high density solids downflow gas-solids reactors density solids downflow gas-solids reactors by ......
TRANSCRIPT
High Density Solids Downflow Gas-Solids Reactors
by
Weidong Liu
Graduate Program in Engineering Science
Department of Chernical and Biochemical Engineering
Submitted in partial fulfiUment of
the requirements for the degree of
Master of Engineering Science
Faculty of Graduate Studies
The University of Western Ontario
London, Ontario
April, 1999
Q Weidong Liu 1999
National Liirary 1*1 ofCanada Bibbthèque nationale du Canada
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Experiments were c d out in a specialiy designed 5 m high, 0.025 m ID high
density solids downfiow gas-soi& fluidized bed to measure the axial pressure gradient
profiles dong the test column and the actual solids holdup in the M y developed region.
FCC particles with a mean particle diameter of 65 j.m and a density of 1550 kg/m3, a
Geldart (1973) A powder, was used.
For CO-current gas-solids downflow, a particle acceleration region and a M y
developed region were identifid dong the column h m the pressure gradient pronles. In
the fully developed region, the apparent solids holdup calculated fiom the pressure
gradient agreed weïï with the actuaï solids holdup measured by a pair of pinch valves
under velocities less than 5.6 d s , but underestimated it at higher gas velocities due to the
increased wall fiction loss. Two different flow regimes were observed in the developed
region, a constant and high density pseudo-aggregative flow regime under low gas
velocities and a reducing density pseudo-particdate flow regime under high gas
velocities, with a boundary between U, = 0.5-1.3 mls, which is the critical gas velocities
deked as U, U, cm be determined either nom the measurement of the solids holdup
in the pseudo-aggregative regime or the Merential pressure fluctuation. Kigh density
downflow operation is defined as operation in the pseudo-aggregative flow regimey where
particle velocity remains constant under ail solids flux and gas velocity conditions and
where the slip velocity is very high, with significant particle agglomeration. A solids
holdup as high as 10% has been achieved in this operating regime. in the more dilute
pseudo-particdate flow regime, the gas-particle slip velocity remains constant and no
particle strands and large particIe clusters were observed. The particle velocity was found
to increase Iinearly with the gas velocity given the constant slip velocity. Conseq~ently~
the solids holdup decreased with increasing gas velocity in this regime, as reported
previously in other riser and downer systems. Cornparison of the resuits obtained here
with those h m an upfiow riser shows inhermt smiilarities between the two gas-solids
CO-cment fbw systemsems
For gas upward-solids downward counter-curent fluidized flow. the flow
development and fiction are discussed in relation to the pressure gradient profiles. The
actual solids holdup measured by a pair of pinch valves and the apparent solids holdup
calculated h m the pressure @-ents am compareci for diffient operaîing conditions-
Based on the changes of the mean particle velocity and the particle slip velocity. the
particle agglomeration was studied. Choking is discussed in relation to both riser and
counter-curent operation. The operable maximum superficial gas velocity and solids
flw in this system for FCC were experimentally determinecl.
The cornparison of the high density downfïow and the counter-cunent flow
regimes with the upflow flow regime were made by using the differential pressure
fluctuations and the particle slip velocity. The flow regimes in the CO-current high
density downflow and the counter-current flow are expected to exhibit the same types of
hydrodynamic behaviour of fast fluidization and pneumatic transport regimes in the
upflow system.
Finaily, an u - e d overall flow regime diagram is proposed This map first gives
a general picture of fluidization which includes all types of fluidized beds. A clear
"operating window" for FCC particles is proposed. The new unifieci flow regime
diagram extends our current knowledge to wider operating ranges.
Keywords: High Density, Downfîow Fluidized Bed, Counter-Cment Flow, Flow
Regllne, Hydrodynamics. Downer
Titie: Characterization of High Density Gas-Solids Downflow Fluidlzed Reactors
Anthors: Liu, LX- Zhu and J- M- Beeckmans
The prelimïnary research, experimental design, testing and experimental runs were
undertaken by W. Liu under the guidance of the CO-advisors J.-X. Zhu and J. M.
Beeckmans. AU drafts of this manuscript were written by W. Liu Modifications were
c d d out under the close supervision of Dr. Zhu. The finai dr& was approved for
submission to the journal Powder Technohgy by the CO-advisors
Title: Characterization of Gas Upward-Solids Downward Counter-Cment Fluidized
Flow
Authors: W. Liu, J.-X. Zhu and J. M. Beeckmans
AU portions of the experiment work were undertaken by W. Liu under the guidance of the
CO-advisors J.-X. Zhu and J. M. Beecham. AU drafts of this manuscript were written by
W. Liu. Modifications were done under the close supervision of Dr. Zhu. The nnal draft
was approved for submission to the journal Powder Technology by the CO-advisors.
The author is sincerely gratefbl to his advisor, Profasor LX- Zhu, for his
continuous encouragement, guidance, and support throughout the completion of the
project.
Much appreciation is also extended to Professor J. M. Beecham for his tutelage
and support.
Sincere tli& to aïi my coiieagues N. =mg, PM. Joiuiston, F. Wang, Y. Ma,
J. Bd, Dr. W. Huang and Mr. J. Z. Wen, who provided assistance and valuable
discussions in operating and the design of the expaimental equipment.
Financial assistance fiom the Natural Sciaces and Engineering Research Council
of Canada is gratefûiiy acknowledged.
FinalIy, special th& is extended to my wiîe for her understanding and great
support during this period of study.
TABLE OF CONTENTS
Page
-. CERTIFICATE OF EXAMINATION ..r..r.....................CC.......C................................~
*.. ABSTRACT ............................................................................................... ..............m
................................................................................................. CO-AUTHORSHIP .V
ACKNOTKLEDGMENTS ....................................................... .......... ...... -................~i
* * TABLE OF CONTF.NTS ......................................................................................... VJ
LIST OF TABLES ...............................................................e............................... .....x
LIST OF FIGURES ................................................................................. ...........-.....~ *. . NOTATION .........~*.*..**.........,........*..............*...*.* .........*.....*.....*......*............... *......xlll
CHAPTER 1 INTRODUCTION .............................. ,. ......................................... 1
............................................................................................. 1.1 Introduction 1
1.2 Objectives .............................................................................................. 4
1.3 Thesis Structure andKey R a t s ....................................... ................ 5
1.4 Bibliography ............................................................................................ 7
CHAPTER 2 EXPERIMENTAL APPARATUS AND PROCEDURES ............... .13
2.1 Description o f Solids Downflow Gas-Solids Fluidized-Bed ................... 14
.......................................................... 2.2 Description of Solids Feed System 18
2.3 Description of the Particulate Mataials ................... .. ..........titi...ti.......... -2 1
..................................................................... 2.4 Measurement of Solids Flux 2 1
...................... 2.5 Measurernent of the Axial Pressure Gradient ,... ............ .23
.............................................. 2.6 Measurement of the Acnial Solids Hoidup .24
vii
............................... 2 -7 Operathg Conditions and Experiment Procedures -25
2.8 Electrostatic Charging and Its Elimination .............................................. 28
............................................................................................ 2.9 Bibliogtaphy 28
CHAPTER 3 CHARACTERlZATION OF HIGH DENSIT'Y GAS-SOLIDS
............................. ............. DOWNFLOW FLUIDIZED REACTOR ...... 30
.............................................................................................. 3.1 Introduction 32
3.2 ExperimentaI and Operathg Procedures .................................................. 33
............................................................................. 3 -3 Results and Discussion 35
3 -3.1 Pressure Gradient Profiles and the Solids Acceleration
Length ............................................. ........................................... 35
3.3.2 Cornparison between the Actual and
the Apparent Solids Holdups .............................. ........ .............. 38
3 -3.3 Solids Holdup. Particle Velocity and Slip Velocity
in the Fully Developed Region ....................................................... 39
3.3.4 Defhition of High Demity Down£iow Operation .......................... 44
.................. 3.3.5 Cornparison between Downûow and Upflow S ystems 45
3 -4 Conclusions ................................ .. .. .... 3 -5 Bibliography .............................................................................................. 49
CHAPTER 4 C"'C"REA'M0N OF THE GAS UPWARD-SOLIDS
DOWNWARD COUNTER-CURRENT FLUIDIZED FLOW ..................... .64
.............................................. ....................... 4.1 Introduction ....................... 66
.................................... .......... 4.2 Experimental Apparatus and Procedures ... 67
.............................................................................. 4.3 Results and Discussion 69
.............................. 4.3.1 Observation ... 4-32 Flow Development and Friction Loss ................ ... .................... 69
............................. 4.3 J Soli& Holdup in the FuiIy Developed Regio a 73
4.3.4 Particle Velocity and Gas-Solids Slip Veloci ty. ............................ -74
............................................................................................... 4.4 C0nc1USi0n.s -77
4.5 Bibliography ............................................................................................... 80
CHAPTER 5 CHARAC-ATION OF TKE FLOW REGIMES AND
................... . UNIFED REG- DIAGRAM GENERAL DISCUSSION 94
5.1 Co-Current Downward Flow Regimes ............ .... ............................. -95
........... 5.1.1 Pseudo-Aggregative and PseudoParticulate Flow Regimes 95
. 5 1.2 Determination of U, by Differential Pressure Fluctuations ............ 100
................................................................. 5 -2 Conter-Current Flow Regime 102
5.3 Cornparison of High Density Downflow and Counter-Current
Flow with Upfiow Flow Reginles ............ ......... ................................ 103
..................... 5.3.1 DBerential Pressure Fluctuations and Solids Holdup 103
....................................... . 5.3.2 Mean Voidage vs Particle Slip Velocity 106
.......................................... 5 -4 UnifIed Flow Regmie Diagram ............ .... 108
.............................................................................................. 5.5 Bibliography 1 IO
CHAPTER 6 . CONCLUSIONS AND RECOMMENDATIONS ...... .. ................. 1 13
................................................................................................ 6.1 Conclusions 113
LIST OF TABLES
Table Description
Table 2.1 Location ofpressure taps dong the test c0Iumn
Table 2 2 The operating conditiofls for various experiments
Table 23 The measured parameters
Table 2-4 The calcuiated parameters
Table 2.5 Valve settings for the c l i f f i t operating modes
Page
24
26
26
27
27
LIST OF FIGURES
Figure
Figure 1.1
Figure 2.1
Figure 2 2
Figure 2.3
Figure 2.4
Figure 2.5
Figure 2.6
Figure 3.1
Figure 3.2a
Figure 3.2b
Figure 3.3
Figure 3.4
Figure 3.5
Figure 3.6
Figure 3.7
Figure 3.8
Figure 3.9
Description
ûperating ranges of various fluidized beds studies in the past
Experiment operathg systern
Generalized schematic of the secondary cycIone
Schematic of the fliridïzed bed feeder
Schemanc oÎiiie fiuiâizeci Deci Îeeâer h e i
Cumulative size distribution for FCC particles measured using Brinhan Particle Size Analyzer (4, = 65 pm)
Caliiration cuve for the Ioad cell
Schema for the experimental apparatus
Pressure gradient profile dong the column at superficial
gas velociîy U, = 0.33 d s
Pressure gradient pronle dong the column at solids
flux G, = 90 kg/m2s
Cornparison of achial and apparent solids holdup
Measured soli& holdup in the M y developed region
as a fimction of U,
Solids holdup in the -y developed region
as a fimction of solids flux
Particle slip velocity as a fiinction of supadcial gas velocity
Mean particle velocity as a fùnction of superficial gas velocity
Operating Wùidow for gas-soli& CO-current dowdhw systems
Solids holdup as a function of U, in the fully developed
region for risers and downers
Page
4
13
17
19
20
22
22
53
54
55
56
57
58
59
60
61
62
Figure 3-10
Figure 4.1
Figure 4.2a
Figure 4.2b
Figure 4.3
Figure 4-4
Figure 4.5
Figure 4.4
Figure 4.7
Figure 4.8
Figure 4.9
Figure 4.10
Figure 5.1
Figure 5.2
Figure 5.3
Figure 5.4
Figure 5.5
Figure 5.6
Figure 5.7
Figure 5.8
Solids holdups vernis adjusted superficial gas velocity
in the M y developed region for both risers and downers
Schematic of the experimental apparatus
Pressure profiles along the column at Gs = 7.8 kglm2s
Pressure gradient pmnles along the column at Gs = 19.5 kg/mZs
Cornparison of actual and apparent solids holdup
Solids holdup in the M y developed region as a hc t ion of U,
SoIids hoIdup in the fûliy deveIoped region as a function of Gs
Gis-soiias counter-current fiow operating range
Mean partice velocity as a hc t i on of Gs
Mean particle velocity as a hc t ion of Ug
Mean particle slip velocity as a fûnction of Ug
Mean particle slip velocity as a function of Gs
Mean particle velocity as a hc t ion of superficial gas velocity
nie relationship between actud particle velocity and actual
pas velocity in the pseudo-particdate flow regime
The relationship between acniai particle velocity and actual
pas velocity in the pseudo-particdate fiow regïme
Particle slip velocity in both downfbw and counter-flow
The cornparison of solids holdup between CO-curent
downflow and counter-cment flow
Differential pressure fluctuation as a furiction as solids holdup
Particle slip velocity as a fiinction as solids holdup
in CO-current downflow
Unifïed flow regime diagram
NOTATION
column cross sectional are+ m2
particle diameter, pm
acceieration due to gcavity, mis2
solids flux,. kg/m2s
distance h m the top of the test column,. m
equivalent heighq. m
pressure, Pa
standard atmospheric pressure, Pa
actual pressure, Pa
rotameter reading at standard atmospheric pressure
actual rotameter re&g
the critical gas velocity,. m/s
superficial gas velocity, m/s
defhed as Ug + 0.57 m/s in the downer and CI, - 0.57 m/s in the riser, m/s
particle terminal velocity, m/s
the omet velocity for signincant solids entrainment
mean particle slip velocity, m / s
mean actual gas velocity, m/s
mean particle veiocity, mis
weight of FCC particles
weiat of FCC particles ~acked in the testiner section
Greek Letters
tolerance for pressure gradient variation, pdm2
distance between the two pressure taps, m
pressure loss due to particle acceleration, Pa
pressure loss due to suspension-to-waU fiction, Pa
pressure loss due to gas phase fiction, Pa
pressure loss due to solids phase fXction, Pa
voidage of packed FCC particles
solids holdup
apparent density, kg/m3
density of gas phase, kg/rn3
density of solids phase, k@m3
xiv
CHAPTER 1. INTRODUCTION
1.1 Introduction
Fluidized bed reactors have many distinct advantages over other gas-solids
reactors. Broadly defined, a fluidized bed is fonned when particdate materials are
partially or completely suspendeci by a flowing fluid The particles are then called
fluidized because the ffuid-particle mixture thPs produced possessa many useful physical
properties of a fluid (Davidson and Harrison 1963). It is those properties which give the
key advantages of fluidized beds: high fluid-solids contact efficïency, high heat and mass
transfer rates, uniform temperature distn'bution, easy addition and withdrawal of solids
into/fiom the fluidized beds etc. (Lim et al. 1995).
The development history of gas-solids fluidization technofogy can be divided into
two perïods. The first period was fiom its inception (in 1920s and 1940s) to late 1970s,
when the conventional fiuidized bed was invented and intensively studieb In a
conventional gas-soli& fluidized bed, gas flows upward through a bed of particulate
materials to form a dense fluidized bed. Particles essentially remain in the bed while gas
continuously passes through the system. It was füst proposed by Winkler in the 1920s
(Kunii and Levenspiel 1969) for coai gasification and then adopted by the peûoleum
industry for catalytic cracking of cmde oil in the 1940s (Jahnig et al. 1980, Squires
1986). Since then, fluidized beds have found many applications in industry (e.g., Matsen
1982)- Extensive studies were carried out in the 1960s to 1980s to characterize this
reactor, as summarized in several key reference books (hnii and Levenspiel 1969,
Davidson and Harrison 1971, Davidson et al. 1985, Geldart 1986).
The second penod starting in the mid-1970s is characterized by high velocity
fluidization or circulating fluidized beds (CFBs). In a circulating fluidized bed, solid
particles are continuously fed into and entraineci out of the reactor by high velocity gas
flow. The added benefits of circulating fluidized bed reactors include even higher gas-
solids contact efficiency and significantly reduced gas and soli& backmirriag-
Yerushalmi and his CO-workers (1976) were the first to propose the concepts of
circulating fluidized bed and fast fluidizattïon flow regirnes, aithough some earlier work
was done by the oil companies (Stemerding 1962, Van Zoonen 1962). A large volume of
research work has been cauÏed out to understand the flow characteristics inside this type
of reactor (Barati et al. 1995, Lim et al. 1995, Grace et al. 1997). Given their distinct
advantages, over 600 CFBs are now in operation mund the world for the Fischer-
Tropsch process (ShlngIes and McDonald 1988, Steynberg 1991), the FCC (Fluid
Catalytic Cracking) process (Avidan et al. 1990. King 1992) and coal combustion @ry
and La Nauze 1990, Engstrom and Lee 1991, Kulïendorff and Andersson 1986). Many
more new applications are also being considered (Contractor 1988, Contractor and
Chaouki 199 1, Zhu and Bi 1995)-
In addition to the cocunent gas-soiids upflow fluidized beds (risers), cocurrent
gas-solids downflow cimilating fiuidized beds (downers) were proposed in recent years
(Shimini et al. 1978, Gross 1983, Gross and Ramage 1983, Kim and Seader 1983,
Niccum and Bunn 1985, Berg et al. 1989, Gartside 1989, Bai et al. 199 1, Graham et al.
1991, Wang et al- 1992, Aubert et al. 1994, Roques 1994, Zhu et al. 1995, Zhu and Wei
1996, Herbert 1997, Johnston et al. 1999, Zhang et al. 1999% Zhang et al. 1999b).
Because both gas and solids travel in the direction of gravity, the flow structure inside the
downer is much more uniforni in the radial direction than in the riser. This radial
unifomity further reduces gas and solids dispersion and leads to nearly plug flow for
both phases in the downer. In addition, the flow accelerates much more quickly in the
downer since soli& are accelerated by both the gas and gravity (compared to the riser
where solids are only accelerated by the gas flow, but resisted by gravity). With these
characteristics (short contact tirne and uniforni residence time distribution), downer
reactors become more advantageous over risa reactors for reactions with very short
residence time and reactions when the intexmediates are the desirable products.
Notwithstanding the numerous advantages of the high velocity riser and downer
reactors, they suffi a common shortcorningr very low volumetric concentration (holdup)
of solids. Conventional fluidized beds are aiso called dense phase fluidized beds, while
circulating fluidized bads are regarded as dilute phase fluidized beds. Typicaüy, a
conventional fluidized bed operates with an average solids holdup of 30%-SWh. A ris-
on the other hand, only contains 1-3% solids by volume in the M y developed region.
The solids holdups achieved in downers as shown by the nported studies are even lower
(mostly below 1%). This represents a serious problem for feactions whem a high
solidslgas ratio is required, since the reaction intensity is limited by the lower solids
concentration. To overcome this weakness Bi and Zhu (1993) proposed the concept of
the high density circdating fluidized bed (HDCFB) riser. Subsequent studies on HDCFB
have shown that solids holdups as hi& as 25% cm be achieved in such a unit @ai et al.
1997, Issangya et al. 1997% 1997b, 1998) With carefiilly controlled operation.
No attempt has been made to achieve high density in a cocurrent downflow
system. In addition, gas upfiow and solids downfiow counter-current fluidization has not
been shidied. Figure 1.1 shows that within the four quadrants formed by U, as x-axis and
G, as y-axis, studies have mainly been in the first quadrant, plus very limited reports in
the M d quadrant. Therefore, it is important to study gas-soli& flow under other
conditions in order to extend our current howledge to wider operating ranges in this
operating map.
Not Possible
Not Studied Scantly Studied
Figure 1.1 Operathg ranges of var*ous fluidized beds studies in the past
1.2 Objectives
The objectives of the present study are:
(1) To build a gas-solids system, which enables the cocurrent high-density gas-solids
dowdow operation and the cornter-cuwnt gas upward-solids downward operation.
(2) To characterïze the gas and solids flow inside the cocurrent high-demsity ges-solids
downfbw fluidized bed (downer) system.
(3) To characterize the gas and solids flow inside the counter-current gas-upward/solids-
downward fluidized bed system.
(4) To iden* possible new operating regimes for gas-soiids fluidized bed systems, and
to map the "'operating windows" of those new regimes.
1.3 Tbesis Structure and Key Resuits
This thesis foUows the "mixed format" as outlined in the UWO Thesis Guide.
Chapter 2 provides the details about the experimental apparatus, the measurement
techniques and the experimental procedures. Chapter 3 presents the study resuits on the
flow characteristics in the cocurrent high-density downer, while Chapter 4 reports on the
hydrodynarnics inside the counter-current gas upward-solids downward fluidized bed
system, both in a manuscript format After that, Chapter 5 discusses the flow regimes in
the above two systems and their relationships with other operating regimes identified
previously. The new "operating wuidow" for gas-solids fluidized bed systems is mapped.
The key findings of this work and in-depth discussions are also presented. Finally,
Chapter 6 iists the key conclusions and recommendations for fiiture work.
In Chapter 3, the r d t s on the flow characteristics study in the cocu~~ent high-
density downer are reported. A particle acceleration region and a fUy developed region
were identified dong the downer h m the pressure gradient profiles. In the fully
developed region, the apparent solids holdup calcuiated from the pressure gradient agrees
weil with the actual solids holdup measured by a pair of pinch valves under velocities less
than 5.6 m/s, but underestimate it at highex gas velocities due to the increased wail
fiction loss.
Two dinerent flow regimes w m observed in the developed region, a constant and
high density pseudo-aggregative flow regime under low gas velocities and a reducing
density pseudo-particdate flow regime under high gas velocities, with a boundary
between U, = 0.5-1.3 mls. The high density downfiow operation is deked as the
operation in the pseudo-aggregative flow regime where particle velocity remaius constant
under all solids flux and gas velocity conditions and where the slip velocity is very high
with very significatlt particle agglorneration. A solids holdup as high as 10% has been
achieved in this operating regime- In the more dilute pseudo-particdate flow regime, the
gas-particle slip velocity rem& constant and no particle strands and large particle
clusters is obsewed. The particle velocity is fond to increase linearly *th the gas
velocity given the constant slip velocity. Consequendy, the solids holdup decreases with
increasing gas velocity in this regime, as reported previously in other riser and downer
systems. Cornparison of the results obtained here with those fiom an upflow nser shows
inherent similarities between the two gas-solids CO-current flow systems.
Chapter 4 discusses the flow behaviour in a gas upward-solids downwards
counter-current fluidized flow systern for the first time. The flow patterns were observed.
Particles were seen to fiow downward as an apparently dispersed suspension. Particle
recirculation at the wail was observed, espbcially at high gas velocities, where particles
flow upward occasionally and sofids holdups were seen to be higher.
Typical axial profiles of the pressure gradient were discussed to identifjr an initial
solids developing region and a iùily developed region. The experimental results indicate
that the pressure gradient provides a simple method to estimate solids holdup without
incurring large enors when the soiids flux is not higher than approximately 15 k&s for
a l l operating gas velocity in the fully developed region.
In this gas-solids counter-c~t~ent flow system, increasing gas velocity under a
given solids flux always leads ta a hear increase in solids holdup. Increasing solids flwz
at fuced gas velocity also causes an increase in the solids holdup. However, W e r
increasing the solids flux beyond some point amund 15 kglids leads to the choking
phenornena which can be used to explain the dramatîc change in soli& holdup, particle
velocity and slip velocity.
In Chapter 5, by pmviding in-depth discussions. the two flow regimes, pseudo-
aggregative and pseudo-particdate flow regimes, which were observed in aii the gas-
soiids CO-currmt downfïow experiments studied, can be determinecl either h m the
measufement of the soîids holdup or the differentiai pressure fluctuation. The cornparison
of the high density downflow and the counter-current flow r e m e s with the upflow flow
regime were made by using the Herential pressure fluctuations and the particle slip
velocity. The flow regimes in the CO-current high density downflow and the counter-
current flow are expected to exhibit the same types of hydrodynamic behaviour of f a t
fluidization and pneumatic transport regimes in the upflow system. FinaUy, an unified
overd flow regime diagram is proposed.
1.4 Bibliography
Aubert, E. D., Bansceau, D., Gauthier, T. and Pontier, R. (1994). Profiles and
Slip Velocities in a Co-Current Downflow FLuidized Bed Reactor", CircuIating
Fhidried Bed Technologv N, (eds A. A Avidan), pp. 403-405, AIChE, New York.
Avidan, A. A., Edwards, M. and Owen, H. (1990). 'Tmovative ïmprovements Highiight
FCC's Past and Future", OiI Gas J , Jan. 8,3348.
Bai, D., rssangya, A S., Zhu, J-X. and Grace, J. R (1997). "Adysis of the Overall
Pressure Balance around a High-Density Circulating Fluidized Bed", Ind. Eng.
Chm. Res., & 3898-3903,
Bai, D., Jin, Y., Yu, 2. and Gan, N. (1991). "Radial Protïles of Solids Concentration and
Velocity in a Concurrent Down£îow Fast Fîuidized Bed (CDFFB)", Circuluting
FIuidized Bed Technology Dl, (eds. P. Bani, M. Hono and M. Hasatani), pp. 157-
162, Pergamon Press, Oxford,
Berg, D. A., Bnens, C. L. and Bergougnou, M A (1989), 'Xeactor Development for the
Vitrapyrolysis Reactof', Can. .l Chem. Eng* a 69-101.
Berruti, F., Chaouki, J., Godfby, L., Pugsley, T. S. and Patience, G. S. (1995),
'TKydrodynamics of Circuiating Fluidized Bed Risers: a Revied', Cm J I.em.
Eng., 73,579-602.
Bi, H. T. and Zhu, J-X. (1993), CLStatic Instabïfity Analysis of Circdating Fluidized Beds
and Concept of High Density Risers", MChE J , 39.1272-1280.
Contractor, R. (1988). ''Butane %dation to Maieic Anhydride in a Recircuiating Solids
Riser Reactoi', Circulating Nuiditd Bed Technology l& (eds. P Basu and J. F.
Large), pp. 467-474, Pergamon Press, Toronto.
Contractor, R and Chaouki, 1. (1991). "Circulating Fluidized Bed as a Catalytic
Reactor", CircuZating FIuidried Bed Technology (eds. P Basu, M Hono, and M
Hasatani), pp. 39-48, Pergamon Press, Toronto-
Davidson, J. F. and Harrison, D. (1963), FZuidWed ParticZes, Cambridge University Press,
Cambridge, England.
Davidson, J. F. and Harrison, D. (eâs.) (1971), FZuidization, Acadernic Press, London.
Davidson, J. F., C l . R. and Harrison, D. (eds.) (1985), Fluidization, 2nd ed., Academic
Press, London.
Dry, R J. and La Nauze, R D. (1990), "Combustion in Fiuidized Beds", Chem. Eng.
Prog., July, 3 1-47.
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Combustion Technology". CirmIating Ruidked Bed TechnoIogy m, (eds. P. Basu,
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CEtAPmR 2. EXPERIMENTAL APPARATUS AND PROCEDURES
AU experiments were performed in the same cold-model solids downflow gas-
solids fluidized bed. The scheme for the exphentai apparatus, designed and
constructeci in houe, is illustrateci in Figure 2.1.
- - VALVE 2
REïüRN PIPE ID 2" (9)
RECYCLE UNE (10)
TEST COLUMN (5) ID 1' 5 m H'gh
*- - - VALVE 5
ROTAMETER (13)
1 VALVE6
Figure 2.1 Experiment operating system
The main wmponents of the solids downflow gas-solids fluidized bed included
the fouowing: a 5 m taIl plexiglass test wlumn of 0.025 m i.d (5). a 6.5 m ta11 steel
return pipe of 0.05 1 m id. (9), two 0.76 m3 solids storage tanks of approximately 1.83 m
tall and 0.76 m 0.d. (1,8), a 0.66 m ta11 feed funne1 of 0.25 m i.d. at the top and 0.025 m
i.d. at the bottom (4), a 1.35 m long viirated inclined feed pipe of 0.10 m i.d. (3). a 0.74
rn ta11 fluidized-bed feeder of 020 m i.d. (2), a steel particle recycle pipe h e of 0.032
m i.d. (IO), aprimary cyclone of 0.39 m high and 0.10 m id , a secondary cyclone of 0.26
m in height and 0.067 m i.d, and a bag house filter. The de- on how to operate this
system are desmied as foilows.
2.1 Description of Solids Downfiow Gas-Solids Fluidized-Bed
There were two pipe lines, the test h e and the particle recycle h e , and hKo
solids storage tanks at the top and the bottom. The operations of the tests were on a batch
basis. Solids flowed down h m the top tank to the bottom one during the tests in the test
lhe and were entrained up to the top tank through recycle h e after each test. Eight
pressure taps were installed dong the column, located at 0.10 m, 0.30 m, 0.50 m, 1.00 m,
2.00 m, 2.50 m, 3.75 m and 4.25 m h m the top entrance of the column, giving five
différentia1 readings between each neighbouring pair of taps (except between 1.00 m and
2.00 m, 2.50 m and 3.75 m). During these experiments, the dinerential pressures dong
the test column were rneasured by the pressure transducers. Pressure gradient d P / M was
then calculated fiom the measued differential pressure. At levels 3.00 m and 4.50 m
from the top of the c o 1 ~ , two pinch valves (6) and (7) were instaiIed to obtain the
a c t d soli& holdup in the fully developed region by collecting and weighing the solids
trapped between the two valves when they were closed simultaneously at the end of each
experirnent. The solids flux was determined by a load celi installed undmeath the
bottom storage tank which monitored the weight changes of the tank.
For the tests, beginning h m the top storage tank (1). solids Ml into the fluidized
bed feeder (2) beneath the tank. The solids feeder system, consisting of a fluidized bed
feeder (2), a viirating pipe (3) and a feeding m e 1 (4). was speciaUy designed to achieve
very smooth and hi& solids flues. The upper portion employs a fluidized bed feeder,
and an incliued vi'brating pipe to regdate the solids flowrate. By changing the level of the
small movable tray, the fluidized feeder delivers a reguiated amount of solids into the
inclined pipe. Through vi'bration, the 20. inclioed pipe (at an angle smder than the angle
of repose) M e r damps the fluctuations in solids flow. The lower portion is the feeding
funnel, within which solids were pre-accelerated by gravity before entering the test he.
Since the particles had an initial velocity close to the terminal velocity (caicdations
performed according to the method suggested by Clift et al. 1978) upon entering the
downer top, choking was avoided and solids flux up to 500 kg/m2s could be achieved by
this novel feeder.
For CO-current downfiow tests, the main gas was introduced in the funnel, and
solids and gas then flowed downward through the test column. The fiowmtes of the main
gas and the fluiâization gas to the feeder were both monitored by rotameters. To avoid the
undesirable effects of electrostatics, 0.5% by weight of commercial Larostat powder was
added to the solids. Once solids and gas feu into the bottom storage tank, the solids were
separated fiom the gas by gravity and deposited in the tank, and gas flowed to a filter bag,
where the remaining fines were collecteci before the gas entered the exhaust system. After
each run, soli& were transported by gas through the particle recycle line (10) Eom the
bottom storage tank to the top one, and the solids were separated by gravity in the two
cyclones.
For counter-current gas upwatd-soli& downward tests, the fluidization air was
introduced into the bottom storage tank and then flowed upward în the test column. In the
fùnnel on top of the test column, air was separated h m the soli& and passed to the bag
house filter for fiuther clean before exhausting the system.
Two commercial size vessels of appmximately 1.83 m height and 0.76 m
diameter wete chosen as top and bottom solids storage tanks. The volume of each tanlc is
0.76 m3. Total weight of 300 kg FCC particles, appmxhately 0.35 m3 of the packed
volume, was employed in the experiments. Packed particles only took 46% of the volume
of each tank, which ensured that there was sufficient space for the separation of solids by
gravity. For each experiment, the soli& inventory was high enough to mure at least 25
minutes of operation calculated by the following relation based on a maximum solids flux
of 400 kg/m2s:
The prirnary and secondary cyclones were designeci based on the standard Zenz-
Cyclone theory. Generalized schematic of cyclones was shown in Figure 2.2 with ail
pertinent dimensions for the secondary cyclone. In order to increase capture efficiency
and be flexible to the change of gas velocities, gates were c o ~ e ~ t e d to adjustable hinges
on the inner wall of each cyclone inlet. By controlling inlet areas, larger gas velocities at
inlet can be obtained.
The air, used for the main air in the test column and for the fluidization of solids
in the fluidized feeder (2), was supplied by a 700 kPa compressor fiom the university
physical plant. The volumetric fXowrates were monitored by individual rotameters.
Unit: mm
Figure 2.2 Generalized schematic of the secondary cyclone
The air flowrates for the fluidization of soiids in the fluidized feeder (2) were
monitored using a commen:ially avaiIabIe rotameta. whieh was cah'brated for standard
atmospheric pressure (Po=101325 Pa) and temperature (T0=293.15 K) by the
manufacturer. The rotameters for monitoring the main air in the test column were
caliirated against by a known rotameter for standard atmospheric pressure and at room
temperature. Pressure gauges were placed irnmediately upstream of the rotameters to
measure the actual pressure (PJ, which was used to conect the reading h m the
rotameters according to foilowing equation:
Pressure gauges were also used to monitor the actual pressure inside the bottom storage
tank and the top funnel. Those values are used to determine the actual air flowrates in the
test column and then superficial gas velocities by assuming that air is ideal gas. For co-
current downflow tests, the gas flowrate in the test column was calculated as the fl owrate
of the main gas plus the flowrate of the fluidization gas. As ail experiments were carried
out at room temperature, the effects of temperature for air flowrates were neglected.
2.2 Description of Solids Feed System
The solids feed system includes three parts, the fiuidized bed feeder? the inclined
pipeline and the funnel, The primary objective is to provide a very stable solids flow at
various flow rates to the test column.
The fluidized bed feeder (2), which was specially designed for these experiments,
includes two parts: the upper and lower columns. Details are shown in Figure 2.3. Solids
were packed in the upper column where a constant solids level was kept because of the
relatively small diameter of the column compared with the storage tank. When solids
dropped into the lower column tbrough a hole under which a movable plate was installe&
the solids mass flowrate was controed by the distance between the hole and the plate
which couid be adjustecf up and d o m L order to stabilize solids flowrate, particles were
fluidized in the lower column to smoot. out large disturbances. Then soli& and
Upper Column
Movable Plate e
1
Fluidization Gas Distributor
0.2 m dia I
Fluidization gas
0.7 m
lower Column
Figure 2.3 Schematic of the fluidized bed feeder
fluidization gas £lew into the vibraîhg pipe (3) of 0.10 m ID and 1.35 m long placed at an
angle of approximately 20' fiom horizontal. Remaining disturbances of the solids flow
were fiirther smoothed out by the vibration to achieve pseudo-stability of the solids mass
flowrate.
Solids entrance \ / 1
Figure 2.4 Schematic of the fluidized bed feeder funne1
When solids drop into the funne1 (4), illustrateci in Figure 2.4, at the axis of the
funne1 and the test column below, they were pre-accelerated by gravity. This was to
ensure efficient feeding of solids into the test column. With pre-acceleration, the particles
had attained the particle terminal velocity (about 0.18 mls as calculated by the method
proposed by Clift et al. 1978) at the entrance of the column if they feu directly into the
test column. Because the funne1 had a very steep angle, those particles which hit the
wall, still had a high downward velocity. Given the high initial solids velocity, high
solids flux could be achieved in the test column. Meanwhile, it also hetped to prwent
clogging at the throat under high solids flows. This Wei is very critical to achieving
high soli& flux operation
2 3 Description of the Particdate Mate-
The particdate solids used for this experiment were FLuidized Cracking Catalyst
PCC) particles, supplieci by Imperid OiI ùi Samia Ontario. A BrinIanan Particle Size
Analyzer was used to determine the particle size distnbutoa FCC particles were found to
have a wide size distribution, shown in Figure 2.5, with an average diameter of 65 pm
and density of 1550 kg/m3. These place the particles as A powder in Gelclart's (1973)
classification.
2.4 Measarement of Solids FIux
The soli& flowrate is determined by measuring the weight changes of the bottom storage
tank over a measured interval of time using the signais the load ceii installed
undemeath the bottom storage tank. This commercialiy available load cell (with a
standard capacity 500 kg) was caiiirated for use in this experiment. In order to Iimit the
effects of other parts on the weight changes of the bottom storage tank, aiI connections
between the storage tank and other parts of the apparatus were flexible. From the
calibration result as shown in Figure 2.6, the output signals (in millivoltage) have a Linear
relationship with the weight of FCC particles stored in the tank. This verined that those
flexible comections did not affect the signai output in the range of weight change.
Figure 2.5 Cumuiative size distribution for FCC particles measured using Brinkman Particle Size Analyzer (d,,=65 pm)
Figure 2.6 Calibration cuwe for the load ceii
To ensure the accuracy when measuring the weight change, especially in Iow
solids flux conditions, the output signal was amplified 500 times and the reference point,
at which the output was zm, was easily adjusteci by using a seIf-desiguecl electrical
differentiator circuit wîth the operatiod amplifier.
2.5 Measurement of the Axial Pressure Gradient
Five differentid pressure transducers fkom &ega@ are comected to pressure
ports at different axial locations along the column in order to measure the pressure
gradient (dl?/@ along the entire test column. The measuring ranges of the top hvo
transducers are -127 - +127 mm H20. Other three transducers with large measuring
ranges (O - 703 mm H,O) are installeci in the lower section of the test column. The top
two tramducers are of positivehegative type given the possibility of a negative pressure
change due to particle acceleration. The transducers had been calibrated by the
manufacturer, but they were each v d e d using a simple differential pressure manometer
to ensure the accuracy. The h e a r equations of the calibration are Iisted in Appendix -1
for aIi transducers. The DAQ software nom National Instruments Company was used to
sample at 100 Hz over a certain sampling period. The resdts are provided as either
instantaneous or time-averaged pressure gradients for each section, dehed by the port
locations Iisted in Table 2.1. For this study, the time-averaged pressure &op are used to
calculate the pressure gradient, which can then be plotted against the median location of
each section to show the pressure gradient profile dong the test column.
filied with FCC particles to find the equivalent height of the trapped solids by the
following relation, excluding the effect of the inegular shape changes of the inner rubber
tubes of the pinch volva whai close&
where cgp is the voidage of packed FCC particIes which is 0.45.
Mer each nm, the height of the packed solids trapped inside the tube, h. was
measured The actual solids holdup was then given by the following equation:
2.7 Operating Conditions and Experiment Procedures
The operating conditions for various experiments and the measured and calculated
parameters are listed in Tables 2.2,2.3 and 2.4.
For each test, pressure drop data were recorded by the cornputer and the soiids
fluxes were measured by monitoring the weight change over a specific time interval.
With changing superficial gas velocity in the operathg ranges, all tests for a same solids
flux were completed together as a group. In order to obtah mean solids holdup, the two
pinch valves were closed at the end of each experiment with the gas and solids flowrates
shut down at the same tune. Then, the mean solids holdups are measured by ident-g
the height of the solids trapped inside the tube between the two valves. After that, the
above steps were repeated with another sol& flowntte.
Table 2.2 The operating conditions for various experiments
I Gas-Solids Downflow Gas UpwardSolids Downward
Superficial gas velocity u, (mm
Table 2.3 The measured parameters
1 Mearured Parameters Measuring Toob 1 Gas flowrates Rotameters
Salids flowrate Load cell
Actual mean solids holdup Pinch valves
Pressures Pressure gauges
Differential pressures Pressure transducers
Table 2.4 The caiculated parameters
--- - -
Amal gas velocity Va = Ua / (I-es)
Mean partide velocity VD = 6 1 (PSES) Mean slip velocity U&) = Va = V, Apparent density PS = PsG + PO (1-ES) Pressure gradient dP/dH
The various valves (see Figure 2.1) were employed to switch between test nuis
and particle recycle operations, and between gas-soli& co-cumnt d o d o w and gas
upward-solids downward counter-cu~zent fiow operations, as shown in Table 2.5.
Table 2.5 Valve settings for the different operating modes
Operation Mode ---
Co-Current
Counter-Cunent
Particle Recycle
Valve Numbers
Closed Closed Open Open Open Closed Closed Closed
Ciosed Open Closed Open Closed Open Closed Closed
Open Closed Closed Closed Closed Closed Open Open
For CO-current downffow test, Valves 3, 4, and 5 were kept open and ail other
valves are closed.
For counter-current flow test, Valves 2, 4, and 6 were kept open and ali other
valves are close&
For particle recycle, Valves 1, 7, and 8 were kept open and all other valves are
closed,
2.8 Electrostatic Charging and Its Elimination
Electrostatic charging occurs by tribelectrification. In general, this means that the
contact and then quick sepration of two different d a c e s with différbg work hctions
causes one surface to be Ieft with a negative charge and the other with a positive charge
(Vonnegut, 1973). The charging of particles in a fluidized system is related to the fiow
conditions and the system parameters. Furthexmore, the effect is magnined by repetitive
collisions with the column walis (Nieh and Nguyen 1987, 1988). The pressure drop
within the testing c o I m may increase due to electrostatics (Smeltzer et al. 1982, Ally
and KIinzing 1983, Chang and Louge 1992).
To eIiminate the effects of electrostatics, 0.5% wt Larostat 519, an ammonium
compound was added to FCC particIes- This has been successfblly applied in other
studies to control electrostatics ( H d e r t et al, t 994).
2.9 Bibiiography
w, M. R. And Rlinzing, G. E. (1983), '%lecectrostatic Effects in Gas-Solid Pneumatic
Transport with Loacüngs to 1009', J m a l of Powder and Bulk Soliak Technology,
7(3), 13-20. -
Chang, H. and Louge, M. (1992). Thid Dynarnic Siniilarity of Circulatlng Fluidizd
Beds", Powdw Technol,, 7Q, 259-270.
Ciiff, K. Grace, I. R and Weber, M. E. (1978), BubbZes, Drups md PartiCles,
Academic Press, New York.
Herbert, P. M. (1994), 66Applicahion of Fik Optic Reflection Probes to the
Measmement of Local Particle Velocity and Concentration in Gas-Solid Flow",
ME.Sc. Dksertation, The University of Western Ontario, London, Canada
Nieh, S., and Nguyen, T. (1987). 'Mea~urement and Control of Electrostatic Charges on
Puiverized Cod in a Pneumatic Pipeliney', Paficulate Science and Techology, 5,
1 15-130,
Nieh, S., and Nguyen, T. (1988), "Effects of H~~nidity, Conveying Velocity and Particle
Size on Electrostatic Charges of Glass Beads in a Gaseous Suspension Flow'',
Journal of Electrostatics, U(I), 99-1 14.
Smeltzer, E. E., Weaver, M. L. And Klinzing, G. E. (1982), ''Pressure Drop Losses Due
to Electrostatic Generation in Pneumatic Transport", Ind. Eng. Chem. Process Des.
Dw., 21,390-394.
C W T E R 3 : CHXRACTERIZATION OF HIGH DENSITY GAS-
SOLIDS DOWNI?LOW FLUIDIZED REACTOR
A version of this chapter is to be submitted for publication to the journal Powder
Technology.
Characteriza45on of Eigh Density Gas-Solids DownRow ETuiàized Reactors
W. Liu, LX Zhu* and J , M, Beeckmcuts
Department of C h m i c d and Biochemical Engineering, University of Western &tano London, Ontario, Canada, N6A SB9
Abstract - Experiments were c e e d out in a specially designed 5 m tall, 0.025 m ID
high density gas-soli& downflow fluidized bed to mesure the axial pressure gradient
profiles dong the downer and the actuai solicis holdup in the fWy developed region. FCC
particles with a mean particle diameter of 65 pm and a d d f y of 1550 kglm3, a Gelciart
(1973) A powder, was used. A particle acceleration region and a fully developed region
were identified dong the columu fiom the pressure gradient profiles. In the fûlly
developed region, the apparent solids holdup calcdated fiom the pressure gradient agreed
well with the actual solids holdup measured by a pair of pinch valves under velocities Iess
than 5.6 m/s, but underestimated it at higher gas velocities due to the hcreased wali
Ection loss. Two different flow regimes were observed in the developed region, a
constant and high density pseudo-aggregative flow regmie under low gas velocities and a
reducing density pseudo-particdate flow regime under high gas velocities, with a
boundary between CI, = 0.5-1.3 d s . High density downfiow operation is dehned as
operation in the pseudo-aggregative fiow regime, where particle velocity remains
constant under aii solids flux and gas velocity conditions and where the slip velocity is
very high, with sîgnincant particle agglomeration. A solids holdup as high as 10% has
been achieved in this operating regime. In the more dilute pseudo-particdate flow
regime, the gas-particle slip velocity remains constant and no particle strands and large
particle clusters were observed. The particle velocity was found to increase hearly with
the gas velocity given the constant slip velocity. Consequently, the solids holdup
decreased with increasing gas velocity in this regime, as reported previously in other rïser
and downer systems. Cornparison of the r d t s obtained here with those nom an upflow
riser shows inherent similarities between the two gas-solids CO-current flow systems.
3.1 Introduction
In the past two decades, there has been considerable industrial and academic
interest in circuiating fluidized beds (CFBs) which have been widely applied for the
Fischer-Tropsch pmcess (Shingies and McDonald 1988. Steynberg 1991), the FCC
(FluidÏzed Cataiytic Cracking) process (Avidan etal. 1990, Kuig 1992) and coal
combustion (Dry and La Nauze 1990, Engstrom and Lee 1991, Kdendodf and
Andersson 1986). Many more new applications are also being considered (Contractor
1988, 1991, Zhu and Bi 1995). In addition to the cocurrent gas-soi& upflow circuiating
fluidized beds (risers)? cocurrent gas-solids downflow circulating fluidized beds
(downers) were proposed in recent yem (Gross 1983, Gross and Ramage 1983, Berg et
al. 1989, Gartside 1989, Bai et al. 1991, Wang et al. 1992, Zhu et al. 1995, Zhu and Wei
1996). With many advantages, such as good gas-solids contact, less gas and solids back-
mixing, a short contact tirne and d o m residence time distribution compared with the
upflow fast fluidized bed (riser), downer reactors become more advantageous over risers
for reactions of very short residence time and ceactions where the intmediates are the
desirable products.
Notwithstanding the nurnerous advantages of the downer reactor, it suffers a
serious shortcoming: very low volumetric concentration (holdup) of solids in the bed. A
typical riser contains 1 3 % solids in the fUy developed region. On the other hand, the
solids holdups achieved in downers as shown by the reporteci studies are much more
dilute (mostly below 1%). This represents a serious problem for reactions where a high
soliddgas ratio is requircd since the reaction intensity is limited by the lower solids
concentration- To ovemme this wealaiess, an attempt was made in this work to achieve
high densities in a C O - c u m t downfhw system.
Because both gas and solids flow in the direction of gravity and the soli& are
accelerated even without the aid of the gas, very high solids fluxes must be achieved at a
relatively low superficial gas velocity to have a high density in the fiilly developed region
of the bed. In the co-current downflow circuiating fluidized bed (downer) studies reported
so far (Zhu et al. 1995, Herbert 1997, Herbert et al- 1998, Zhang et al. L999a,b ), it was
impossible to reach high densities due to feeder restrictions and the pressure balance in
the system. To facilitate high density operation, a specialiy designed feeder system was
employed to achieve high solids flux.
It was important to study gas-soüds downflow at high density under different
operathg conditions for the potential applications and for extendhg our current
knowledge. The objectives of this work were: (i) to achieve high density downflow; (ii)
to characterize the gas and solids flow in cocunent high-density downflow; (iii) to
compare the characteristics of dowdlow with those of upflow; and (iv) to find the
relationship between the operating parameters and the solids density.
3.2 Experimental and Operating Procedures
A schematic of the experimentai apparatus is illustrateci in Figure 3.1. There were
two conduits, the test Iine and the particle recycle line, and two solids storage tanks, at the
top and at the bottom. The operations of the tests were on a batch basis. Solids flowed
down fcom the top tank to the bottom one during the tests in the test Line and were
entrained up in the recycle line d e r the test. The test colum. was made of plexiglass,
with an inner diameter of 0.025 m and a full length of 5.0 m. Eight pressure taps were
installeed along the column, located at 0-10 m, 030 m, 0.50 m, 1.00 m, 2.00 m, 2.50 m,
3.75 m and 4.25 m m m the top entrance of the culumn, giving five differential readings
(except between 1-00 m and 2.00 m, 2.50 m and 3.75 m). D u ~ g these experiments the
differential pressures along the test column were measured by pressure tramducers.
Pressure gradients were then caicdated b m the measured differential pressures. At
levels 3.00 m and 4.50 m fiom the top of the column, two pinch vaIves were installeci to
obtain the achial solids holdup in the M y developed region by collecting and weighing
the sol& trapped between the two valves whai they were closed simultaneously at the
end of each experiment. The solids flux was determinecl by a load cell installed
underneath the bottom storage tank which monitored the rate of weight change of the
tank
The soli& feeder system, consisting of a fluidized bed feeder, a vibrahg pipe and
a feeding -el, as shown in Figure 3.1, was specially designed to achieve very smooth
and high solids fluxes. The upper portion employs a fluidized bed feeder of 0.10 m ID
and 0.71 m height, and an inclinecl vibrating pipe of 0.10 m ID and 1.35 m in length to
regulate the solids flowrate. By changing the level of the small movable tray, the
fluidized feeder delivered a regulated amount of solids into the inclined pipe. Through
vibration, the 200 inclined pipe (at an angle srnalier than the angle of repose) fkther
darnped the fluctuations in solids flow. The lower portion containeci a 0.66 m high
feeding funnel, within which solids were pre-accelerated by gravity before entering the
test line. Since the particles had an initial velocity close to the terminal velocity
(calculations perfonned according to the method suggested by Clifi et al. (1978) upon
entering the downer top, choking was avoided and solids fluxes up to 500 kg/m2s wuld
be achieved by this novel feeder.
The main gas flow was introduced in the top of the -el, and solids and gas then
flowed downward t b u g h the test column. The flowrates of the main gas feed and the
fluidization gas to the feeder were both monitored by rotameters. The gas ffowrate in the
test column was caiculated as the fïowrate of the mai . gas feed plus the flowrate of the
feeder fluidization gas. Experiments were carried out over a wide range of superficial gas
velocities between 0.16 d s and 10.4 mls and soiids fluxes fiom 23 kg/m2s to 400
kglm2s. FCC particles with an average diameter of 65 pn and a density of 1550 kg/m3, a
Geldart (1973) A powder, were used for the tests- To minimize the undesirable effects of
electrostatics, 0.5% wt of commercial Larostat powder was added to the solids. Once
solids and gas fell into the bottom storage tank, the solids were separated nom the gas by
gravity and depositcd in the tank, and the gas flowed to a fdter bag, where the maining
fines were coiiected before the gas entered the exhaust system.
3.3 Results and Discussion
3.3.1 Pressure Gradient ProIües and the Solids Acceleration Length
Typical pressure gradient profiles in the axial direction are shown in Figure 3.2
for different operating conditions. The pressure gradient is initiaily low, either positive or
negative at the column top, but rapidly increases withui a distance of less than 1-2 m, and
then gradually approaches a constant value dong the column length. The very low
pressure gradient near the downer top is due to the rapid acceleration of solids which
leads to a large pressure loss. An increase of solids flux increases the value of the
pressure gradient at a given gas velocity Figure 3.2a) and an increasing gas velocity
decreases the pressure gradient for a constant solids flux (Figure 3.2b). These trends are
generdy consistent with what had been observed in previous studies in dilute downfiow
fluidized beds (Wang et al. 1992, Herbert et al. 1998, Johnston et al. 1999, Zhang et al.
1999b).
The pnssure gain in the downwd direction for the gas-solids downflow in the
column may be estimateci using the foiiowing relation
where Af, is the pressure drop due to solids acceleration and M'is the pressure &op
due to waU fiction. When the flow of gas and soli& is fixlly developed and has reached a
steady state, thete is no acceleration of particles, no change of solids holdup, and no
change in the wali fiction, so the pressure gradient in this region must be constant-
Therefore, the pressure gradient profiles shown in Figure 3.2 c m also be used to identifjr
the two flow regions dong the downei: the initial downer section with varying pressure
gradient is the solids acceleration region and the remaining section with constant pressure
gradient is the M y developed region. That is, (~ZP/&) = O in the M y developed
region. In practice, it was assumed in this study that the fully developed region has been
reached when
Compared to the maximum values of 1000 pa/rn2 for the slopes of the curves at the top of
the column, this srnail tolerance is considered reasonable.
With the above method, one can see h m Figure 3.2 that increasing gas velocity
ancilor solids flux both lengthen the solids acceleration region. Figure 3.2a shows that an
increase of solids flux slightly increases the solids acceleration length. This is
understandable since more solids are fed into the system and may take longer to reach the
fùlly developed state. An increasing gas velocity is shown in Figure 3.2b to lengthen the
solids acceleration region more significantly, since the "cquilibriumyy particle velocity in
the hilly developed region increases with the gas velocity.
In Figure 32% a positive pressure gradient at the top of the column is m e a d
for most of the t h e under a low gas velocity. This suggests that a pressure gain due to
solids holdup under low gas ve1ocities is larger than the pressure loss caused by the
particle acceleration and the fiction between the waii and the gas-solids suspension.
When the solids flux is higher, the pressure gain due to solids holdup is larger so that the
absolute value of the pressure gradient is also higher. The pressure gradient in the
entrance region onîy becornes negative at low soIids flux, as s h o w in Figure 3.2%
because the pressure loss dong the column due ta particle acceleration and fiction
between the gas-soiids suspension and the wall, items AP, and dPf in eqn (1). exceeds
the pressure gain due to gas-soli& weight in the given section.
At higher gas velocities, the pressure gradient is mostly negative in the solids
acceleration region (Figure 3.2b). At very high gas velocity (Ug=10.4 ds ) , the pressure
gradient is negative even in the fblly developed region. This is due to the significant wali
friction in the small diameter test column, which results in a larger pressure loss than the
pressure gain fiom the gas-soi& holdup. That is, dP'> gp,~#T in eqn (1). in the M y
developed region.
However, the pressure gradient is not always negative at the downer top for all the
operating conditions, as shown in both Figures 32a and 3.2b. This is different h m that
observed in the dilute dowdow fluidized beds as reported by Wang et al. (1992), Aubert
et al. (19941, Zhu et al. (1 993, Herbert et al. (1998) and Johnston et al. (1999). Wang et
al. (1992) suggested that the pressure gradient profiles be used to iden- the h t particle
acceleration region (where the particles are accelerated by both gravity and gas drag with
VBY,), the second particle acceleration region (where the particles are m e r accelerated
by gravity but resisted by gas drag with VgcVP) and the M y developed region.
Neglecting wall fiction, the zero pressure gradient point signifies the boundary between
the first and the second acceleration regions. For this study, because of the signincant
particle pre-acceleration in the spaciaiiy designed feeder, the particles sometimes e n t d
the column at velocities higher than the gas velocity, su that the fht acceleration region
was eliminated- Furthenmore, the wall fiction also becomes more significant at high
suspension density. Therefore, no clear demarcation can be found between first and
second acceleration regions h m the pressure gradient pronle alone.
3.3.2 Cornparison between the Actuaï and the Apparent Solids Holdups
In the M y developed region, the pressure gradient can be used to calculate the
soli& holdup if fiction is neglected:
The solids holdup, E, thus obtained is called the apparent solids holdup, which has been
used by previous researchers to cfiaracterize the flow in the nser (e.g. Bai et al. 1992). On
the other hand, Zhu et al. (1995) have shown that the pressure gradient cannot be used to
estimate the actual solids holdup in a dilute downer given the lower solids holdup and the
relatively hîgh suspension-to-wall fiction. It would therefore be interesting to examine
the pressure gradient behaviour in gas-soiids downfiow in the current experiments. Figure
3.3 compares the actuai and apparent solids holdups obtained in this study. At low gas
velocities, therc is no signifïcant difference between the actual solids holdup measured by
the pinch valves and the apparent solids holdup calculated from the pressure gradient.
When the gas velocity is increased, the fiction between the wall and the gas-soli&
suspension become larger, leadhg to a significant ciifference between the actual solids
holdup and the apparent solids holdup. In such cases the pressure gradient cannot be used
to accurately calculate the solids holdup. Within acceptable maximum errors of f 15%
(dashed lines in Figure 3.3), a gas velocity higher than 5.6 mls seems to cause the
apparent solids holdup to deviate signincantly h m the actual solids holdup, makhg the
calculated apparent solids holdups lower than the actual soüds holdups and even resuIting
in negative values for the apparent solids holdup at a very high gas velocity of 10.4 mls.
Nonetheless, Figure 3.3 does suggest that the pressure gradient can still provide a simple
and reliable method to estimate the solids holdup at different SOI& fluxes in the M y
developed region of the high density downer for gas velocities lower Hhan 5.6 d s .
3.33 Solids Holdup, Particle Velocity and Slip Velocity in the FuUy Developd
Region
Solids holdup is one of the key parameters which characterize a gas-solids system.
In a CO-current gas-solids system, either upflow (Bai et al. 1992) or downflow (Zhu et al.
1995), gas velocity and soüds flux are the main operating variables influencing the solids
holdup. Generally, an increase of gas velocity decreases the solids holdup at a constant
solids flux and an increased solids flux results in an increase in the solids holdup when
the gas velocity is fixed. This, however, is wt always tme for the high deasity gas-solids
CO-current downflow system reported in this study.
Figure 3.4 shows the mean solids holdup in the M y developed region as a
fünction of solids flux at different gas velocities. For the entire range of gas velocities
tested, solids holdup in the M y developed region is seen to increase with the solids flux,
just as in a gas-solids riser. Furthexmore, a lhear relationship clearly exists between the
solids holdup and the solids flux, something not obswed in a nser. This is Wrely due to
the fact that in downfiow situation particles need not to be cmied by the gas. By mass
balance, one has
A linear relatiomhip between 4 and G, indicates that Y, must be a constant, that is Y, is
independent of solids flux. In a fdly developed dow&w system, since particles need
not to be carried by the gas, an inmase in solids flux would not increase the %mien" of
the gas and therefore wiii not decrease the particle velocity (as in a riser). In this case, the
particle velocity is only a function of the gas velocity, as can be seen by the change in the
slopes of the hes [ I f @ , G)] in Figure 3.4 wah the gas velocity. Since the gas flow,
which travels slower than the solids in the M y developed region in the downer, ex- an
upwards drag on the solids flow, an increased gas velocity would increase the "final" or
cceqUiiibnum" particle velocity achieved in the downer.
It is interesthg to note in Figure 3.4 that soiids holdup becomes independent of
the gas velocity at 1ower gas velocities. This is more clearly shown in Figure 3.5 where
the solids holdup in the M y developed region is plotted against the superficial gas
velocity at diffkrent solids fluxes. The plot can be divided into two dinerent segments: an
initial constant and high density segment under low gas velocities and a reducing density
segment thereafter. In the high velocity segment, the solids holdup is seen to decrease
rapidly at first and then more graduaily with an hcrease of gas velocity. This is the same
as observed in nsers and is also as expected, since increased gas velocity leads to an
increased particle velocity, which in tum reduces the solids holdup. In the initial low
velocity segment, however, the solids holdup is seen to be at its maximum value for a
specific solids flux and remains almost constant in a narrow range of low gas velocities.
Given the constant E- 5 must also be constant as weU in this initial segment for a given
solids flux. Such phenornena have not been observed by earlier researchers (e.g., Herbert
et al. 1998, Zhang et al. 1999b).
To understand the existence of this special low gas velocity operating range with
constant and high suspension density, it is userul to examine the particle velocity and the
slip velocity unda différent flow conditions. The variations of the particle velocity and
the slip velocity with the gas velocity at ciiffirent soli& fluxes are shown in Figures 3.6
and 3.7. Figure 3.6 shows that the p d c l e velocity is essentially constant in the initial
region, except for G, = 46 kg/m2s where the data seem to be somewhat out of the space as
shown in Figure 3.4. This re-confimis the earlier argument of constant Y , in the low
veiocity segment. In the high velocity segment, the particle velocity increases linearly
with the gas velocity. Ln other word, the mean gas-particle slip velocity defïned as:
is constant in this velocity range. Figure 3.7 shows a ciramatic decrease of the slip
velocity with gas velocity in the low velocity segment and then a fairly constant slip
velocity in the high velocity segment.
Apparently, there edsts an operating range under low gas velocities where the
particle velocity remains more or less constant, being neither a fimction of the gas
velocity nor a function of the solids flux; where the solids holdup is only a fiuiction of
solids flux, and where the slip velocity is very high and decreases quickly with increasing
gas velocity. Such an operating condition has never been rqorted before in either risers
or dilute domers.
The very hi& slip velocity (1.7 to 0.6 m/s, much higher than the single particle
temiinal velocity of 0.18 m/s) in this hi&-density region indicates that particle
agglomeration is fjlirly severe. (Particle agglomeration produces large effective particle
sizes in the gas stream, which leads to higher e f f d v e terminal velocities and therefore
high observed gas-particle slip velocities.) This is indeed what has been observed during
the experiments: In this region, the particles were seen to flow downward in apparently
continuous strands and clustem both in the core region and dong the wall. On the other
han& the flow structure was observed to be more homogeneous with more uniform gas-
solids suspension and without apparent strands and large clusters beyond the velocity
which demarcates the low velocity high density region h m the reducing density region.
In the latter region, the slip velocity was found to ranain fairly constant around 0.57 mis,
about 3 times of the singIe particle termuial velocity of 0.18 mls as calculated using the
method proposed by Clifk et al. (1978). This result is consistent with the resuit of Yang
et al. (1993) who measured the local gas and particle velocities simultaneously in a dilute
nser and found that the actual gas-solids slip velocity, excluding any effect of radial
segregation, in the M y developed region in a FCC riser was more or less constant and
was about 3 times the particle temrinal velocity. With Ut = 0.18 mfs for the FCC particles
used here, the value of the constant slip velocity (-0.57 m/s) in this downer is also
approximately three times CI,. It would appear that the particles are fdIy suspendeci in this
high velocity segment so that the value of the slip velocity becomes similar to that in a
fÙUy suspended dilute riser. Therefore, this region may be called the pseudo-homogenous
regime of operation. The constant slip velocity in this regime indicates that the particle
agglomeration does not change signifïcantly with increasing gas velocity. In other word,
the gas-solids flow has achieved a fûUy suspended equilibnum state-
To the contrary, the slip velocity in the low velocity and high-deasity segment
was very high, with clearly observed particle strands and large clusters both in the core
region and dong the wall. Therefore, this may be refmed to as the pseudo-aggregative
operating regime. In this hi&-density regime, a very hi& loading ratio can be achieved
because of the low gas velocities. In addition, an increase in gas velocity results in a
decrease in the slip velocity and therefore reduced particle agglomeration. On the other
hand, higher solids fluxes lead to higher particle slip velocities because the flow becomes
denser and produces more particle agglomeration.
In this regime, it would appear that the particles tend to accelerate themselves by
gravity but the low gas velocity cannot "catch up". Since the very slow flowing gas
would exert resistance to the particle flow, particles need to form larger clusters to create
high slip in order to accelerate. Figures 3.6 and 3.7 shows that the bounda~~ between the
two regimes lay within 0.5-1.3 mls and the constant particle velocity in the pseudo-
aggregative regime ranained between 1.2-1.9 d s for Mirent soi& fluxes. The
maximum value of this transition gis velocity, between 0.5-1.3 d s , coincides with the
value of the omet velocity for signiscant solids entrainment in the FCC riser
(U, = 1.29 d s ) , as calculated using the foliowing equation recommended by Bi et al.
(1 995)
This omet velocity for signincant solids entrainment is denned by Bi et al. (1995) as the
velocity below which the solids flow in a riser is no longer M y suspendeci by the gas
flow and the flow regime enters the turbulent fluidization region with high bed density. In
the dowdiow system, the same velocity seems also to demamate the transition between
the high density pseudo-aggregative flow regime and the pseudo-particdate flow regime.
It is worth noting fiom Figure 3.6 that solids flux has little effect on the particle
velocity in the pseudo-particulate regime. In other word, according to eqn (4), an increase
of soiids flux at a constant gas velocity only results in a linear increase of solids holdup
but does not change the ratio of the solids flux to the soîids holdup in the downer, being
dilute or dense. Furthemore, sol& flux also has Little effect on the relationship between
the gas and the particle velocity in the downer, in either the pseudo-particdate or the
pseudo-aggregative regimes, although this relationship is distinctly different in the two
regimes. This would be a special feature of downeis given the fact that particles can
accelerate without the assistance of the gas, so that an increase of particle flux ody
changes the solids holdup but not the particle velocity and its relationship with the gas
velocity .
With the independent relationship between the gas and particle velocity, the
particle velocity and therefore the soli& holdup can be predicted for given U, and Gr For
FCC particles, Y, = 1.249 mls in the pseudo-aggregative @me and
Y, = U, + 0.55 mls in the pseudo-particdate regime. Then, the solids holdup can be
cdculated using eqn (4) with known Gr This way, a simple procedure is available for the
estimation of solids holdup, an extremely important parameter in gas-solids dowdiow in
a downer.
It should be pointed out that the results presented in Figure 3.7 are different h m
those reported by Herbert et al. (1998), who showed increased Us,* with increasing U,
under constant Gr Since it is understood (Yang et al. 1993) that iacreasing U, tends to
decrease the extent of particle agglomeration, it would be more reasonable for Us[$ to
decrease with Ur
3.3.4 Definition of High Density Downfiow Operation
From the above discussion, high density downer operation may be dehed as the
operation of a gas-solids CO-curent downfiow system in the constant and hi&-density
pseudo-aggregative flow reghe, where the particle velocity is constant, and the solids
holdup is a fùnction only of solids flux, and not of the gas velocity.
The operating window for gas-solids downfiow achieved in this study, for both
the high density and the reguiar downer operations, is shown in Figure 3.8. This figure
clearly shows the trend of increasing solids holdup with increasing solids flux and
decreasing gas velocity. The high density regïme is achieved with a combination of
higher solids flux and Iower gas velocities, as in the case for a high density riser (Zhu and
Bi 1995). %th the cumnt experimentaL se-, a solids holdup bigher than 10°h can be
achieved under the high density conditions.
3.3.5 Cornparison between Downflow and Upflon Systems
As discussed above, most of the general trends for the effects of gas velocity and
solids flux on the flow behaviow obsmed in the fully developed region of the riser aiso
hold tnie in the fbUy developed region in the gas-solids co-current downer. In Figure 3.9,
solids holdups are plotted against superficial gas velocity in both nsers and downers, with
the riser data fiom Huang and Zhu (1998). It seen that these c w e s have similar trends,
with solids holdup decreasing with the gas velocity, but the two groups of curves do not
match weU. The solids holdups in the riser are consistently h i g k than those in the
downer at the same gas velocity and solids flux. This is not surprishg since it is the
particle velocity, not the gas velocity, which determines the solids holdup. The fact is that
particles travel slower than the gas in the nser but faster in the downer. To make a more
direct cornparison at similar particle velocities, the solids holdups are plotted in Figure
3.10 as a function of superficial velocity plus the average slip velocity of 0.57 mls for the
downer, and as a ninction of the superficial velocity minus the average slip velocity of
0.57 m/s for the riser. The figure shows that, after shifting the abscissa (U,) by 2 times of
the mean slip velocity, the curves for both the downer and the riser now coincide very
weU with each other in the higher velocity range where the condition Us[, =O57 m/s
holds. Cledy, some inherent similarity begins to show up between the CO-current upflow
gas-solids riser and the co-cment downfiow gas-solids downer. More investigation is
needed before more definite conclusions can be drawn.
3.4 Conclusions
The principal conclusions of this study are as foJiows:
(1) Along the downer colrrmn, a particle acceleration region and a fblly developed region
were identifiai under all operating conditions testeci. The two regions Ca. be
established by examining the measured pressure gradient profiles, with a constant
pressure &radient signifiing the M y developed region.
(2) The experimental r d t s indicate that the pressure gradient provides a simple and
reliable method to estimate the solids holdup without incufTing large errors when the
gas velocity is lower than approximately 5.6 m/s at différent soüds fluxes in the M y
developed region. Beyond this gas velocity, wali Ection leads to a significant
underestimation of the actual solids holdup when using the pressure gradient
method,
(3) Two different flow regimes have been identifieci in the developed region, a constant
and high density pseudo-aggregative flow regime at low gas velocities and a low
density pseudo-particdate flow regime at high gas velocities, with a boundary within
the range U, = 0.6- 1.3 m/s for dinerent solid fluxes.
(4) Hi&-density downfiow operation is defined as operation in the pseudo-aggregative
flow regime. In the high-density flow regime, particle velocity remains constant for a
given solids flux and is independent of the gas velocity, and the slip velocity is very
high with very significant particle agglomeration.
(5) In the more dilute pseudo-particdate flow regime, the gas-particIe slip velocity
remains constant and no particle strands and large particle clusters are observecl. The
constant slip velocity suggests that equilibrium has been reached in particle
aggregation so that it no longer changes with the gas velocity. The particle velocity
increases linearly with gas velocity given the constant slip velocity- Consequentiy,
the soüds holdup decreases with hcreasing gas velocity in this regime, as reported
previously in otha n s a and downer systems.
(6) The operating window for gas-soli& do&w in the experimental
apparatus has been mapped- For the experimental conditions employed in this study,
solids densities above 10% cm be reached in the high density flow regime.
(7) Cornparison of the results obtained in this study with those h m an upfIow riser
shows some inherent similarities between dowdiow and upfbw gas-solids CO-
current flow systems. The variations o f solids holdup with gas velocity becornes
consistent for both systems when the slip velocity is deducted h m the nser gas
velocity and is added onto the downer gas velocity.
Financial assistance nom the Natural Sciences and Engineering Raearch Council
of Canada is gratefidly aclcnowledged. The authors are also gratefbl to H. Zhang and PM.
Johnston for providing valuable discussions in the design of the experimental equipment.
Notations
g acceleration due to gravity? m/s2
G, soli& flux kg/m2s
H distance h m the top of test column, m
P pressure? Pa
U, superficial gas velocity, mls
u,' defined as U, + 0.57 mls m the downer and U' - 0.57 mis in the riser, mls
Use the omet velocity for Signifïcant solids entrainment
U s mean particle slip velocity, mls
V, mean actuai gas velocity, mls
V, mean particle velocity, mls
S tolerance for pressure gradient variation, pa/rn2
AH distance between the two pressure taps, m
dP,, pressure loss due to particle acceleration, Pa
LW' pressure loss due to suspension-to-waIl friction, Pa
E, soli& holdup
pg densiîy of gas phase, kglm3
p, density of solids phase, kgfm3
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Cornmerciaikation of Rapid Biomass Pyrolysis for Fuel and Chcmical Productîoa",
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Herbert, PM. (1997), 'Tlydrodynamic Study of a Downflow Circdating Fluidized Bed",
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study of a O.OSm diameter downflow circulating fluidized bed", Powder Technol.,
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Huang, W-X. and Zhu, J X , '%xperimental Study on Solids Acceleration Length in a
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Johnston, P. M., de Lasa, H. 1. and Zhu, 1. (1999), "Axial Flow Structure in the Entrance
Region of a Downer Fluidized B e b Effects of the Distributor Design", Chern. Eng.
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Kullendorff, A. and Andersson, S. (1986), "A General Review of Combustion in
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pp. 83-96, Pergamon Press, Toronto.
Shingles, T. and McDonald, A F. (1988), "Commercial Experience with Synthol CFB
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pp. 43-50, Pergamon Press, Toronto.
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Wang, Z., Bai, D. and Jin, Y. (1992), b'EZydrodynami~~ of Cocunent Downflow
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the Circulating Fluidized Bed", AIC7i.E Symp- Ser., @(296), 81-90.
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Laguerïe), pp. 50 1-51% Engineering Foundation, New York.
Figure Captions
Figure 3.1
Figure 3.2a
Figure 3.2b
Figure 3 -3
Figure 3 -4
Figure 3.5
Figure 3 -6
Figure 3.7
Figure 3 -8
Figure 3.9
Figure 3.10
Schema for the experimental apparatus
Pressure gradient profile dong the column at superficial gas velocity
U, = 0.33 d s
Pressure gradient profile dong the column at soli& flux Gs = 90 kg/m2s
Cornparison of actual and apparent solids holdup
Measured solids holdup in the fully developed region as a h c t i o n of U,
Solids holdup in the hlly developed region as a fimction of solids fia
Particle slip velocity as a Mction of superficial gas velocity
Mean particle velocity as a b c t i o n of supeficial gas velocity
Operating window for gas-solids co-current downfiow systcms
Solids holdup as a hction of Cl, in the M y developed region for risers and
downers
Solids holdups versus adjusted superficial gas velocity in the M y
developed region for both risers and downers
I - UPPER STORAGE TANK
PIPE ID 2"
PARTICLE RECYCLE UNE
TEST COLUMN
Pinch valves - : : -
VALVE4 - - -
BOTTOM STORAGE TANK
L
VALVE 6 VALVE 7 ROTAMETER
VALVE 8
AIR
- - - VALVE 2
Sdids
Schematic diagrarn of the fiuidued bed feeder
* - - - VALVE 5
Figure 3.1. The schematic of the experimental apparatus
Distance from the top entance, H (m)
Figure 3.2a Pressure gradient profile dong the column at difFerent U, and G,
Distance from the top entrance, H (m)
Figure 3.2b Pressure gradient profile h g the column at soiids flux G,=90 kglm2s
0.00 0.02 0.04 0.06 0.08 0.10 0.12
Actual solids holdup, E,
Figure 3 3 Cornparison of actual and apparent solids holdup
m 266 250 300
Solids f l u ~ G, (kglrn2s)
Figure 3.4 SoIids holdup in the M y developed region as a fùnction of solids flux
Supeiriaai gas velocih/, U, (m/s)
Figure 3.5 Measined solids holdup in the M y developed region as a fùuction of CI,
O 2 4 6 8 10 f 2
Superficial gas velocity, Ug (rnls)
Figure 3.6 Mean particle velocity as a hction of superficial gas velocity
- - - - - - -
O 2 4 6 8 10 12
Superficial gas velocity, U, (mls)
Figure 3.7 Particle slip velocity as a fùnction of superficial gas velocity
w I w I m I w 1 m 1 8
es=ll.6 %- I
~ ~ ~ 7 . 5 % _
a -
m
œ
I
O 1 I L I 1 1 I I 1 I 1 L
O 2 4 6 8 I O 12
Superficial gas wlocity , U, (rds)
Figure 3.8 Operating window for gas-solids CO-cment downflow system
Superlicial gas wlocity, Ug (mls)
Figure 3.9 Solids holdup as a function of U, in the W y developed region for both the riser and the downer
nnn I 1
Figure 3.10 Solids holdups vs. the adjusted superficial gas velocity in the fûUy developed region for both the nser and the downer
CEAPTER4: CEARAC'IEREATION OF GAS UPWARD-
SOLIDS DOWNWARD COUNTER-CURRENT
FLUIDIZED FLOW
A version of this chapter is to be submitted for publication to the journal Powder
TechnoI~gy~
Characterization of Gas Upward-Solids Downward Cornter-Current Fluidized Flow
K Liu, J-X Zhu' and J M- Beecham
Department of Chmicul and Biochenricu~ Enginee- University of Western Ontario, London, Canada, N6A 5B9
Abstract - Experiments were camed out in a specially designeci 0.025 m ID, 5 m high
gas upflow-solids downflow fluidized bed to measure axial pressure gradient profiles and
actual solids holdups. FCC particles with a mean particle diameter of 65 pm and a density
of 1550 kg/m3, a Geldart (1973) A powder, were used. The flow development and fiction
are discussed in relation to the pressure gradient profiles. The actual SOI& holdup
measured by a pair of pinch valves and the apparent solids holdup calculateci h m the
pressure gradients are compared for different operating conditions. B a d on the changes
of the rnean particle velocity and the particle slip velocity, the particle aggiomeration was
studied. Choking is discussed in relation to both nser and counter-current operation. The
operable maximum superficial gas velocity and solids flux in this system for FCC were
exp erimentally detemiined.
4.1 Introduction
Saidies on gas-solids fluidized beds have mainly bcen in gas-solids co-current
upflow systerns, which includa the conventional fluidized bed and the cimilating
fluidized beds (CFBs or risers). In ment y-, some b t e d studies on CO-current gas-
solids dowdhw circulating fluidized beds (dowriers) have been reported, as reviewed by
Zhu et al. (1995) and Zhu and Wei (1996). With many advantages, such as good gas-
soli& contact, less gas and solids back-mïxing, a short contact time and unSiorni
residence t h e distri'bution compared with the upfiow configuration (risers), downer
reactors become more advantageous than risers for veiy short residence time reactions
and for reactions where the intermediates are the desirable products. Because both gas
and solids travel in the direction of pvity, the flow accelerates much more quickly in the
downer since solids are accelerated by both the gas and gravity (compared to the riser
where solids are only accelerated by the gas flow, but resisted by gravity) and the
acceleration length in the downer is therefore shorter (Wang at el. 1992, Johnston et al.
1999). Notwithstanding the numerous advantages of the downer reactor, it SUffers a
senous shortcoming: very low volumetric concentration (holdup) of solids in the bed. A
typical riser contains 1-3% solids in the fully developed region. On the other hand, the
solids holdups achieved in domers as shown by the reported studies are much lower
(mostly below 1%). This represents a problem for reactions where a high soliddgas ratio
is required, since the reaction intensity is limiteci by the lower solids concentration. To
overcome this weakness, the hi&-density co-current dowdow reactor (Liu et al. 1999a)
was proposed.
Another alternative to overcome this limitation of low gas-soiids suspension
density is to ernploy a gas upflow and solids downflow counter-cunent fluidized bed.
However, no similar work on counter-current flow has been published before. Therefore,
it is important to study gas-solids flow under these conditions in order to extaid our
curent knowledge of nuicikation to wider operating ranges.
The specinc objectives of this worlc were: (i) to h e the pomile operation
range; (ii) to characterize the gas and soli& flow in counter-curent mode; Ci) to fïnd out
the relatiomhip between the operating parameters and the gas-solids suspension density;
and (iv) to discuss the choking phenornenon in a new context.
4.2 Experimental Apparatus and Procederes
Experiments were camed out over a range of upflow superficial gas velocities
between O and 0.39 m/s and dowdow solids fluxes h m 4.5 to 19.5 kglm2s. Running at
higher G, became problematic due to blocking caused by particle choking at the entrance
of the test column.
A schernatic of the experimental apparatus is given in Figure 4.1. There were two
flow conduits, the test iine and the particle recycle line, and two solids storage tanks at
the top and at the bottom. The tests were operated on a batch basis. Solids flowed down
fiom the top tank to the bottom tank during the tests in the test line and were entrained up
in the recycle line &er the test. The test column was made of plexiglass, with an inner
diameter of 0.025 m and a full length of 5.0 m. Seven pressure taps installed dong the
column were located at 0.30 m, 0.50 m, 1.00 m, 2.00 m, 2.50 m, 3.75 m and 4.25 m h m
the top entrance of the column, giving five differential readings (except between 2.50 m
and 3.75 m). During these experirnents the dinerential pressures dong the test column
were measured by pressure tramducers. Pressure gradients were then calculated fiom the
measured differential pressures. Two pinch valves were instaiied at distances of 3.00 rn
and 4.50 m fiom the top of the column to obtain the true solids holdup in the M y
developed region by collecting and weighing the soli& trapped between the two valves
when they were closed simultaneously at the end of each experiment. The soli& flux was
detennined by a load cell installed undemeath the bottom storage tank which monitored
the weight of the tank as a hc t ion of time.
The solids feeder system, consisting of a fluidized bed feeder, a vibrating pipe and
a feeding funnel, as s h o w in Figure 4.1, was specially designed to achieve very smooth
solids fluxes and to amid blochg at the entrance of the test coIumn. The upper portion
employed a fluidized bed fder of 0.10 m ID and 0.71 m height, and an inclined
vibrating pipe of 0.10 m ID and 1.35 m length to regdate the solids flowrate. By
changing the level of the smaU movable tray, the fluidized feeder delivered a regulated
amount of solids into the inclined pipe. Through vibration, the 20' incluied pipe (at an
angle smailer than the angle of repose) M e r damped the fluctuations in solids flow.
The lower portion of the feeder contained a 0.66 m high feeding funnel, within which
solids were pre-accelerated by gravity before entering the test line. Since the particles had
an initial velocity close to the terminal velocity (calculations performed accordhg to the
method suggested by Clift et al. 1978) upon entering the downer top, localized choking at
the entrance at low solids flux was avoided. The main air feed was introduced at the top
of the bottom storage tank, and gas then flowed upwardly through the test column. The
flowrates of the main air feed was monitored by a rotameter. The gas flowrate in the test
column was calculated as the fiowrate of the main air feed plus the displacement of air in
the bottom tank by the downflow particles. But due to the small operation range at low
solids fluxes, the displacernent air was very smaU compared with the main air feed. Even
at the maximum G, the displacement air was only about 6.4 x 1 0 ~ m3/s or 0.01 mls. So a
zero (0.0 m/s) displacement gas velocity was assumed when the main air feed was
positive. FCC particles with an average diameter of 65 prn and a density of 1550 kg/m3, a
Geldart (1973) type A powder, were used for the tests. To mitigate the undesirable effects
of electrostatics, 0.5%wt of commercial Larostat powder was added to the soli&, Solids
f W g into the bottom storage tank were separated nOm the gas by gravity. Air fbm the
column top flowed to a filter bag, whexe the srnaII amount of fines entraineci was
collected before the gas entered the exhaust system.
4.3 Results and Discussion
4.3.1 Observation
During the experiments with gas upward and solids downward counter-current
fïow operation, particles were seen to flow downward as an apparently dispersed
suspension. In the centre of the column particles travelled faster and solids holdups were
lower than in the w d area. Particle recirculation at the waU was obsemed, especially at
high gas velocities, where particles flew upward occasionally and solids holdups were
seen to be higher. Visual observations indicated that the flow was turbulent. Relatively
large pressure fluctuations were also observed and were probably caused by periodical
particle backmixing at the wall.
4.3.2 Flow Development and Friction Loss
Typical axial profiles of the pressure gradient (f?om top to bottom) are shown in
Figure 2, which can also be used to identify the different flow regions inside the column.
The pressure gain in the dowdow direction for the gas-solids counter-current flow in the
test column is composed of the following terms:
where gp,&gW and gp, (1-E,)& represent the contniutions of gas and soli& holdup,
AP, is the pressure loss due to solids acceleration, and MI is the Kction loss @ut a
pressure gain in the downward direction) for the gas-soli& suspension flow. When the
flow of gas and soIids is fully developed and has reached a steady state, thae is no
acceleration or deceleration of particles, and no change of solids holdup and fiction, so
the pressure gradient should be constant,
The pressure gradient profile for G, = 7.8 kg& and different gas velocities is
shown in Figure 4.2a For the range of gas velocity testeâ, the pressure gradient is seen
not to Vary much dong the colmnn h m the first measured point at 0.4 m k m the top
entrance to the bottom of the co1umn. This suggests that solids development is very
quick for low solids flux conditions (G, = 7.8 kg/m2s), and that the solids flow is hilly
developed almost instantly upon entering the column at lower gas velocities, and within
about 1.5 m for higher U, up to at least 0.39 m/s.
Increasing the soli& flux to 19.5 kg& augments the solids development
(acceleratiod deceleration) process, as shown in Figure 4.2b. In this case, the pressure
gradients are initially higher at the top of the column under 1ow gas velocities, but
decrease within a distance of about 1.5 m and then g r a d d y approach a constant dong
the column length. An initial solids developing region and a fùiiy developed region can
be identified at this G, with a bomdary around 1.5-2.0 m. In the initial flow developing
region, the pressure grradient experiences a positive deviation from that in the fUy
developed region, rather than a negative deviation as observed in CO-current downflow
systems (Zhu et al. 1995, Liu et al. 1999a). This is Wcely due to particle deceleration.
Because the particles have been accelerated by gravity in the specially designed funne1 to
their terminal velocity (Ut = 0.18 d s ) before entering the test column, and because the
particle velocity is lower than U, in the entrance region, particles initiaiiy decelerate upon
entering the column. This deceleration leads to a negative dP, value in eqn (l), which in
him results in an increase of the pressure gradient in the entnmce region. Iohnston et al.
(1999) also obsened a similar phenornenon in a co-current gas-solids downer when the
particles were accelerated to very high velocity by high velocity nozzles upon entering
the downer column. They also reported a sipnincant demase of pressure gradient in the
top section of their 10 m downer. Under higher gas velocities, the pressure gradient in the
top section is lower, consistent with those observeci in CO-current downers.
Figure 4.2 also shows the effects of gas velocity and solids flux on the pressure
gradient and the solids fiow development: (1) an increase of gas velocity ancilor soli&
flux inrreases the absolute value of the pressure gradient; and (2) an increase of gas
velocity or a decrease of solids flux both shorten the flow development length.
An increase in the pressure gradient is, among other things, related to the solids
holdup. This explains phenornenon (1) since increasing either solids flux or gas velocity
increases solids holdup. It is worth mentionhg that increasing gas flow tends to increase
the solids holdup in this system given the counter-current fiow nature in this system. For
phenomenon (2), it is obvious that decreasing solids flux would shorten the acceleration
Iength due to the reduced solids momentum. Increasing gas velocity would also reduce
the relative magnitude of the solids momentum, so it also shortens the solids acceleration
length.
In the fdly developed region, the pressure gradient can also be used to estimate
the apparent solids holdup provided that the fiction is relatively s m d and c m be
neglected. Many researchers (e.g. Arena et al. 1986) have used this method to obtain the
apparent solids holdup in upflow gas-solids risers and to characterize flow behaviour.
Those authors fond that the apparent solids holdup thus obtained can weil represent the
actuai solids holdup in the N l y developed region away nom the bottom solids
acceleration region on the riser. On the other hand, Zhu et al. (1995) commented that this
method may not be very usefùl in a dilate gas-solids co-current downflow system since
the contribution of fiïction to the piessure gradient becomes relatively larger given the
very low solids holdup (c 1%)). More recently, the same method has ais0 been employed
by Liu et al. (1999a) to characterize soli& holdup in a high-density gas-solids CO-curent
downer. They found that the apparent solids holdup inferrd h m the pressure gradient
agrees very well with the actually measured soli& holdup in the low gas velocity range
(< 5.6 ml') where the hoIdup is relatively high, but begins to underestimate the solids
holdup at higher velocities because of increased Ection Ioss. nius, the significant
deviation of the apparent solids holdup h m the actual measured soli& holdup is also an
indication of the contribution of the fiction loss to the pressure gradient-
The same method was used in this study to characterize the filction loss in the gas
upflow/solids downflow counter-curent flow system. As shown in Figure 4.3, there is
no significant diffaence between the actual solids holdup measured by the pinch valves
and the apparent soiids holdup calculated h m the pressure gradient at low solids fluxes
for the range of gas velocities tested. Beyond a solich flux of 15 kg/m2s, the fiction
between the wall and the gas-solids suspension becomes more signincant, resulting in an
obvious difference between the actual and apparent solids holdups. A positive deviation
indicates that the fiction is in the downward direction, agahst the upflowing gas which
suspends the solids in the counter-cment flow system.
It is seen that when the solids flux is higher than 15 kg/m2s the apparent solids
holdup no longer agrees with the -al soli& holdup. Even at zero gas velocity there is
still a significant difference between the actual and apparent solids holdups. It was
concluded that fiction, mainly effected by solids flux, played a dominant role in
determinhg the difference between the actual and apparent solids holdups.
43.3 Soïids H o h p in the Filly Developed Region
The actuai soli& holdups in the M y developed region as measured by the two
pinch valves are plotted in Figures 4.4 to 4.6. Increasing gas velocity tends to 6'slow
down" the solids dowdiow and therefore leads to higher solids holdup. An almost linear
relationship is observed between the gas velocity and the solids holdup for the solids flux
range tested (Figure 4.4). Increasing soli& flux at fixeci gas velocity also causes a nearly
hea r increase in the solids holdup (Figure 4.5). However, M e r increases in soli& flux
beyond some point between 15-19 kg/m2s leads to a reduction in solids holdup. This
resuit could have been caused by the increase in p d c l e velocity
In a gas-solids fiow system, solids holdups is a fiuiction of both solids flux and
mean particle velocity:
An increase in soiids flux normally leads to an increase in the mean particle velocity.
The variation of the solids holdup with increasing solids flux wouid then depend on the
rates of change of Gs and Y,. It appears that V, increases much faster than Gs beyond a
point within the range 15-19 kg/m2s, so that the general trend of g with Gs is downwards.
Below this point the mean particle velocity does not Vary much (see Figure 4.7). leaving
the solids flux as the dominating factor in the change (increase) of solids holdup. Beyond
this point, a m e r increase in solids flux leads to a very significant increase in the mean
particle velocity, to such an extent that the particle velocity increases more quickly than
the solids flux, leading to a decrease in solids holdup. In 0th- word, the particle velocity
was dorninating the change (decrease) of solids holdup.
Figure 4.6 provides the overail relationship between the solids holdup in the
counter-current gas-solids fIow system and the two main operating conditions: the
superficial gas velocity and the soli& circulation rate (flux). The generai trend is that the
soiïds holdup increases with both the solids flux and the superficial gas velocity. This is
dinerat h m the situations in bai gas-solids co-cwrent upflow risers and cocurrrnt
downflow downers, where the solids holdup increases with solids flux but decreases with
the gas veIocity. This is due to the counter-current nature of the flow system reported
here, where the gas flow resists rather than assists the soli& flow.
4.3.4 Particle Velocity and Gas-Solids Sllp Velocity
The mean particle velocity inside the fUy developed region can be caIculated
using eqn (2) fiom the measured solids holdup and flux. These r d t s an plotted against
the solids flux in Figure 4.7 and against the gas velocity in Figure 4.8. Figure 4.8 shows
an almost linear decrease in particle velocity with increasing superficial gas velocity.
This is reasonabfe given the couriter-current nature of the gas and solids flow. On the
other hand, the increase of the particle velocity is not linear with the solids flux, as shown
in Figure 4.7. The increase in mean particle velocity is initidy very small, but becomes
very significant beyond a point between G, of 15 and 19 kg/m2s.
For a known particle velocity the mean slip velocity c m be calculated using the
following equation:
As shown in Figure 4.9, for all solids fluxes tested, particle slip velocity increases more
or less lùiearly with increasing superficial gas velocity, except at the high solids flux of
19.5 kg/m2s, where the particle slip velocity is high but only increases slightly with
increasing gas velocity. An increase of gas velocity would have two effi ts on the slip
velocity: a positive one due to the increased gas velocity itself and a negative one due to
the decreased particle velocity g i v e the higher gas cornter flow velocity. It seems that
the increase in gas velocity is dominating here so that the overail effect is an hcrease of
slip velocity with increasing gas velocity.
The variation of slip velocity with solids flux is plotted in Figure 4.10. As in
Figure 4.7, the slip velocity initially increases slightly with the solids flux and then very
dramaticdiy beyond about 15-19 kg/ds. This is again due to the significant increase of
particle velocity beyond this point.
In all cases, the particle slip velocity is higher than the terminal velocity of the
FCC particles (Ut = 0.18 d s ) . This suggests the formation of particle agglomerates
inside the system. (The slip velocity may be considerd as the effective terminal velocity
of the particles in the gas-solids suspension. A slip velocity higher than the termina1
velocity for a single particle would indicate that the "effective particles" are larger than
single particles, i.e., particle agglomerates have been formeci.) From Figures 4.9 and
4.10, one c m see that the slip velocity increases with both solids flux and gas velocity. It
is specially worth noting that the slip velocity increases sharply when the solids flux is
raised over a value amund 15-19 kg/m2s, suggesting severe particle agglomeration
beyond this point.
It appears that there is some dramatic change to the gas-solids flow at a G, higher
than 15-19 kg/m2s, which causes significant changes in solids holdup, particle velocity
and slip velocity.
4.3.5 Choking
In a gas-solids CO-current upflow system, decreasing gas velocity unda a given
solids flux or increasing solids flux at a fixed gas velocity cm both Lead to unstable gas-
solids operation, refmed to as choking @i et al. 1993). In this study with the gas-soli&
counter-current flow system, the choking gas velocity is seen to d-e with the sol&
flux. At a lower solids flux of G, = 4.5 k&s the choking velocity is 0.39 mls. At othet
higher soli& fluxes the choking velocity is slight lower, about 0.33 mls. At any gas -
velocity above this key point, particles tend to be entraineci by upflowing gas, which
blocks particles flowing down. Counter-flow choking velocities are larger than the mean
partick terminal velocity which was calculated as 0.18 d s due to the existence of the
particle agglomeration.
An increase in solids flux at a Gxed gas velocity also results in unstable gas-solids
operation, in 0th- word, there is a dramatic change to the gas-solids flow when G, is
higher than 15 kg/m2s. This can be M e r discussed in relation to the experiment &ta at
G, = 19.5 kglm2s by cornparison with choking phenomena in gas-soiids upfiow.
Choking has been found to depend on the properties of both gas and solids
particles, as weil as on the size and geometry of the column containhg the flow system
(Zenz 1949, Yousfi and Gau 1974). As summarized by Bi et al. 1993, three types of
choking were identifieci for gas-solids upflow. The "accumulative chokingyy, or type A
choking, which will occur nrSt with decreasing gas velocity, was attributed to an abrupt
change in voidage (Matsen 1982, Yerushalmi and Cankurt 1979, Brereton 1987, Rhodes
1989). The solids circulation rate (solids flux) at the type A choking velocity co~esponds
to the saturation canying capacity (Zenz and Weil 1958, Wen and Chen 1982, Sciazko et
al. 199 1). To fbd the relatiomhip between choking phenomena in counter-cumnt and CO-
current upflow, the Bi and Fan (1991) equation was used to calculate the choking velocity
for upflow, at G,= 19.5 kg/&, yielding Uch=2.12m/s. Assurning slip velocity
= 1.1 mfs was the same as for counter-current flow at the choking point, the particle
velocity at the chokmg point for the co-current upflow condition is estimated as foUows:
The particle velocities tested in counter-current flow at G, = 19.5 k&s are in the range
of 0.80 - 1.12 d s which are close to the partice velocity at the chobg point in CO-
current upflow. The soli& flux G, = 19.5 k&s corresponds to an abrupt change in
solids holdup at this point and is similar to the saturation carrying capacity in co-current
upflow. Because of the very naww range of gas velocities, the effects of gas velocity on
cholong at a specinc solids flux are not signincant, It is concludeci that, for a given solids
flux at the choking point, the flow behaviour seems to depend on the particle velocity at
the choking point for both counter-current flow and co-current upflow. In other word,
with increasing solids flux, the choking occurs at about the same particle velocity whether
solids flow upward or downward. The chobg phenornenon at G, = 19.5 kglm's can be
used to explain the ciramatic change in the gas-solids flow at a G, higher than 15 kg/m2s,
which causes signincant changes in solids holdup, particle velocity, and slip velocity.
4.4 Conclusions
Flow behaviour in a gas upwd-solids downwards counter-current fluidized flow
system was stuclied for the fint thne. The following conclusions were found:
(1) Particles were seen to flow downward as an apparently disperseci suspension. Particle
recirculation at the wail was observeâ, especially at high gas velocities, where
particles flowed upward occasionally and solids holdups were seen to be higher.
(2) Typical axial profiles of the pressure gradient were discussed and used to identify
initial solids developing region and a fully developed region. The experimental
results indicate that the pressure gradient provides a simple method to estimate soiids
holdup without incurring large exrors when the soli& flux is not higher than
approxhately 15 kg/m2s for all operating gas velocities in the fùliy developed
region.
(3) hcreasing gas velocity at a given solidï flux always leads to a hair increase in
soli& holdup. Increasing soiids flux at fixed gas velocity a h causes an increase in
the solids holdup. However, further increasîng the solids flux beyond some point
around 15 kg/m2s leads to a reduction in solids holdup.
(4) The choking gas velocity is seen to decrease with the SOI& flux. The chokhg
phenornenon at Gs=19.5 kgkm2s can be used to explain the dramatic change to the
gas-solids flow at a G, higher than 15 kglm2s, which causes significant changes in
soli& holdup, particle velocity and slip velocity.
The authors are gratefbl to the hanciai support fiom the Naturai Sciences and
Engineering Research Council of Canada for this research project.
Notations
g: acceleration due to gravity, m/z G,: solids flux, kg/m2s
Hi distance h m the top of test column, m
P: pressure: Pa
Uk supdcial gas velocity, d s
U s : particle slip velocity, m/s - V, : mean actuaf gas velocity, m/s - V, : mean particle velocity, m/s
6: tolerance for pressure gradient variation, palrd
LW distance between the two pressure taps, m
dP,: pressure loss due to particle acceleration, Pa
dPk: pressure loss due to gas phase fiction, Pa
dP/,: pressure loss due to solids phase fiction, Pa
4: soli& holdup
4: density of gas phase, kg/m3
p,: density of solids phase, kg/m3
Bai, D., Jin, Y., Yu, 2. and Gan, N, (1991), ''Radial Profiles of Solids Concentration and
Velocity in a Concurrent Dowdiow Fast Fluidized Bed (CDFFB)", CircuIating
Fluidked Bed Technology m, (eds. P. Basu, M. Hono and M. Hasatani), pp.157-162,
Pergamon Press, Oxford.
Bai, D., Jin, Y., Yu, 2. and Zhu, J-X (1992), 'The Axial Distri'iution of the Cross-
Sectiondy Averaged Voidage in Fast Fluidized Beds", Powder TechnoC., = 51-58.
Bi, H. T., Grace, J. R and Zhu, J-X. (1993), "Types of Chokuig in Vertical Pneumatic
Systems", IntC J , MuZtiphase Flow, 19.1077-lO92.
Bi, H. T., Grace, J. R and Zhu, 1-X. (1995), "Regime Transitions Mécting Gas-Solids
Suspensions and Fluidized Beds", C'hem. Eng. Des. Dev., a 154-161.
Clift, R, Grace, J. R and Weber, M. E. (1978), BubbZes, Drops and Particies, Academic
Press, New York.
Geldart, D. (1973), "Types of Gas FIuidization", Powder Technol., 2,285-292.
Johnston, P. M., de Lasa, H. 1. and Zhu, J- (1999), "Axial Flow Structure in the Entrance
Region of a Downer Fluiâized Bed - Effects of the Distributor Design", Chem.
Eng. Sei., in press.
Liu, W., Zbu, J-X., and Beeckmans, J. M. (1998), "Characterizaion of High Density Gas
Solids Downflow Fluidized Reactor", Powder TechnoZ., to be submitted (Chapter 3).
Matsen, T. M. (1 982), 'Mechanisms of Choking and Entrainment", Powder Technol., 32. 21-33.
Wang, Z., Bai, D. and Th, Y. (1992), 'Hydrodynamics of Cocurrent Downflow
Circuiating Fluidized Bed (CDCFB)", Powder TechnoL, 271-275.
Yerushalmi, J. and Cankurt, N. T. (1979), 'Turther Shidies of the Regime of
Fluidization", Powder Technol., 24.1 87-205.
Yang, Y-L., Jin, Y., Yu, 2-Q., Zhu, LX. and Bi, H-T. (1993). 'Zocal Slip Behaviors in
the Circulating Fluidized Beâ", MCnE S p p . Ser., =(296), 81-90.
Z e t q F. A. (1949), 'Two-Phase F1uidized-Solid Flow", Ind. Eng. Cnm., Pi. 2801-2806.
Zenz, F. A. and Weil, N. A. (1958). "A Theoretical-Ernpirical Approach to the
Mechanism of Particle Entrainment h m Fluidized Beds", NChE JI, &, 472479-
Zhu, J-X. and Bi, H-T. (1995), "Distinctions between Low Density and High Density
Circdating Fluidized Beds", Con. J. Chm. Eng., 23.644-649.
Zhu, J-X-, Jin, Y., Yu, 2-Q., Grace, J. R and bsangya, A. (1995), "Cocurrent DoWaaow
Circulating Fluidized Bed @owner) Reactors - A State of the Art Review", Cun- J.
Chem. Eng., a 662-677. Zhu, J-X. and Wei, F. (1996), "Reecent Developments o f Downer Reactors and other
Types of Short Contact Reactors", Flicdikution Ym, (eds. J. F. Large and C.
Laguerie), pp. 50 1-5 1 O, Engineering Foundation, New York.
Figure Captions
Figure 4.1 : Schematic of the experimental apparatus
Figure 4.2~~: Pressure gradient profiIes along the wlumn at G,=7.8 ks/rn2s
Figure 4.2b: Pressure gradient profiles along the column at Gs=19.5 kglm2s
Figure 4.3 : Cornparison of actuai and apparent solids holdup
Figure 4.4: Soli& holdup in the fally developed region as a hction of U,
Figure 4.5: Solids holdup in the fully developed region as a fhction ofGS
Figure 4.6: Gas-soç-solids counter-current ffow operating range
Figure 4.7: Mean particle velocity as a fünction of Gs
Figure 4.8: Mean particle velocity as a hct ion of (Tg
Figure 4.9: Mean particle slip velocity as a bction of U.
Figure 4.10: Mean particle slip velocity as a hction of G,
w CYCLONES
1 1 UPPER I STORAGE TANK
FUlDlZED BED FEEDER W
FEEDING FUNNEL
PARTICLE RECYCLE LlNE
TEST COLUMh ID 3" 5 m High
- TO FILTER BAG
BOTTOM STORAGE TANK L] ROTAMETER
n
- - VALVE 4
V E E T V Z E ~ $. AIR
Schematic diagram of the fluidized feeder
Figure 4.1 The schematic of the experimental apparatus
Distance from the top, H (m)
Figure 4.2a Pressure gradient profile almg the column at G17.8 kgf&
O 1 2 3 4
Distance from the top, H (m)
Figure 4.2b Pressure gradient pronle dong the column at Gs=15.0 kg/m2s
0.000 0.005 0.010 0.015 0.020 0.025 0.030 0.035
Actual solids holdup,
Figure 4.3 Cornparison of the actuai and the apparent solids hoIdup
Superficial gas ualocity, U, ( m k )
Figure 4.4 Soiids holdup in the W y developed region as a function of U,
Solids flux. G, (kg/m2s)
Figure 4.5 Solids holdup in the W y developed region as a fiuiction of G,
14 -
. 1
4 - 2 - I
O m 1 1 , I I . 1 L
0.0 0-1 0 2 0.3 0.4 0.5 0.6
Superficial gas velocity, Up (mls)
Figure 4.6 The operating range for gas-solids countercurrent flow
Solids flux. Gs ( k g h 2 s )
Figure 4.7 Mean particle velocity as a fùnction of G,
Supeficial gas \ialocity, Ug (mis)
Figure 4.8 Mean particle velocity as a fimction of U,
Superficial gas velocity, U, (mls)
Figure 4.9 Mean particle slip velocity as a fùnction of U,
Solids flux Gs (kglm2s)
Figure 4.10 Mean particle slip velocity as a fiuiction of G,
CEAPTER 5. rFT4DLACTERIZATION OF THE FLOW REGIMES
AND UNIF'IED REGXME D I A G W E N E R A L DISCUSSION
To characterize gas-soli& flow, the basic flow regimes and corresponding
transitions have been studied for several decades. Various flow regïme maps made by
different approaches were presented. From previous works in the literature, there are at
least five different fluidization regimes d e W . particdate fluidization (for group A
particles ody), bubbling (slugging) fluidization, turbulent fluidkation, fast fluidization,
and pneumatic transport. In an mly study, Zenz (1949) proposed the dense fluidization
and the CO-curent pneumatic fiow regimes in a flow diagram in which pressure gradient
was plotted against superficial gas velocity. Simitar flow regime maps were presented by
Yerushalmi et al. (1976, 1978, 1979) Ui which bed voidage was ploîted against
superficial gas velocity and gas-particle slip velocity, to show the transitions among the
packed bed, bubbling fluidization, turbdent fluidization and fast fluidization regimes.
The regime map presented by Li and Kwauk (1980) also plots voidage against superficial
gas velocity. Squires et al. (1985) expanded such a map to include the pneumatic
Cransport regime and choking points, and this was M e r modined by Rhodes (1989).
Grace (1986) extendd and modified the approach of Reh (1971) to propose a unified
regime diagram based on fiterature data to show the operating ranges of conventional
fluidized beds, spouted beds, cimilating beds and transport systems. To combine gas-
solids fluidized beds and pneumatic transport, a flow regime diagram, deveioped from the
flow regime diagram of Grace (1986), was presented by Bi and Grace (1995). Bai et al.
(1993) proposed a flow regime map for circulating fluidized bed in which solids flux was
plotted against the superficial gas velocity and two flow regimes, fast fluidization and
dense phase conveying, were de- Fiow regimes were aIso chiuacterized by Bai et al-
(1996) using the différentia1 pressure fluctuations. The study on the flow regimes for the
conventional fluidized bed by Saxena (1990) using the local heat-transfer data were also
reported.
For pneumatic transpor&, the flow regime diagrams plotted by superficial gas
velocity vs- sofids flux were proposed by Leung (1980), KlinPng (1981), Yang (1984)
and Mok et al. (1989) in which gas-soli& transport was =ded into dense-phase and
dilute-phase flow regimes. Takeuchi et aL (1986) proposed a flow rnap based on theu
experùnental data to define the boudaries of f a fluidization. This flow regime map was
later modined by Bi and Fan (1991). Hirama et al. (1992). on the 0 t h hand, tried to
extend such a diagram to the transition from high-velocity to low-velocity fiuiâization.
From the above literature, one cm see that the extensive work on the flow regimes
for gas-solids fluidization and upwiml sansport has been c b e d out. However, no work
has been published on the flow regimes for gas-solids co-current downfiow and for gas-
solids conter-current flow. In this work, a flow regime rnap was proposed based on the
experimental fïndings.
5.1 Co-Current Downward FIow Regimes
5.1.1 Pseudo-Aggregative and Pseudo-Particalate Flow Regimes
Figure 5.1 shows mean particle velocity in the M y developed region of gas-solids
CO-current domward flow as a fùnction of superficial gas velocity at Mirent solids
flues. The data obtained h m wide operating range, solids fluxes h m 90 to 3 0 kg/m2s
and solids holdup h m 1% to 11% , are plotted. For a specific solids flux, mean particle
96
velocities remain constant in a small range of the low superficial gas velocity and then,
Superficial gas velocity, Ug ( d s )
Figure 5.1 Mean particle velocity as a fiuiction of superficial gas velocity
after a critical value of gas velocity, they increase linearly with increasuig superficial gas
velocity under aU soli& flux conditions. The nrrning points in Figure 5.1 wrresponding
to a range of 0.65-1.35 mls of superficial gas velocities are the critical gas velocities
defined as U, which divides the flow into two sections for different solids ffuxes:
pseudo-aggregative and pseudo-particdate flow ~egimes. Those two flow ngimes were
observed in all the gas-soli& co-cuuent downflow experiments studied.
In the pseudo-aggregative flow regime, mean particle velocity ;exnains constant
and the measured mean solids holdups are at theu maximum values for a given solids
flux. The particles were seen to flow downward in apparently continuous strands and
clusters in the core region and dong the waU In the centre of the column, particles
travelled faster and solids holdups were lower than in the wall area. In this flow regime,
solids are only accelerated by gravity and the gas flow exists as a kind of resistance which
would counter-balance particle flow to reach a steady state.
The pseudo-particdate flow regime appears immediately after the pseudo-
aggregative flow regime when the gas superficial velocity was beyond U,. In this flow
regime, the mean particle velocities inmeases linearly with superficial gas velacity and
solids holdups decreases with increasing gas velocity. Both gas and solids flow
downward and more ULLiform gas-solids suspension without apparent clusters was
observed. The flow structure becomes homogeneous compared to the pseudo-aggregative
flow regime. The mean particle velocity is only detennined by the superficial gas velocity
and the specific particle properties (mean particle size and density, particle size
distribution), but independent of the solids £lux, although particles are accelerated by both
gravity and the gas flow to reach the steady state. In this flow regime, particles seems to
be fUy disperseci because of the existence of relatively high volume of gas.
In gas-SOUS iipnow circulating fluidized bed, the mean particle slip velocity is
always higher than the particle terminal (fiee fd) velocity because fine particles have the
tendency to aggregate into particle clusters to maintain a stable state (Grace and Tuot
1979) or to minimize energy dissipation (Li et ai. L988), as well as to reduce the drag
force between gas and particles (Zenz and Othmer 1960, Fujima 1991).
In the gas-solids downtlow system, the mean particle slip velocity is defined by
the following equation:
From all experimental data in the pseudo-particdate fiow region, mean particle velocity
was plotted in Figure 5.2 as a hction of actual gas velocity. The figure shows that the
mean particle velocities in this flow regime have a linear relationship with actual gas
velocity and was independent of solids flux and solids holdup. The linear regression for
the data gives the foliowing:
U* in eqn (5.1) could be determineci to be 0.57 mis by comparing eqn (5.2) and eqn
(5.1). This U&, of 0.57 mis could be considered to be an apparent taminal velocity of
particles for gas-solids CO-current downflow. For this study, it is seen that, in the pseudo-
particdate flow regime, Us,* h m eqn (5.1) is a constant or the apparent terminal
velocity of particles remains constant. In other word, it is thought that the solids
suspension approached fdly dispersai.
Figure 5.2 The relationship between actual particle velocity and actual gas velocity in the pseudo-particdate flow regime
Apparently, the particle slip velocity, even in the pseudo-particdate flow regime,
is much higher than the single particle terminal velocity. This enhancement of particle
slip velocity can be attnauted to the formation of particle agglomeration. Due to a
constant particle slip velocity in the pseudo-particdate flow regime, the ratio of the
partide slip velocity to the particle terminal velocity, Usl& , is also constant as 3.17,
where &=O. 18 m/s as calculateci by using the method pmposed by Clift et al. (1 978). The
inaease of gas flowrate does not cause a significant change in the degree of particle
aggiomeration in the pseudo-particdate flow regime. Unlike the pseudo-aggregative flow
regime and the flow regimes in the riser, in which the value of varia with the
specinc solids holdup (Matsen 1982). in the pseudo-particdate fiow regime in the
downer is independent of gas velocity and soli& flux
5.1.2 Determination of U, by DUTerentiai Pressure Pi~ctuatiois
Figure 5.3 shows the standard deviation fluctuation of the differential pressure,
obtained from dinerentiai pressure transducers, in the Mly developed region for the co-
current downûow as a fùnction of superficial gas velocity at different solids flux. It is
seen that the dif3ierenta.i pressure fluctuation initiaily decreases for ail given solids flux
condition within a range of superficial gas velocity and then gradually approaches a
constant value with increasing gas velocity. The superficial gas velocities correspondhg
to the tuming points of those c w e s were at around 0.55 m/s at Gs 23 kglm2s, and 1.08
mls at Gs 90 kg/m2s to 210 kg/m2s, which are in h e with the U, values indicated in
Figure 5.1. In the pseudo-aggregative flow regime, the degree of partide agglomeration,
which was reflected by the fluctuation of the différentid pressure signal, would decrease
with an increase of gas velocity. The higher particle agglomeration wouid result in a
higher fluctuation of the differential pressure because of the existence of bigger strands
and clustm. When gas velocity is beyond U,& comsponding to the turning point, the
constant differential pressure fluctuation indicates that the gas-soli& flow is in the
pseudo-particdate flow repime. From Figure 5.3, it muld be concludad that high soi&
holdup under high solids flux resulted in high pressure fluctuation at a given gas velocity.
This is in agreement with the results of gas-solids upward fluidized bed @ai et ai. 1996).
As discussion above, U, can be detennined either h m the measurment of the
solids holdup (mean particle velocity) in the pseudo-aggregative m g h e or the dinkrential
pressure fluctuation.
Superficial gas velocity, U, mis
Figure 5.3 Dinerential pressure fluctuation as a hction as U,
During the experiments with gas upward-soli& downward counter-ciment flow
operation, particles were seen to flow downward as an apparently dispersed suspension.
In the centre of the column, particles travelied faster and solids holdirps were Iower than
in the waiI region. Particle recirculation at the wali was observe& especially at high gas
velocities, where particles flow upward occasionally and solids holdups were seen to be
higher-
The particle slip velocity for both the high density dowdow and the counter-
cment downfîow is plotted against the superficial gas velocity in Figure 5.4. The particle
slip velocities in the counter-current flow for all operating conditions (except at the
choking point) are in the range of 0.3-0.6 mis, most of which are lower than the apparent
terminal velocity of the pseudo-particdate flow regime for the high density downflow,
but stiil higher than the single particle terminal velocity. It can be concluded that the
degree of particle agglomeration in the counter-curent h w is lower than those in the
high density downfiow.
As the operating limitation of counter-cment flow, the solids holdups under
steady operating conditions are in a small low value range. The cornparison of solids
holdups for both counter-current and CO-current downflow using the apparatus of the
same geometry was shown in Figure 5.5. To have a same value of the solids holdup for
counter-current flow, only one third of the solids flux for cocunent downflow was
employed. This result is thbught to be very important in the design of a chernical reactor
where a low solids circulating rate and a high solids holdup are expected.
Superficial gas veiocity, U, (mis)
Figure 5.4 Particle slip velocity in both downflow and counter-flow
5.3 Cornparison of High Density Dovvnfïow and Counter-Cnrrent Flow with Upfiow Flow Regimes
5.3.1 DLBierentiai Pressure Fïuctaations and Soïids Holdup
The dineratid pressure fiuctuaîions have been used by some researchers to
characterize the gas-solids upflow fluidized bed in bubbling, turbulent, fast fluidization
and pneumatic transport regimes. Yerushalmi and CO-workers (1978). in an extensive
study, fond the transition to hubulent fiuidization by the amplitude of pressure
fluctuations across the bed- Fan et al. (1981) reporteci that the amplitude of the pressure
fluctuations in a bubbhg beâ is related to both the bed density and the size of bubbles
which are the sources of the pressure fîuctriations. For fast fiuidized bed, Bi (1994)
concluded that the standard deviations of Merential pressure fluctuations has a sole
relationship with the apparent solids density for aIi operating conditions. Bai et al. (1996)
proposed a flow regime map by using the standard deviation of pressure ffuchiatio~~~.
Superficial gas velocity (downflow), Ug (mls)
Superficial gas velocity, U, (mls)
Figure 5.5 The cornparison of solids holdup between co-current downflow and counter-cuxrent flow
For this study, the standard devlations of the diflrerentiaï pressure for the co-
current high density downfhw and the counta-current flow fluidized bed in fWy
developed regions, as weli as the upflow (Bai et al. 1996). were plotted against mean
solids holdup in Figure 5.6. It is shown that the standard deviations of differential
pressure changed with gas velocity and solids flux for the same solids holdup, but the
general trend of standard deviatiom of the differential pressure was obviously observai
with changing
Riser Downer
6.001 0.01 0-1
Solids holdup, E
Figure 5.6. Differential pressure flucîuation as a hction as solids holdup
solids holdup. The values of ai i data h m both the co-current high density downfiow and
the counter-cunent flow are consistent with thosc obtained h m upflow at the same
soli& holdup. It can also be concludeci that the co-current high density downflow and the
couuter-cument flow correspond to the f& fiuidization and the pneumatic transport
regimes in the iipflow systems. as proposed by Bai et al. (1996). But no fkher b
demarcation between the pseudo-aggregative ff ow and the pseudo-particdate flow regime
for CO-current high density downfiow could been found h m Figure 5.6.
5.3.2 Mean Voidiige vs. Partick Sïip Vetoeity
Figure 5.7 plotted the partic1e slip velocity against the mean solids holdiip. As
presented by Yemshalmi et al. (1979) in a similar plot for the fast fluiàization and the
pneumatic transport regimes, the particle slip velocity increased sharply with increasing
solids holdup and then approached a constant in a small range of very low solids holdup.
For the CO-current high density downflow. Figure 5.7 shows the particle slip velocity in
the CO-current high density downfiow can &op to approach a constant quickly with
changing solids holdup. The general trends are similar to those in the fast fluidization and
the pneumatic transport regimes for upfiow system. For CO-current hi& density
downfiow, the pseudo-aggregative and pseudo-particdate flow regimes seems to
correspond to the fast fluidization and the pneumatic transport regimes in the upflow
system.
O .O0 O .O2 O .O4 0.06 O .O8 0.1 O 0.1 2
Solids Holdup, E
Figure 5.7 Particle slip velocity as a fhction as solids holdup in CO-cment dowdow
The solids holdups correspondhg to the omet of the pseudo-particdate fiow
regime, in which particle slip velocity remains constant, varied in a large range at
different solids fluxes. This is due to no restriction of "gas carrying ability" for CO-
current downflow and both gas and solids flow in the direction of gravity, so the
transition between regimes does not have to correspond to the same value of soli&
holdup. Furthemore, it is related to the value ofsolids flux.
Based on the discussions above, the flow regimes in the CO-current high density
downflow are expected to exhibit the same typa of hydrodynamic behaviour as on the
fast fluidization and pneumatic transport regimes in the upflow systern. To complicate
matters M e r , the pseudo-aggregative flow regime in the cocument hi& density
downflow is smiilar to the fast fluidization regime; while the pseudo-particdate flow
regime is close to the pneumatic transport regime.
5.4 Unified Flow Reghe Diagram
By combining the literatwe data for FCC in the conventional fluidization and the
CFBs (risers) @uang and Zhu 1998, Bai et al. 1993) and the present study in both the
high density downfiow and the cou11teri:urrent downfiow, an unined flow regime
diagram is proposed. Figure 5.8 shows that, within the four quadrants formed by U, as x-
axis and G, as y-axis, the upflow operation (riser) is in the first quadrant and the
downflow and the counter-current flow operations are located in the third and the fourth
quadrant. This map first gives a generd picture of fluidization, which includùig ali types
of fiuidized beds. A clear "operathg window" for FCC particles is proposed. The high
density CO-current downflow and counter-cment flow overcome a common shortcornhg
of the high velocity riser and downer: very low volumehic concentration (holdup) of
solids. The results of this study is very important in the chernical reaction where a high
solîddgas ratio is required, since the reaction intensity is limited by the lower solids
concentrationn The new unified flow regime diagram extends our cumnt knowledge to
wider operating ranges.
Figure 5.8 Unified flow regime diagram
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113
CHAPTER 6. CONCLUSIONS AND RECOMMENDATIONS
6.1 Conclusions
The principal conclusioll~ of this study on gas-solids CO-current downflow are as
follows:
Along the downer column, a particle acceleration region and a fiilly developed region
were identined under ali operaihg conditions tested. The two regions can be
established by e x a . g the mainaed pressure gradient profiles, with a constant
pressure gradient signifying the M y developed region.
The experimental d t s indicate that the pressure gradient provides a simple and
reliable method to estimate the solids holdup without incming large errors when the
gas velocity is lower than appmximately 5.6 m/s at dinerent soli& fluxes in the M y
developed region. Beyond this gas velocity, waU fiction leads to a signifïcant
underestimation of the actual solids holdup when using the pressure gradient
method.
Two diBerait flow regimes have been identifieci in the developed region, a constant
and high density pseudo-aggregative flow regime at low gas velocities and a low
density pseudo-particulate flow regime at high gas velocities, with a boundary within
the range U, = 0.6- 1 -3 m/s for different solid fluxes.
High-density downflow operation is defined as operation in the pseudo-aggregative
flow regime. In the hi@-density flow regime, particle velocity remains constant for a
given solids flux and is independent of the gas velocity, and the slip velocity is very
high with very significant particle aggiomeration.
In the more dilute pseudo-particdate flow regime; the gas-particle slip velocity
remains constant and no particle strands and large particle clusters are observed. The
constant sIip velocity suggests thaî equili'brium has been reached in particle
aggregation so that it no longer changes with the gas velocity. The particle velocity
iucreases 1ùiearly with gas velocity @en the constant slip velocity. Conse~uently,
the solids holdup decreasa with increasing gas velocity in this regime, as reported
previously in other riser and downer systems.
The operating wuidow for gas-solids downflow in the existing experimental
apparatus has been mapped. For the experimental conditions employed in this study,
solids densities above 10% c m be reached in the high density fiow regime.
Cornparison of the results obtained in this study with those h m an upflow riser
shows some inherent similarities between downfiow and upflow gas-solids CO-
current flow systems. The variations of solids holdup with gas velocity becomes
consistent for both systems when the slip velocity is deducted fiom the riser gas
velocity and is added ont0 the downer gas velocity.
Flow behaviour in a gas upward-solids downwards wunter-current fluidized flow
system was studied for the first t h e e The following conclusions were found:
Particles were seen to flow downward as an apparently dispersed suspension. Particle
recirculation at the wall was observed, especially at high gas velocities, where
particles flowed upward occasionally and solids holdups were seen to be higher*
Typical axial profiles of the pressure gradient were discussed and used to identify
initial sol& developing region and a fblly developed region. The experimental
results indicate that the pressure gradient provides a simple method to estimate solids
holdu~ without incurring large emrs when the solids flux is not hifier than
appmximately 15 kg lds for afl operating gas velocities in the M l y developed
region.
Increasing gas velocïty at a given solids flux aiways leads to a linear increase in
solids holdup. I n d g solids flint at nxed gas veIocity aïs0 causes an increase in
the solids holdup. However, increasing the solids flux beyond some point
around 15 kg/m2s leads to a reduction in solids holdup.
The choking gas velocity is seen to decrease with the soli& Bux. The choking
phenornenon at G'=19.5 kg/m2s can be used to explah the ciramatic change to the
gas-solids flow at a G, higher than 15 k g d s , which causes signincant changes in
solids holdup, particle velocity and slip velocity.
The foIlowing are genael conclusions of the flow regimes in high deosity co-current
dowdow and comter-curzent flow:
(1) For the high density co-current downfiow, the cntical gas velocities de- as U,
divide the flow into two sections for dinerent soiids fluxes: pseudo-aggregative and
pseudo-particdate flow regimes. U, cm be deteTmined either fkom the measinexnent
of the solids holdup in the pseudo-aggregative regime or the differential pressure
fluctuation.
(2) The comparison of the high density d o w ~ w and the conter-current flow regimes
with the upflow flow regime were made by using the differential pressure fluctuations
and the partide slip velocity. The flow regimes in the cocurrent high density
downflow and the counter-current flow are expected to exhibit the same types of
hydrodymmic b e h a . 0 ~ of fmt fluîdization and pneumatic transport regimes in the
upfl ow system-
(3) Finally, an unifieci overaii flow regime diagram is proposed. This map nrst gives a
general picture of fluidization which including ail types of fluidized beds. A clear
"operathg window" for FCC particles is proposed. The d t s of this study is very
important in the chernical reaction where a high soiiddgas ratio is required, since the
reaction intensity is limiteci by the Iowa solids concentration. The new unifid flow
regime diagram extends our cunent knowledge ta wider operating ranges.
6.2 Recommendation
1- It would be helpfd to improve the design of the equiprnent for the high density CO-
current downflow and CO-current fiow so that the test operation can be done on the
continuos basis as same as the circulathg fluidized bed reactor. The new continuos
operation, which will have more potential industrial applications, need to be fiirther
studied with the effects of the back pressure and inventory.
2. The geometry of experimental apparatus wil l mect the flow characters significantly..
For this study there is a obvious wall efféct in t&e test wlumn of 0.025 m diametter.
Shidies in larger diameter column would be needed for m e r research with more
details of the radial profiles.
3. Future studies shodd be undertaken to test the effects of the different solids
properties such as particle densityy particle size, and the distribution of the particle
size.
4. Further research work should also be conducteci in the fluiàized feed system. For the
column of larger diametaY the solids distriiutor need to be studied.
APPENDIX
Appendix-1 Cdibration lineu curve for each düferentiai pressure transducer
used in the erperiment
Transducer No Range (mm H2O)
Best fit linear equation
Appendix-2 Esperimenûü dam for figures used in manuscripts
3.1 High Density Co-Current Downilow
Distance from Ug-5-56 m/s
Pressure gradient, dp/dh (Mm)
Pressure gradient, dpldh (Palm) Gs=9û kglm2s
Solids Holdup, E
Mean Parücle Slip Velocity, Vp (mis)
Pressure gradient, PM (Pdm) G W . 8 (kd-)
Preuure gradient, PM (Palm) Gs=iS.O (kNrn2s)
Solids Holdup, E
Mean Particle Slip Valocity, Vp (ds)