp-v 8
(3)
with velocity positive in the downward direction. Therefore, the fraction of auxiliary gas that passes through the annular bed of solids, vs, is determined from the pressure drop in the dependent portion of the loop. F_,q.(2) bulds until the point of incipient fluidization, where the pressure drop through the bed becomes constant with
changes the riser solids mass flux. Thus the feeder design ig consistent with the valve mode of opuratiou of non-mechaw ical valves described by Knowlton [ 15]. Eq. (2) also indicates that for a movingpacked bed, incipient flaldizarion conditions are reached when the relative velocity is equal m the incipient fluidizatiou velocity, as opposed to the superficial gas velocity in a stationary packed bed. Therefore, based on Eq. (3), if the downward solids velocity is much larger than the upward gas velocity, the annular bed may become fluidized even though the gas velocity in the annular bed is very small. A large downward solids velocity is indicative of a high riser solids mass flux, which consequently requires a larger pressure drop in the annular bed in order to close the pressure loop. In order to maximize the solids mass flux attainable in the system, it is desirable to achieve a high solids mass flux with as little auxiliary gas flow as possible. This highlightsthe importaxce of the radial distributor design in this work. The auxiliary gas enters with a zero vertical velocity component, and has a greater tendency to pass through the orifices. Therefore, more solids axe dragged through the orifices at a lower au~liasy gas flow, and less gas is available potentially to pass upward through the annular bed. With velocity defined as negative in the upward direction, it is seen from Eq. (3) that if mum gas flows upward for a given downward solids velocity ( or given solids mass flux), the relative velocity is higher, and the bed fluidizas earlier to give a lower value of maximum solids mass flux. The success of the radial distributor in reducing the flow of gas through the annular bed is illustrated in Fig. 6, which plots the flow of gas through the orifices versus the auxiliary gas flow. With the exception of the data points represented by the open circles, better than 95,% of the auxilimy gas flows through the orifices in all cases. Thus the fraction of gas 4.0
S.5 3.0
:~ zs
increasing velocity. This sets the upper l i m i t on the riser solids
mass flux. Further increases in the auxiliary gas flow can no longer raise the solids mass flux because the maximum pressure drop has been realizod in the annular bed. As the pressure loop must always close, the condition of incipient fluidization corresponds to the maximum solids mass flux that can he achieved in the system. The previous studies using fluidized bed solids feeders [10,11] maintained the bed at incipient fluidizatiou conditions. Thus, the bed could not be used to control the solids flow rate and the solids mass flux was simply controlled by the riser gas velocity. Since die solids feeder described in this work is operated below minimum fluidizatiou conditions, varying the auxilimy gas flow
].o o.s o,o 0.0
i
i'
i
,
,
i
i
0.5
1JO
1.5
ZO
2.5
3.O
3.S
Fig.6. Frncdonof auxi]ia~gas passing0=roughII~ odfic~s.
q,O
130
Z S. Pussley et aL /Powder Technology88 (1996) 123-131
diverted through the orifices with the radial distributor design is greater than that in the ICFB of Milne [ 12 ], where orifice gas was approximately 85% of the auxiliary gas flaw. The points specially designated by open circles indicate that only roughly 70 to 80% of the auxiliary gas flows through the orifices. These points correspond to a range of solids mass fluxes of 8 to 25 kg/m 2 s and riser gas superficial velocities of 2.6 to 3.6 m/s. These operating conditions are typical of the turbulent fluid~zation regime. Even though the net solids mass flux is low, the dense bed characteristic of the turbulent regime forms at the riser base and results in a large total riser pressure drop. As a result, the pressure gradient in the annular bed of solids must be large in order to close the pressure loop. As the solids mass flux is low, the downward particle velocity in the annular bed in Eq, (3) is low. but, in order to generate the required pressure drop, the relative velocity between solids and gas must be high. Therefore, the upward gas flow through the annular bed in Eq. (3) is necessarily very high. In summary, the fraction of auxiliary gas that passes through the orifices (or upward through the annular bed) is a result of the dependent pressure drop in the annular bed of solids required to close the pressure loop. The radial distributor provides a design that is able to direct 95% or more of the auxiliary gas flow through the orifices. Such a design maximizes the solids mass flux. A lower upward gas velocity in the annular bed (Vg in Eq. (3)) allows for a very high downward solids velocity without fluidizing the annular bed, and hence, a very high solids mass flux in the riser. 4.4. Design and scale-up o f the solids feeder
Only one feeder of the dimensions given in Fig. 2 was constructed for this work. However, the study of the pressure loop provides valuable insight into the proper design and scale-up of this device for CFB risers of different sizes. In order to increase the maximum solids flux for a given diameter of the annular bed of solids, the height of the bed may be increased. The larger head of solids provides a higher pressure drop in the annular bed, and hence a higher solids mass flux can be achieved before incipient fluidization conditions are reached. The influence of the diameter of the annular bed is illustrated in the fraction of auxiliary gas which bypasses through the annular bed. Consider a certain solids mass flux and gas superficial velocity in the riser. These conditions create a certain pressure drop in the independent side of the circulation loop, and hence a pressure drop is established in the annular bed in order to close the loop. This fixes the relative velocity according to Eq. ( 1). If the annular bed diameter is increased, the downward solids velocity necessary to deliver this solids mass flu x to the riser will decrease, and, as a result, the upward gas velooity required to maintain the relative gas-solids velocity in the annular bed will increase. Therefore, it may be concluded that the annular bed gas bypassing will increase when the annular bed diameter is increased.
From this consideration of the annular bed height and diameter, the rule of thumb for the design of the feeder is that the height of solids controls the maximum solids mass flux attainable, and the bed diameter dictates the amount of gas bypassing. A scenario may be envisioned in which a certain amount of gas bypassing in the annular bed is desired. This may be the case if the auxiliary gas is to be used as a heat cartier or stripping gas for the solids. Knowledge of the expected pressure drop in the annular bed would allow the determination of the required diameter. Thus, proper design of the solids feeder requires a knowledge of the pressure drop expected in the riser, cyclone, and across the orifices for a given value of riser solids mass flux and superficial gas velocity. Results of experiments in the laboratory scale riser indicate that the riser pressure drop dominates at high solids loadings, with the cyclone and orifices contributing 20 to 30% of the total pressure drop in the independent portion of the loop (for example, Fig. 4). Reliable hydrodynamic models such as that developed by our group [22] are becoming available to predict the pressure drop in the riser for design purposes. Such models when combined with the experimental work presented here, provide design tools for a promising new type of solids feeder device for CFB risers.
5. List of symbols AP dp G~ L Q~u~ Qo Uo v
pressure drop (Pa) average particle diameter (p.m) solids circulation rate (kg/m 2 s) length (m) auxiliary gas flow rate (sL/s) orifice gas flow rate (sL/s) riser gas superficial velocity ( m / s ) velocity (m/s)
Greek letters
e tz p d~
voidage viscosity (Pas) density (kg/m ~) sphericity
Subscripts
p pb g r
particles packed bed gas relative
6. Acknowledgement The Authors wish to acknowledge Mr Dan Fantini for his technical support in the design and construction of the apparatus. This work has been funded by the Natural Sciences and
T.X Pugsley er al. / Powder Technology 88 (1996) 123-131 E n g i n e e r i n g R e s e a r c h C o u n c i l o f C a n a d a operating grant and the U n i v e r s i t y o f C a l g a r y .
References [ 1] R.J. Dry and RD. La Nauze, ChorL Eng. Prog.. 86 (1990) 3 I. [2] R M. Comractorm~d J. Chaouki, in P. Base, M. Horio. and M. H~atani ( eds.), Circulating Fluidized Bed Technology IlL Pergamon, Toronto, 1991, p, 39. [3] A, Avidan and R, Shinnar, Ind. Eng. Chcm, Bes,. 29 (|990) 931, [4l D. King, in A. Avidan (ed.), CircRlating Flaidlzed Bed Technology IV, American Institute of Chemical Engineers. New York. 1994, p. 500. [5] R.M. Contractor. G.S. Patience, D.L Gameu, H.S. Horowite. G.M Sis[er and H.E, Bcrgna. in A, Avidaa (ed.)• Circulating Fla[dizedBed Technology IV. American Institute of Chemical Engineers, New york. 1994, p. 387, [6 ] G.S. Patiene.e, J, Chaouki and B.P.A, Gram[jean. Powder Technol., 61 (1990) 95. [7] J. Happel, Ind. Eng. Chem..41 (1949) Ileal. [ 8] BJ. Milnc, F. Bermti and L.A. Behle, in O.E. potter a~d D.J. NicHin (eds.). Fluldization VII, E~gineering Fom'~lation. New "~ork. 1992. p. 235. {9] B J, Harris and J,F, Davidson, in O.E, Pot~'.r aad DJ, Nicklin (eds,), Fluidization VII, E ngianering Foundation. New ~ ~;,'k, 1992, p. 219. [10] T. Hirama, H. Takeuchi, S Chiba, and T. !'idba, in p. l ~ l :~ d J.F. Large (eds.). Circulating Fluidized Bed ",¢che~lagy li, :'.~,~anon, Tol~anto, 1998, p. 511. [11] L Fusey, M.A.$c. Thesis, The University of British Columbia, Vancouver, BC, Canada, 1985. [ 121 BJ. Milne, Ph,D. Dissertation, The University of Calgaff. Calgary, Alta, Canada, 1993. [ 13] S.B. Reddy KalTi and T.M, Knowlton, in O.E. Porter and D,L Nicklin ( eds. ). Flaidizmion VII, Engineering Foundation, New York, 1992. p, 345, [ 14] K,-E, Wit|h, in Flu[dization VIII Preprints, Bnginee~ng Foundalion, New York. 1995, p. 105. [ 15] T.M. Knowlton, in P. Bssu and J.E Large (eds.). Circulating Flaidized Bed Technology II, Pergamon, Toronto, 1988, p. 31. [ 16] L. Massimilla, V. Betta and C, Della Rocca, AICh£ 1., 7 [ 1961) 502.
131
[17] D.R.M. Jones amd I.F. Davidso~ RheoI. Acta, 4 (1965 ) 180. [ 18] M.R. Judd end P.D. Dixon. AIChE Syrup. Set. 74, (176) (1978) 38. [ 19] T.M. Take,shita, M, Hamda~ H. Yosbim~ R, Ym~azaki and G. )imbo, Int. Chem. Eng.. 30 (1990) 268. [20] ~.J. Chin, RJ. Mmz. and J.R. Grace, Powder TechnoL. 25 (19'30) 191. [ 21 ] M.J. Rhodes and p. Lausma~n, Can. J. Chem, Eng., 70 (1992) 625. [22] T.S. Pugsley and F. Befruri, Powder Technol.. (1996) m:c~ed for peblicalion [23] R, Bader, J, Findlay and T,M. Knowllon, in P. Basu ~ J.F. [Ah']~g [eds, ), Grcu[ating Fluid[zeal Bed Technology II, Pergamon. Toronto, 1988, p. 123. [24] E.-U. Hinge, D, Ransaor tnd J, Wet|her, ia P. Basu and J.F. Larl~ ( cds. ), Circulating Flaidlznd Bed Technology II, Pexg~aoa, Teeon~o, 1988, p. 165. [25] M, Horio, K. Mo~ishita. O. Tachibe~a and M. Murata. in P. Baste and J,F, Large (eds.), Cb~ularing Flaid~zed Bed Technology II, ~ , Toronto, 1988, p. 14"/. [26] U, A~ena, A. Canallaml& L. Massin~lla and D. Piroz~,in P. Barn ami J.F. Large ( eds. ) •Circ idaflng Fla idized Bed Technology II, Pe~amon, Toronto, 1988, p, 223. [27] C.M.H. B ~ t o n and J.R. Glace, ChemL En&. SoL, 48 (1993) 2565. [28] P. Bodeleio, Y. Molodtsof and A. Delebarle, in A.A. Avidem ( ed, ) ,Circulating Flaidized Bed Technology/V. American Insril~.~:of Chetm,.,~alEngineer, New Ymk, 1994, p. 118. [29] M. Schnitzlem and H. Weinstein. in P. Base and |.1~. Large (eds.), Circular[rig Fluidized Bed Technology II, PeqpJnon, Torr~o, 1988, p. 205. [30] M J. Rhodes, X.S, Waag, H. Chen8 ami T. Hirama, Chem- Eng. Sc/., 47 (1992) 1635. [ 3 [ ] D.-R. Bai, Y.Jin, Z.-Q. Yu and J.-X. ~ Powder Technot, 71 (1992) 51. [32] M. Louge, DJ. IAscher and H. Chang, Powder Technot. 62 [ 1990) 269. [33] A. Miller and D. Gidespow.AIChEJ., 38 [ 1992) 1801. [34] R.Wong, M.Sc. Thesis, TheUnivemityofCalgevJ, Alta.C~ada, 1991. [35] S. Mort. y. Yah, K. Kato, K. Matsuhexa ami D. LiQ. in P. Ba~u, M. Horio and M. Hasatani (nds.), Circulating Finldlzed Bed Technology III. Pergamon. Toronto, 1991, p. 113. [36] M. A.-zi, p. Tudier, J,F. Lar~ am[ J.R. Bernaxd. in P. Basu, M. Horio amd M. Hasatani (eds.), Circulating Flafitiznd Bed Tectmok~ HI, Pergamon, Tofomo, 1991, p. 189. [ 3"/] P.I. Viitanen, In& Eng. Cher~ Re~., 32 (1993) 577.