Heat Transfer Studies in a Gas-Solids Downflow Circulating Fluidized Bed (Downer)
by
Ying L. Ma
Faculty of Engineering Science
Deparmient of Chernical and Biochemical Engineering
Submitted in partial hlfilrnent of the requirements for the degree of
Master of Engineering Science
Faculty of Graduate Studies The University of Western Ontario
London, Ontario April, 1998
Q Ying L. Ma 1998
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Heat transfer between gas and solids flow suspension and the surface immersed in
the bed was studied in a 9.3 m high, 100 mm inner diameter gas-solid CO-current
downflow fluidized bed (downer) with 65 Fm FCC particles. The radial and axial
distributions of heat tnnsfer coefficients between the suspended surface and the gas-
solids flow suspension were obtained using a miniature cylindrical heat transfer probe
under different operating conditions. The gas-solids miring eficiency in the entrance
region of the downer was studied based on a thermal method, with which the gas-solids
contact efficiency was estimated through measuring the temperature change of hot
fluidized air in the bed by a shrouded thennocouple. The solids concentration and particle
velocity, which are considered as hKo of the most influential factors to the gas and solids
flow, were measured by two separate optical fiber probes. The results show that heat
transfer and gas-solids mixing behavion are al1 controlled by the hydrodynamics in the
bed, and there is a close relationship between the heat transfer coefficient or gas-solids
contact efficiency and the solids suspension density. Utilizing this characteristics and
through studying the heat transfer behavior with different types of distributors, the effects
of distributor stmcture on the flow pattern and development in the entrance region of the
downer were also investigated and characterized. The operating conditions and entrance
structure (dishibutor) have been found to have significant effects on the gas and solids
flow structure, gas-solids mixing and heat transfer behaviors in the downer.
Keywords: Heat transfer, Circulating fluidized bed, Contact efficiency, Discributor,
Hydrodynarnics, Downer
CO-AUTHORSHIP
Title: Experimental Study of Heat Transfer in a Co-Current Downflow Fluidized Bed
@owner)
Authors: Y. Ma and J.-X. Zhu
Al1 portions of the experiment work were undertaken by Y. Ma under the guidance of
advisor J.-X. Zhu. Al1 cirafts of this manuscript were written by Y. Ma. Modifications
were carried out under the supervision of the advisor. The final draA was subrnitted to the
j oumal Chemical Engineering Science.
Title: Characterizing Gas and Solids Distributors with Heat Transfer Study in a Gas-
Solids Downer Reactor
Authors: Y. Ma and J.-X. Zhu
The experiment were undertaken by Y. Ma under the guidance of advisor J.-X. Zhu. The
designs of the circulating fluidized bed system and the distributors were assisted by H.
Zhang and P. M. Johnston. Al1 drafts of this manuscript were written by Y. Ma and
modified by the advisor. The final draft was submitted to Chemical Engineering Journal.
Title: Gas-Solids Contact Efficiency in the Entrancr Region of a Co-Current
Downflow Fluidized Bed
Authors: Y. Ma, J.-X. Zhu, and H. Zhang
Al1 portions of the experiment were undertaken by Y. Ma under guidance of advisor J.-X.
Zhu. This work was assisted by H. Zhang who helped to install the electrical air heater
and P. M. Johnston who designed the distributors used in this study. Al1 drafts of this
manuscript were written by Y. M a Modifications were done under the supervision of the
thesis advisor before the final draft was submitted to Chem. Eng. Res. Des.
ACKNOWLEDGMENTS
The author is sincerely grateful to her advisor, Professor J. Zhu for his
guidance, encouragement and continuous support throughout the period of study
and the completion of the project.
Sincere thanks to my fellow student and fiend Mr. H. Zhang, who helped
me build the electrical gas heater and some other measurement apparatus and
provided many assistantance during the experiments.
To al1 my colleagues P.M. Johnston, F. Wang, J. Bal1 and W.-D. Liu, W.
Huang, 1 extend gratehil appreciation for their helpful advice and cooperation in
operating Our downer equipment.
This research was carried out under the financial support from NSERC,
the Natural Science and Engineering Research Council of Canada.
Finally, thanks is extended to my husband for his great support and help
during this two years.
Page
TABLE OF CONTENTS
CERTIFICATE OF EXAMINATION .............-....-.....-....-.....-.................................. i i
.-. ABSTRACT .......................................-...--.............-....-......-..................................... 111
CO-AUTHORSMP.. ................................................................................................. .iv
ACKNO WLEDGMENTS.. ....................................................................................... .v
TABLE OF CONTENTS ......................................................................................... vi -. - LIST OF TABLES .................................................................................................... viir
LIST OF FIGURES .................................................................................................. ix
NOTATION ............................~................................................................................. xv
C W T E R 1 . INTRODUCTION ..................... .. .................................................. 1
CHAPTER 2. LITERATURE REVIEW .............................................................. 5
2.1 . Circulating Fluidized Beds ......................................................... 5
2.1.1. Fundarnentals and Applications of Co-Current
................................. Downflow Circulating Fluidized Beds 5
2.1.2. Hydrodynamics in Upflow Circulating Fluidized Beds .......... 7
2.1.3. Hydrodynamics in Downflow Circulating Fluidized Beds ..... 8
2.1.4. Gas and Solids Flow Structure at the Top Entrance Zone
in the Downer ...................................................................... 16
......................................... 2.1 S. Downer Distributor Configuration 1 7
2 -2. Heat Transfer in Up flow Circulating Fluidized Beds ....................... .2 1
................................................................ 2.2.1. Particle Convection 2 3
2.2.2. Gas Convection ....................................................................... 27
2.2.3. Radiation ................................................................................. 30
2.2.4. The Influence Factors on Heat Transfer
............................................... in Circulating Fluidized Beds 33
2.2.5. Heat Transfer Modelling ........................................................ -43
........................ 2.3. Gas and Solids Mixing in Circulating Fluidized Beds .. 46
. CHAPTER 3 EXPERIMENTU EQUIPMENT ...................................................... 49
3.1. The Circulating Fluidized Bed Unit ....................................................... 49
................................................................... 3.2. Distributors for the Downer 5 1
.................................................................... 3.3. Fluidized Air and Particles 55
3.4. Heat Transfer Measurements ............................................................. 56
3.4.1 . Measurements for Heat Trans fer between Flow
...................................... Suspension and Suspended Surface 56
............. 3.4.2. The Measurement for Gas-solids Contact Efficiency 62
3.5. Measurement Techniques for Gas and Solids Flow .............................. 66
.............. 3.5.1. Axial Pressure Gradient and Pressure Measurement 66
............... 3.5.2. Particle Concentration and Velocity Measurements 66
CHAPTER 4 . HEAT TRANSFER BETWEEN SUSPENSION FLOW AND
SUSPENDED SURFACES IN THE DOWNIER ............................. 68
4.1. Axial and radial distributions of the heat transfer coefficient ............... 73
........................ 4.2. The Effect of Solids Circulating Rate on Heat Trans fer 76
...................... 4.3. The Effect of Superficial Gas Velocity on Heat Transfer 76
....... 4.4. The Relationship between Suspension Density and Heat Transfer 78
............. 4.5. The relationship between Particle Velocity and Heat Transfer 79
vii
.................................. 4.6. Cornparison with the Heat Transfer in the Riser 79
CHAPTER 5 . CHARACTERIZING GAS AM) SOLIDS DISTRIBUTORS
WTH HEAT TRANSFER STLJDY IN A GAS-SOLIDS
.................................................................. D O W R REACTOR 103
5.1. The Relationship between Heat Transfer and Solids Concentration ..... 109
5 .2 . Radial Distributions of Heat Transfer Coefficient
with Distributors A and B .................. .... ....................................... 110
5.3. Radial Distributions of Heat Transfer Coefficient
......................................................... with Distributors 4 C, D and E 112
.......... 5.4. Effect of Distributor Structure on Gas-solid Flow Developrnent 113
CHAPTER 6 . GAS-SOLIDS CONTACT EFFICIENCY IN THE
ENIRANCE REGION OF A CO-CURRENT DOWNFLOW
...................................................................................... FLmIZED BED 133
.........,................................ . 6.1. Gas Temperature Distribution Profiles .... 140
6.2. The Axial Distribution of Average Gas-solids Contact Efficiency ....... 141
6.3. Effect of Operating Condition on initial Gas-sotid Mixing .................. 144
6.4. Relationship between Local Solids Holdup and the Gas-solid
Contact Efficiency ......................................................................... 144
............................ CHAPTER 7 . CONCLUSIONS AND RECOMMENDATIONS 160
.................................................................................................. REFERENCES 164
....................................................................................................... APPENDIX 175
............................................................................................................. VITA,.. -192
LIST OF TABLES
Table Description
Table 3-1. The operating conditions in heat transfer measurements
Table 4-1. Comparison of the measured heat transfer coefficients and
the predicted values for forced air convection
Table 4-2. Comparison of the heat transfer coefficients in the domer
and in the riser
Table 5- 1. Operating conditions utilized in this study
Page
56
86
LIST OF FIGURES
Figure
Figure 2- 1.
Figure 2-2.
Figure 2-3.
Figure 2-4.
Figure 2-5.
Figure 2-6.
Figure 2-7.
Figure 2-8.
Figure 2-9.
Figure 2- 10.
Figure 2- 1 1.
Page Description
Axial and radial flow structure of gas-solids suspensions
in upflow circulating fluidized beds reported by several
researc hers (Horio, 1997) 9
Typical axial flow sections in the downer (Zhu et al., 1995) 11
Radial so1ids concentration distributions in a 140 mm i.d.
downer (Zhu et al., 1995) 13
Radial distribution of particle veiocity in a 140 mm i.d.
downer (Zhu et al.. 1995) 14
Radial distribution of gas velocity in a 140 mm i.d. downer
(Zhu et al., 1 995) 15
Radial solids distribution in the entrance of a 140 mm i.d. downer
(Wei et al., 1997)
A typical multi-tube downer distributor design used by Tsinghua
University (Zhu et al., 1 995)
Downer distributor used by institute of French Petroleum
(Herbert et al.. 1994)
Mechanism of particle convection
Radial and axial distribution of heat transfer coefficient
in a riser of Silicagel A particles (Bi et al.. 1989)
Cornparison of heat transfer mechanisms at different
temperatures and solids concentrations (Gliskman et a l , 1997).
Figure 2-12.
Figure 2- 1 3.
Figure 2-14.
Figure 2- 1 5.
Figure 2- 16.
Figure 2-1 7.
The relationship between the heat transfer coeficient and
suspension density reported b y di fferent researc h groups
(Basu and Nag, 1996) 35
The relationship between the heat trans fer coefficient and
suspension density (Basu and Nag, 1996) 36
The relationship between the heat transfer coefficient and bed
temperature (Wu et al., 1989a) 37
The relationship between the heat transfer coefficient and
particle size (Wu et al.. 1989a) 39
Schematic of the membrane heat tram fer surface 31
Heat transfer coefficient measured in the different region
of the wall (Lockhart et al.. 1 995)
Schematic of the riseddowner circulating fluidized bed
Schematic of the distributor designs (Johnson et al., 19%)
Schematic of the distributor designs
Schematic of a long heat transfer probe (Bi el al., 1991 )
Schematic of an instantaneous heat transfer probe
and measurement circuit (Wu et al., 1989)
Schematic of the miniature cylindrical heat transfer probe
used in this study
The measurement for gas-solids contact efficiency in
a 0.09 m i.d. riser (Dry et al, 1990)
Figure 3- 1.
Figure 3-2.
Figure 3-3.
Figure 3-4.
Figure 3-5.
Figure 3-6.
Figure 3-7.
Figure 3-8. Schernatic of the probe for the gas-solids contact efficiency
used in this study
Schematic of the riser/downer circulating fluidized bed
Schematic of the miniature heat transfer probe
Axial distribution of the average heat transfer coefficient
Axial distribution of the average suspension density
Axial distribution of the pressure gradient
Axial distribution of the local heat transfer coefficient
at different radiai positions
Radial distribution of the heat transfer coefficient
along the downer
The effect of solids circulating rate on the heat transfer
The effect of solids circulating rate on the radial distribution
of the heat transfer coefficient along the downer
The effect of superficial gas velocity on the heat transfer
coefficient 98
The effect of gas velocity on the radial distribution
of the heat transfer coefficient dong the downer 99
The radial distribution of the solids holdup along the bed 1 O0
Cornparison of the local heat transfer coefficient
with the local solid holdup IO1
The relationship between the local heat transfer coefficients
and local suspension densities the corresponding
Figure 4- 1.
Figure 4-2.
Figure 4-3.
Figure 4-4.
Figure 4-5.
Figure 4-6.
Figure 4-7.
Figure 4-8.
Figure 4-9.
Figure 4-10.
Figure 4- 1 1.
Figure 4-1 2.
Figure 4- 1 3.
Figure 4- 14.
xii
Figure 5-1.
Figure 5-2.
Figure 5-3.
Figure 5-4.
Figure 5-5.
Figure 5-6.
Figure 5-7.
Figure 5-8.
Figure 5-9.
Figure 6- 1.
Figure 6-2.
Figure 6-3.
Schematic of the riseddowner circulating fiuidized bed
Schematic of the distributor design and the probe
for heat transfer
The relationship between heat transfer coefficient and
local solids concentration
(a-Wu et al., 1989; b- Lockhart et al., 1995)
The relationship between the heat transfer coefficient and the
local solids concentration with Distributors A and B 126
Radial distributions of heat transfer coefficient with
Distributor A
Radial distributions of heat transfer coefficient with
Distributor B
Radial distributions of heat transfer coefficient with
Distributors A, C, D and E 130
Axial distributions of the average heat transfer coefficients with
Distributor A and B 131
Axial profiles of RN1 (h) for different distributor types
Schematic of the riseddowner circulating fluidized bed
Schematic of the distributor design and the gas-solids
mixing rneasurement technique
Radial distribution measured gas temperature
with distributor #1
Figure 6-4. Axial distribution of average gas temperature with
distributors #1 and #3
Figure 6-5. Axial distributions of the gas-solids contact efficiency
with distributor # 1 , #2 and #3
Figure 6-6. Effect of operating conditions on the gas-solids
contact efficiency with distributors #1 and #3
Figure 6-7. The relationship between the gas-solids contact
efficiency and solids holdup
xiv
NOTATION
surface to volume ratio of particles, lm
heat transfer surface area, m2
Archimedes number, (= dip, (p, - p, )g / p,' )
solids concentration, kg/m3
cluster concentration, kJ/m'
specific heat of packet, k.J/m3
specific heat of gas, k.J/m3
specific heat of particle, kl/m3
particle size, pm
diameter of fluidized bed, m
contact efficiency between gas and solids
the cluster-to-wall view factor
acceleration due to gavity, m/s2
solids circulating rate, kglm's
heat transfer coefficient, w/m2~
average heat transfer coefficient, W I ~ ~ K
heat transfer coefficient of "packet", w / ~ * K
heat transfer coefficient due to gas film, w 1 m 2 ~
gas convective heat transfer coefficient, w / ~ ' K
heat transfer coefficient between gas and solids, W I ~ ' K
highest possible heat transfer coefficient observed
in the downer, w / ~ ' K
minimum (gas only) heat transfer coefficient in the downer,
w / ~ ' K
particle convective heat transfer coefficient, w / ~ ' K
radiative heat transfer coefficient, w / ~ ' K
axial distance from the distributor, m
electrk current through the probe, rnA
empincal constant as defined in eqn. (2-5)
thermal conductivity of cluster, W/mK
thermal conductivity of packet, W/mK
thermal conductivity of gas, W/mK
length over which particle remains in contact
with wall, m
the length of the probe, m
Nusselt number (= hP&)
pressure, kPa
pressure drop across furnace, Pa
Prandtl number (= cm pk%)
heat flow, w/m2
xvi
heat flow of radiation, w/m2
heat flux of gas, w/mL
heat flux of gas-solids, w/m2
radial position
column radius, m
reduced radia1 position
Reynolds nurnber based on slip velocity (dppg( Vg- Y,)/&)
Radial Nonuniformity Index of heat trans fer
Radial Nonuniformity Index of solids holdup
temperature, OC
bed temperature, O C
cluster temperature, O C
initial gas temperature, O C
gas temperature, OC
average gas temperature in a given bed section, OC
gas temperature at the top of a given bed section, OC
gas temperature at the bottom of a given bed section, O C
average solid temperature in a given bed section, O C
initial solid temperature, O C
ratio of the increase in temperame
surface temperature of the probe. OC
xvii
Greek Letters
4 P
4
E
wall temperature, O C
time of particle contact on the surface, s
superficial gas velocity, m/s
voltage applied on probe, V
gas velocity, m/s
nozzle gas velocity, m/s
solids velocity, m/s
terminal velocty of cluster on wall, m/s
the location of the cluster in heat transfer surface
height of a given bed section
empirical constants used in eqn. (2-5)
volume fraction of heat transfer exposed to clusten
void fiaction
void fiaction of cluster
cluster effective emissivity
void fraction of 'packet"
wall emissivity
density of bed, kg/m3
density of cluster, kg/m3
density of the up-flowing gas containing dispersed solids
cluster, kg/m3
xviii
density of packet, kg/rn3
densiîy of gas, kg/m3
partic le density, kg/m3
suspension density, kg/m3
gas visconsity, Pa s
Stefan-Boltzamnn constant, 5 . 6 7 ~ 1 o - ~ w/rn'~'
standard deviation of the radial heat transfer coefficient
highest possible standard deviation of heat transfer coefficient
CHAPTER 1. INTRODUCTION
Circulating fluidized bed (CFB) technology h a been widely used for various gas-
solid reactions such as catalytic cracking, combustion and other reactions which
commonly require heat transfer during the reactions. A clear understanding of heat
transfer behaviors in CFB will help to control the bed temperature and energy exchange
in the reactor and is thus essential for the proper design of some CFB reactors. During the
design of fluidized bed units, it is often necessary to know the amount of heat transfer to
the walls of the equipment, to the surfaces irnmersed in the bed, and between gas and
solids phases. Under the proper operating conditions and with correct design based on
reliable knowledge of the mechanism of heat transfer, a circulating fluidized bed c m be a
very satisfactory thermal system with the best use of energy sources. Therefore, heat
transfer studies in CO-current upflow circulating fluidized beds (risers) have increased
significantly in the last two decades. Many researchers have presented comprehensive
reviews on this subject (Grace, 1986; 1990a; Glicksman, 1988; 1997; Leckner, 1991 a;
Basu and Nag, 1 996).
With the development of CFB, CO-current downflow circulating fluidized bed
(downer) has also been proposed more recently as an alternative to the riser. It has
c&kly attracted the attention b m many researchers in different areas. Because the gas
and solids flow directions are downwards in the sarne direction as gavity, downer
reactors have been shown to have many district advantages over riser reactors (Zhu et al.,
1995): it has much shorter gas-solid contact time, more uniform axial and radial flow
structure and more uniform gas and solids residence times. Thus, downers have the
ability to bring gas and solids into contact in a fairly uniform manner, making the flow
pattern significantly closer to plug flow than risers. The short and uniform contact time
for both gas and solid phases in the downer leads to a better reaction selectivity and more
uniform product distribution. This feature is likely to lead the downer reactor to new
applications with gas-solids operation processes such as Fluidized Catalytic Cracking
(FCC) where short contact tirne and uniform gas and solids residence time distribution
are extremely important and where extremely high temperature regions (hot spots) should
be avoided. Although a lot of fundamental and applied research has been carried out (Zhu
et al-. 1995; Zhu and Wei, 1996; Aubert et al., 1994; Herbert et ai., 1994), most of the
previous research in the downer hydrodynamics were carried out mainly in the fully
developed section and no result has been reported on the heat transfer and the gas-solids
mixing in the downers. The characteristics of heat transfer and gas-solids mixing are very
important elements in the downer reactor design and development, so that it warrants
careful study.
The objective of this research is to study the heat transfer and gas-solids mixing
behaviors in CO-current downflow circulating fluidized bed. It consists of three parts:
(1) the heat transfer between the suspended surface and the gas-particle flow suspension;
(2) the contact efficiency between the gas and solids phases in the entrance region;
(3) the effect of gas and solids distributor configuration on the heat transfer behavior and
flow development in the downer.
In the first part, the local heat transfer coefficients between the gas-solids
suspension flow and a srnall suspended surface were measured at different operating
conditions in various radial and axial bed locations. The radial and axial distributions of
heat transfer coefficient and the effects of operating conditions on the heat bansfer
coefficient were investigated. Analyses of local heat transfer coefficient demonstrate the
signifiant influence of the solids circulating rate and the gas velocity on the heat transfer
and the close relationship between the heat transfer behavior and the solid suspension
density in the downer. To compare the heat transfer in the downer with that in the riser,
additional expenments on heat transfer coefficient between the gas-solids flow and the
suspended surface in the accompanying riser were also canied out under similar
operating conditions. The axial and radial distribution profiles of the heat transfer
coefficients in the downer and the riser were found to be significantly different. Sorne
factors affecting heat transfer were also found to be different.
In the second part, the effect of distributor configuration on heat transfer between
the gas-solids Bow suspension and the suspended surface in the entrance region of the
downer were examined under five different types of distributors. Through measuring the
axial and radial distributions of the heat transfer coefficients, and due to the close
relationship between the heat transfer behavior and the gas-solids Bow pattern in the
downer, those five distributors are characterized in terms of their influence to the flow
conditions in the entrance region of the downer.
In the third part, the gas-solids interphase heat transfer study was used to estimate
the contact efficiency between the gas and solids phases in the entrance region of the
downer, since the gas-solids contact efficiency is mainly dependent on the
hydrodynamics and heat transfer behaviors in the downer. Ln other words. a thermal
rnethod was utilized to investigate the gas-solids mixing phenornena in the downer. The
gas-solids contact efficiency was obtained through measuing the temperature change of
hot fluidized air in the bed. The axial distributions of the contact efficiency were then
obtained at different operating conditions. Three different types of gas and solids
distributors were employed to investigate the effects of distributor design on the gas-
solids mixing behavior. Both the operating conditions and the disaibutor design were
found to have obvious effects on the gas-solids rnixing behavior. The appropriate
distributor structure can also improve the gas and solids contact effectively.
Al1 of this infornation significantly impacts the design and operation of downers.
CHAPTER 2. LITERATURE REVIE W
2.1. Circulating Fluidized Beds
2.1.1. Fundamentals and Applications of Co-Current Downflow Circulating
Fluidized Beds
Circulating fluidized bed (CFB) riser is a type of gas-solid reactor which has been
widely applied and developed in the chernical industry due to many intrinsic properties.
Its applications include fluid catalytic cracking, polyethylene production, calcination
operating and combustion etc. Compared with conventionai bubbling and turbulent
fluidized beds, circulating fluidized beds (risers) have such advantages as high gas-solids
contact efficiency, high solids throughput, reduced axial dispersion of both gas and solids
phases, high tumdown ratios, and the ability to handle wide range of different particles.
On the other hand, the concurrent upflow circulating fluidized beds (riser), in which the
gas-soiid suspension is transported upward, stiIl have some disadvantages, including
relatively severe solids back-mixing and non-uniform gas and solids flow. The radial
segregation of gas and solids Ieads not only to reduced contacting between the two
phases, but also to less uniform distribution of the desired product and a reduced
selectivity. In fluid catalytic cracking (FCC), for exarnple, the radial non-uniform gas-
solids contact and solids dispersion in the wall region c m cause severe over-cracking.
Due to these shortcomings of the riser reactor, a concurrent downflow circulating
fluidized bed (downer) was proposed as an alternative, where gas and solids flow
directions are downwards in the same direction of gravity. It has been found that downer
reactors have several distinct advantages: short gas-solid contact time, more uniform gas-
solids contact, more uniform radial distribution of the gas and solids flow and
significantly reduced axial gas and solids dispersion. These advantages should result in:
(i) more efficient gas-soli& contact; (ii) more uniform gas-solids contact time; (iii) hi&
temperature regions (hot spots) being avoided; (iv) better product selectivity; and (v)
higher operating temperature closer to the maximum design temperature without the risk
of localized over-heating.
The first application of downer reactor appeared to be the plasma ultrapyrolysis of
coal in the 1960s and 1970s in the former USSR and Germany (Beiers et al., 1988;
Brachold et al., 1993; lin, 1994). From the 1970s, Stone and Webster Engineering
Corporation began to develop a new type of reactor referred to as the "Quick Contact"
(QC) reactor (Gartside, 1989). The QC reactor is reported to offer very short residence
times (- 200~1, near plug flow and a high temperature reaction environment. Murphy
(1992) proposed an FCCheavy oil cracker unit which incorporates a downflow reactor
and a nser regenerator. Mobil and Texaco have both patented downer reacton for the
FCC process (Gross and Ramage, 1983; Gross, 1983; Niccum and Bunn, 1985). They
claim uniform distribution of catalyst, decreased contact time of cataiyst with the feed
and reduced coking. Berg et al. (1989) proposed a downflow Ultra-Rapid Fluidized
m) reactor, which has now been successfÙlly applied to biomass pyrolysis (Graham es ai., 1991). To respond to the potential industrial applications of the downer, fundamental
studies in downer were started by Shimizu et al. (1978) and Kim and Seader (1983).
From the 1980s, researchen at Tsinghua University have camied out a senes of downer
hydrodynamic studies (Bai et al., 1991% b; 1 992a, b; Yang et al., 199 la; Wei et al.,
1994; 1995; Zhu and Wei, 1996). The researchers at the French Institute of Petroleum
also published their hydrodynamic results from a 50 mm diameter downer (Aubert et al.,
1994; Herbert et al., 1994; Herbert, 1997). More recently, cornprehensive hydrodynarnics
studies have also been carried out at this univenity (The University of Western Ontario)
to achieve better understanding of the downer (Johnston et al., 1998a; 1 998b). However,
there are still many unanswered questions such as the initial gas-solids mixing and flow
development in the downer, the effect of the gas and solids distributor design and heat
transfer between the gas-solids flow suspension and the suspended surfaces
2.1.2. Hydrodynamics in Upflow Circulating Fluidized Beds
The hydrodynamics of upflow circulating fluidized beds (nsers) have been studied
for decades. The major characteristics of the riser reactor have been presented very clear:
Ln general, a riser reactor may operate either in the fast fluidization regime or in the
pneumatic transport regime. When the gas velocity is increased, the bed regime changes
fiom a packed bed to a bubbling bed, to slug flow, to the turbulent regime. to fast
fluidization and eventually to pneumatic transport. So for its axial gas-solids flow
structure, a riser reactor usudly can be divided into a dense bottom region and a dilute
upper region. The bonom region generally operates either in bubbling or turbulent
fluidization mode depending on the superficial gas velocity used. in the upper dilute zone
the solids volumetric concentration becomes very low and solids density aimost remains
constant. The overall radial flow structure of the nser bed is better explained by the core-
annulus mode1 which consists of two vertical zones: a relatively dilute up-flowing core in
which solid particles are entrained upward by high-velocity gas Stream; and a much
denser annular layer near the column wall in which solid particles congregate and either
rise slowly or fa11 d o m as dense structures sirnilar to waves of strands or streamers
(Rhodes, 1990; Harris and Davidson, 1994; Grace, 1 WOb).
The axial and radial flow structure have been confirmed by many experimental
measurements (Herb et a1.,1989; Knowlton, 1995; Bader et al., 1988). Figure 2-1 clearly
shows that the solids concentration is always much higher in the bottom region than in
the upper region and decreases with the bed height. The radial suspension density
gradients are reported with a maximum at the wall and a minimum at the centre, the
solids concentration is seen to have a dramatic increase in the annulus region, which
agrees well with the core-annula approximation to the riser flow structure. Also, the
radial distribution profiles of the particle velocity and solids flux across the cross-section
of risen are found to be approximately parabolic (Figure 2 4 , often with negative values
along the riser wall. This indicates that the particles near the wall have the tendency of
flowing downward. Furthemore, the axial and radial flow structure, including the height
of the bottom dense region and the radial solids concentration and velocity distributions
c m vary with the colurnn design and operating conditions.
From the hydrodynamics shidies, it is clearly shown that radial variations of
solids concentration fiom the core to the wall is closely linked to the radial distributions
of mass and heat transfer behavior. The concentration and velocity of the descending
particles and their residence time at the wall are important parameten that will affect the
heat transfer between the gas-solids suspension and the heat transfer surface. Obviously,
the non-uniformity of the radial and axial gas and solids flow structure have caused the
non-uniform heat tram fer distribution in risers (Horio et al., 1988; Gliskrnan, 1997).
2.1.3. Hydrodynamics in Downflow Circulating Fluidized Beds
(1) Axial Gus and Solids Flow Structure
In a downer reactor, gas and particles are fed fiom the top of the downer through
gas and particle distributors. Solids acceleration is caused by both gravity and drag. There
typically exist three distinct flow sections along the axis of downer reactors (Figure 2-2).
The first one is the first acceleration section, which is from the top to the position where
particle velocity is equal to the gas velocity. In this section, particles are accelerated by
both gravity and gas flow so that the pressure gradient is negative and the absolute
pressure, P, decreases dong the downer. In the second acceleration section, solids are
further accelerated by gravity, and particle velocity increases m e r until the slip velocity
Figure 2- 1. Axial and radial flow structure of gas-so!idr suspensions in upflow circulating fluidized beds reported by several researchers
(Horio, 1997)
between the particle and gas reaches a value where the drag force counter-balances the
gravitational force. Hence the pressure gradient is positive and the absolute pressure
increases along the downer. In the third section, the gravitational force is in balance with
the drag, both particle and gas velocities remain constant downstream. It is named the
constant velocity section, in which pressure gradient remains constant and the absolute
pressure increases linearly along the downer. For this three-section the axial flow
structure has been confinned by pressure and pressure gradient measurements in the
downer (Zhu et al., 1995; Iohnston et al., 1998b).
(2) Radial Gas and Solids Flow Structure
Radial distributions of solids concentration, solids velocity, solids flux. and gas
velocity are typically used to characterize the hydrodynarnics of the flow in the downer.
Compared with the radial gas and solid flow structure in risers which is characterized by
a core-annulus flow structure, the radial distribution of gas and solids flow in downers are
much more uniform over the column cross section.
Researchers fiorn Tsinghua University have presented typical radial profiles of
gas and particle velocities and solids concentration in the downer of 140 mm i.d., 4.7 m
in height with 54 prn FCC particles (Bai et al., 1991a; Cao et al., 1994; Wang et aL.
1992):
(i) The solids concentration is seen to remain relatively constant in the centre region untii
r/R reaches 0.8, where a dense ring with significantly higher solids concentration begins
to develop. After reaching a maximum value at r/R = 0.85 - 0.9, the solids concentration drops towards the wall (Figure 2-3).
(ii) The solids velocity distribution profiles also have a peak near the wall. It reaches a
maximum in this dense ring region at r/R a 0.85 to 0.96, with much smaller peaks
compared to that in the solids concentration profile (Figure 2-4). In the downer, the local
2.0 4.0 6 .O 0.002 0.004 0.006
V,, m/s 1 -€
Figure 2-2. Typical axial flow sections in the downer (Zhu et al.. 1995)
solids velocity can be higher than the local gas velocity, which can not exist in the nser.
(iii) The radial distribution profiles of gas velocity, s h o w in Figure 2-5, are rather flat.
and do not change significantly over the bed cross-section. From the profiles, it can be
seen that the particle velocity also attains a maximum at r/R = 0.85 to 0.96, corresponding
to the maximum local solids concentration discussed above. With increasing superficial
gas velocity, the shape of the radial profile of gas velocity does not vaiy obviously. but
the local gas velocity over the cross section increases evenly.
On the other hand, the experiments of Herbert (1997) taken in a downer of 50 mm
interna1 diameter and 5.0 m hi&, indicated that the distribution profiles of solids
concentration are rather flat in the centrai core region and but increase to a maximum
value around r/R = 0.6 rather than for an r/R = 0.8 to 0.95 as measured by Bai et al.
(1991b).
The solids velocity profiles reported are also more uniforrn in the downer than
these in the nser. However, the profile of solids velocity has a parabolic shape with
higher values in the central core region and lower values towards the wall than that of the
Tsinghua results. The different downer reactor size and the column material may be the
reason for these differences.
Results reported by Tsinghua researchers (Bai et al., 199 1 a; b; Yang er ni., 199 1 a;
b) show that three different regions rnay exist across the downer cross section: a dilute
core fiom r/R = 0.0 to about 0.85, where the local solids concentration, particle velocity,
and solids flux are rather uniform; a denser annular region fiom about r/R = 0.85 to 0.96,
where al1 three variables have a maximum value; and a wall region fiom r/R 2 0.96,
where al1 three variables decrease towards the wall.
Figure 2-3. Radial solids concentration distributions in a 140 mm i.d. downer (Zhu et al., 1995)
Figure 2-4. Radial distribution of particle velocity in a 140 mm i.d. downer (Zhu et a l , 1995)
Figure 2-5. Radial distribution of gas velocity in a 140 mm i.d. downcr (Zhu et al., 1995)
This three section radial flow structure is reasonable due to the flow
characteristics of the gas-solids suspension in downers as thoroughly explained by Zhu et
al. (1995). It can be considered that particles initially distributed uniformly by the top
distributor and then falling at the same velocity in the entrance region. In order to satisQ
the no-slip condition at the wall, the gas velocity in the wall must become lower, leading
to lower particle velocity in the same region. Following, since lower particle velocity
must correspond to a lower solids concentration in the downer, the solids concentration
will be lower in the wall region. Particle velocity and concentration consequently reduce
in the wall region and cause the migration of particles from the wall to the annular region,
leading to increase of particle concentration and velocity in the annular region. Because
of the influx of solids to the annular region at the top of the downer, the concentration and
velocity is higher than in the central core region. Furthemore, the more dense annular -
more dilute core and wall flow structure is a stable flow structure.
2.1.4. Gas and Solids Flow Structure at the Top Entrance Zone in the Downer
The gas and solids flow structure at the top entrance region in the downer is one
of the most important parts in hydrodynamic studies of the downer. Chen er al., (1992)
studied the influence of entrance structure on the axial pressure profile in the downer.
Wei et al. (1997) studied the radial solids fraction profile and the solids residence time
distribution in the entrance region of a 140 mm i.d. downer under hvo types of solids
distributors. The results indicated that two regions may exist in the entrance region: the
disûibutor effect region and the turbulence control region. in the distributor effect
region, as shown in Figure 2-6, solids are non-uniformly distributed across the radial
direction near the disûibutor. It was reported there exists a dense region near the solids
distribution tubes. As solids leave the distributor effect region and radial solids dispersion
occurs, the dense region becornes dilute and the profile develops into a flatter shape,
marking the end of the distributor effect region. The turbulence control region is
characterized by the development of a dense ring at r/R = 0.8 to 0.95 as solids aggregate
towards the wall region, which must be comected with the gas-solids turbulence flow.
They found that the distributor structures have a large influence on the initial profiles of
the radial solids fraction. The distributors were reported also to have an effect on the
Iength of the distributor effect region. But both the superficial gas velocity and the solids
circulation rate seern not to have significant effect on the lengths of the entrance regions
and the characteristic shapes of the radial profiles.
These results fiom Wei et al. (1 997) provided some of the characteristics of the initial gas
and solids flow structure in the entrance region, but more comprehensive study is still
needed to fully understand the effect of the distributor design and operation conditions on
the flow structure and development and the gas-solids mixing behaviors in the entrance
region.
2.1 .S. Downer Distributor Configuration
In the downer, a unifom distribution of solids at the entrance is more important
than for risers, since solids acceleration in a downer is caused by both gravity and gas
drag; whereas in risers, solids acceleration relies entirely on gas drag, so that a uniform
distribution of gas become more important in nsers (Zhu et ai., 1995). For this reason,
the top distributor for the downer should be primarily designed to evenly distribute solids
across the downer cross-section while the corresponding bottom disaibutor in nsers is
designed mainly for effective gas distribution over the cross-section.
Figure 2-6. Radial solids distribution in the entrance of a 140 mm i.d. downer (Wei et al., 1997)
A typical solids feeding system is illustrated schematically in Figure 2-7, which is
employed in Tsinghua University (Zhu et al., 1995). The solids feed is fiom a fluidized
bed, which is situated at the top of the downer. It consists of many small diameter vertical
solids-delivery tubes to evenly deliver soli& into the downer. The main fluidization gas
is introduced below the top distributor bed, and gas deflecting devices may be installed to
direct the gas flow downwards. The distributor bed is either semi-fluidized or kept around
minimum fluidization to allow unifonn solids feed (Tesch et al., 1994). Bubbles in this
bed should be avoided to prevent fluctuations in the solids feed. Small orifices may be
dnlled on the wall of the distributor tubes within the bed to improve solids distribution.
The solids flowrate may be adjusted by the bed height andior by the flowrate of
distributor fluidization air.
Herbert et al. (1994) and Aubert et al. (1994) used a distributor similar to that
shown in Figure 2-8. The downer portion of the reactor continues up into a fluidized bed
feeding area. The solids flow downward the column when the bed height surpasses the
height of the downer. The air is simply injected through a narrow slanted ring around the
entire downer colurnn. For the studies in the entrance region, this design seems not to be
very optimal because the solids could not be easily distributed over the column cross-
sectional area but just along the downer wall. The gas injection will tend to impinge
around the solids flow to cause an even redistribution of the solids.
Al1 of the radial distribution profiles of gas velocity and solids concentration and
velocity reported in the downer review paper by Zhu et al. (1995) have been measured in
the sarne cold downer mode1 apparatus with one type of distributor. It consisted of seven
vertical tubes distributing solids from a semi-fluidized bed into the downer top; in the
meantirne, the gas was fed fkom a perforated plate with a large open area. Solids and gas
are evenly distributed in an equilateral pitch over the column cross-section. However,
little radial gas-solids mixing is promoted because the design of disû-ibutor is to form
Solids inlet
Downer
Figure 2-7. A typical multi-tube downer distributor design used by Tsinghua University (Zhu et al., 1995)
Figure 2-8. Downer distributor used by the Institute of French Petroleum (Herbert et al., 1994)
stable jets in the axial direction and does not provide good radial mixing. Most of the
papers published (Bai et al., 199 1 a; 1992a; Cao et al., 1994; Qi et al., 1990; Wang et al.,
1992; Wei et al., 1994; 1995; Yang et al., 199 1 b) reported the hydrodynarnics of flow in
the fully developed region in which the distributors have little effect on the gas-solids
flow structure. For this reason, these studies are not directly relevant to the distributor
design.
Gas and solids distributors will obviously influence the flow hydrodynamic near
the distributor, in the acceleration and flow development regions. A good distributor
would provide excellent gas-solid mixing and uniform distributions of gas and solids over
the column and enhance the flow development. Furthemore, in order to optimize the
downer reactor for industrial application, the initial gas and solids flow structure in the
entrance section is very important for the downer designs. For this reason, development
and improvement of gas and solids distributor is very essential in the downer research
area.
2.2. Heat Transfer in Upflow Circulating Fluidized Beds
Many circulating fluidized beds involving combustion or other exotherrnic
reactions cotnrnonly require heat exchange during the reaction. With CFB riser becoming
more and more popular, especially in the last two decades(Grace et al., 1997), an accurate
understanding of heat transfer in circulating fluidized beds is very important for the
proper design of CFB reactors. Many studies have been carried out to test the effect of
different design and operating parameters on heat transfer (Grace, 1990a; Glicksman,
1988; Basu 1990)
Even though no results have been reported so far on the heat transfer in the
downer, the mechanism and influencing factors of the heat transfer in downers are
expected to be similar to those in risers. So a review of the heat transfer in risers can help
to better understand the heat transfer behavior in downers which is to be studied in this
work. Previous research indicate that the heat transfer rate and its mechanism in the
circ ulating fluidized bed are largely govemed b y the hydrodynamic conditions and are
directly related to the gas-solids flow patterns, especially to the particle and gas behavior
near the heat transfer surface in the bed (Grace, 1990a; Glicksman, 1988; Basu and Nag,
1996; Leclmer, 199 I a).
The heat transfer in a CFB generally consists of four most important parts: ( 1 )
heat transfer between bed (gas-solids mixture) and column or other suspended surfaces;
(2) heat transfer between gas and the particle surface; (3) heat transfer between the
particle surface and its core (4) heat transfer between particles. Because the temperature
of particles can be considered uniform due to the severe solids-mixing, the heat transfer
resistant between particles can be reasonably neglected. And because particles used in
CFB are usually very small and particle conductivity is much larger than that of gas, the
temperature gradients inside individual particles can also be neglected. Therefore the
most important heat transfer processes in a CFB are the heat transfer between the gas-
solids suspension flow and heat exchange surfaces; and the heat transfer between gas and
solids phases. The latter has been scantily reported because in the nser reactor the heat
transfer between the gas and solids just exists in the very short section in the entrance
region of the riser and the two phases reach thermal equilibrium very quickly (Grace,
1990a). Therefore, only heat transfer behavior between gas-solids flow suspension and
heat transfer surfaces is reviewed here.
2.2.1. Particle Convection
In a circulating fluidized bed, heat c m be transferred fiom the gas-solids flow of
the bec3 to heat transfer surface by several different mechanisms. The overall heat transfer
behavior c m include (1) particle convection (2) gas convection and (3) radiation. In most
CFB risers, particie convection is the primary heat transfer mechanism, which has been
confirmed by Glicksman et al. (1 993) and Ebert et al. (1 993).
When the heated particles at the bed temperature move to the heat transfer
surface, the heat exchange takes place between the particles and the contacting surface.
Since in general the particles seldom touch the surface directly, most of the heat transfer
would take place through a gas layer separating the particle and the surface. The overall
process is properly termed particle convection (Glicksman, 1997). The particle
convection is the dominating heat transfer process at the heat transfer surface, normally
covered by comparatively dense particles, clusters, swanns or strands. In particle
convection, the motion of the particles fkom main suspension flow to the surface is the
key mechanism for energy transfer. Therefore, particle convective heat transfer is mostly
controlled by the particle concentration and residence time of the particles on the heat
transfer surface. The schematics of particle convection cm be illustrated as shown in
Figure 2-9: Particles (a cluster) first appears at the wall at location XI, they fa11 with the
flow along the surface and in the meantirne, transfer the heat to the heat exchange surface.
M e r falling some distance L I , they separate fiom the surface and mix with the core flow
again. The average heat transfer coefficient is made up of the contributions of the whole
array of clusters which arrive at diflerent positions and fa11 different distances before
departue.
Figure 2-9. Mechanism of particle convection
The particle convection heat transfer rate and its mechanism in the circulating
fluidized bed have been widely studied. A detailed reviews of experimental data collected
kom both laboratory and commercial size circulating fluidized beds are presented by
Glicksman (1 988). Grace (1 990a) and Basu and Nag (1 996).
The earliest study on heat transfer in non-circulating fluidized beds was made by
Mickley and Trilling ( 1 949). They conducted numerous expenments to estimate heat
transfer coefficients at different superficial gas velocities and particle sizes. A strong
effect of particle size and suspension density was noted. They presented one of the very
useful models "Packet" mode1 (Mickley and Fairbanks, 1955), which explains that the
heat transfer could be represented as a transient process hom the "packet", a group of
particles, to the wall with the resulting heat flux related to the residence time of the
"packet" at the surface.
Kobro and Brereton ( 1986) measured heat transfer coefficients in a 3 rn long and
0.2 m diameter CFB combustor using a srna11 100 mm heat transfer probe. Heat transfer
coefficients were found to be 70 to 280 w/rn2~ at 25OC and 850'~. It was found that the heat transfer coefficients increase with the increasing suspension density. Basu and Nag
(1 987) found that for a given solids circulation rate, the heat transfer coefficient decreases
with an increase in the superficial gas velocity. This is still related to the solids
suspension density in the bed, because the solids suspension densities decrease with
increasing of gas velocity at a fixed solid circulating rate. Mahalingam and Kolar (1991)
measured heat transfer coefficients using a long cylindrical heat transfer probe of 4.64 m
long in a 100 mm square, 5.5 m ta11 CFB, operated at superficial velocities ranging from
4.2 to 8.2 d s and solid circulation rates fiorn 17 to 110 kg/m2s, with sand particles of
rnean diameters of 156,256 and 362.5 Pm. They also reported that the suspension density
had a strong influence on heat tramfer. For a fixed solid circulating rate, it was found that
heat transfer coefficients decreased with increasing superficial velocity.
Wu et al. (1987; 1989a) reported heat transfer data obtained in two senes of tests
at different bed temperature ranges, one was at 35OC and the other was 340 - 880°C, with sand particles. It was noted that at high temperature and low suspension density
conditions, radiation plays a significant role. The heat transfer coefficients increase
almost linearly with local suspension density ranged fi-orn O to 70 kg/m3 and the heat
transfer coefficient also Vary significantly with lateral position. Wu et al. (1989b. L991)
rneasured the instantaneous and time-averaged local bed-to-wall heat transfer coefficients
in a 9.3 m tall, 152 mm i.d. cold mode1 circulating fluidized bed nser, where the heat
transfer mechanisms and local hydrodynamics were studied by an instantaneous heat
transfer probe and a capacitance probe. It was found, in general, that the tirne variation in
the instantaneous heat transfer coefficient corresponds closely with that of the local
solids density. This suggests that the abrupt increases in heat transfer coefficient are
caused by the anival of particle packets or stands at the surface of the heat transfer probe.
It again confirmed that there exists a close relationship between heat trmsfer and
hydrodynamics in circulating fluidized beds and the important role of those particle
stands on heat transfer in circulating fluidized beds.
Mahalingarn and KoIar (1991) measured heat transfer coefficients in a 100 mm
square and 5.5 m ta11 CFB for different size particles. The strong influence of suspension
density on heat transfer had been noted again, but the direct influence of gas velocity was
not observed. It was found that the effect of superficial gas velocity seems not to be
separable from the dominant effect of the suspension density. Nag and Ali (1992) also
presented experimental results on the effects of operating parameters like bed
temperature, suspension density and superficial gas velocity on bed-to-wall heat transfer
in a hi&-temperature circulating fluidized bed and found similar results.
Bi et ai. (1989; 1991) presented an expenmental study of heat transfer in a fast
fluidized bed made of Plexiglas with 186 mm i.d. and 8 m in height, with silicagel A
particles. It was found that:
(1) The heat transfer coefficients decrease with the axial location, because the solids
density decrease with bed height.
(2) The heat transfer coefficient profiles are complex in fast fluidized bed. Generally, it is
uniform in the centre region and has a steep increase in the region near the bed wall,
which correspond to the solids concentration distribution.
(3) Solid concentration is the dominant factor influencing heat transfer; cornparatively,
the effect of gas and particle velocities are less significant.
(4) When solid concentration is high, the heat transfer coefficient profile is consistent
with that of the solid concentration; for low solid concentration, a minimum point of heat
transfer coefficient will appear in the reduced radial position from 0.5 to 0.8; and for very
low solids concentration and high gas velocity, gas convection becornes notable, so that
the heat transfer coefficient profile becomes more complex with higher values appearing
in the region near the axis rather than near the bed wall (Figure 2- 10).
From the previous research, it c m be concluded that the particle convection is the
most important heat transfer mechanism in CFB reactors and the dominating influence
factor is the solids suspension density in the bed.
2.2.2. Gas Convection
In the dilute region of a circulating fluidized bed, the suspended heat transfer
surface may only be contacted by gas or a very dilute particle-gas mixture most of the
time. In this case, gas motion becomes the primary rneans of transfemng energy from the
Us 3.7 mfa
0.0 QI 1.0
r lR
Figure 2- 10. Radiai and axial distribution of heat tram fer cocnicient in a riser of Silicagel A particles (Bi et ai., 1989)
bed to the heat exchange surface. This mechanisrn for heat transfer by gas motion to the
uncovered surface is termed gas convection. Gas convection is controlled by the gas
turbulent motion near the surface, so that high gas velocity can increase the contribution
of gas convection in the heat transfer process.
Relatively few studies have been carried out for the gas convection heat transfer
between the gas-solids flow and the uncovered portion of the heat transfer surface in
circulating fluidized beds due to the diffïculties in separating gas convection from particle
convection. Some investigators calculated the heat transfer coefficient of the gas
convection using the single phase correlation for the flow of gas alone (Wu ez al. 1989a;
Kunii and Levenspiel, 1991). Others implicitly assumed that the entire heat transfer
surface was covered by clusters and ignored the gas convection in the CFB bed (Dou,
1990, Sekthira et al., 1988, Mahalingarn and Kolar, 199 1). However, in very dilute flows
where only a small fraction of the heat transfer surface is covered by clusters at a given
time, gas convection must be very important.
Lints (1992) sweyed the CFB heat transfer measurements at cross-section
average densities below 50 kg/m3 and obtained an estimate for h, by linearly
extrapolating measurements to zero density. It is found that the zero density
extrapolations are larger than the correspondhg values obtained fkom existing single
phase gas flow correlations. This trend is in agreement with the results found by Ebert et
al. (1993) in a circulating fluidized bed using a mass transfer measurements. In this
expenment the overall mass transfer coefficient increased with the presence of the
particles but was insensitive to the particle concentration. These results suggest that the
presence of even a modest fraction of particle clusters tend to enhance the gas convection
on the measured surface. The low vaIues of the particle act as roughness elements or
turbulence promoters. They also found that the gzs convection heat transfer coefficient
varies from 20 to 10% of the total measured heat transfer for values of suspension density
between 12 and 79 kg/m5. However, fkom their expenment, it seemed that gas convective
component is not related directly to the superficial gas velocity.
Basu and Nag (1 987, 1990) argued that even in the very dilute region of a CFB, it
could not be entirely solids-fkee. A small number of particles are dispersed in this up-flow
gas. These particles have an important effect on the gas convection. The following
correlation of dust-laden gas (Wen and Miller, 1961) can be used to estimate the gas
convection component.
where pd, is the density of the up-flow gas containing dispersed solids, Y' is the terminal
velocity of those solids having an average diameter d,.
An empirical correlation for the gas convective component was also derived by
Botterill and Denloye (1 978) as:
This correlation also shows that the particle size has an effect on gas convection.
However this equation should be used with caution, because it was based on expex-imental
results with a fairly narrow particle size range.
2.2.3. Radiation
At high bed temperatures (> 400°C), the rate of heat transfer to the wall increases
due to the increase in gas conductivity and the contribution of radiation heat transfer.
Radiation acts in parallel with gas convection at the uncovered wall areas and improves
the heat transfer kom the gas-solids flow to the wall.
Assurning that (1) a cluster is initially in contact with the wall, (2) it has a uniform
temperature, (3) the cluster is considered as having an effective emissivity E,, and (4) the
wall and cluster act as gray bodies, the radiant transfer c m be written as (Basu, 1990)
If the clusters are isothermal, the effective emissivity will be higher than the particle
surface emissivity because of the reentrant geometry of the particle array. The radiation is
obviously dependent on the cluster temperature and the wall temperature.
Figure 2-1 1 shows a cornparison of heat tramfer mechanisms at different
temperatures and solids concentrations. At low volumetric solids concentration, the
contribution of the radiation and gas convection to the surface uncovered by clusten is
the largest mode of heat transfer. As the solids concentration increases, particle
convection becomes a more important process and its effect increases with increasing
solids concentration. While the heat transfer coefficient of particle convection increases,
both heat transfer coefficients of gas convection and radiation decrease, because the wall
shielding effect by cooler particles in the annulus become significant as the solids
concentration increases.
Only a limited amount of data on radiation heat transfer in CFB boilers is
available in the open literature due to experimental difficulties and proprietary
safeguards. Andersson and Leckner (1992) measured heat transfer coefficient on a 12Mth
CFB boiler using a heat flux probe, choral thermocouple and the overall heat balance
technique. Boyd and Freidrnan (1991) used the sarne technique to measure the heat
transfer in a 1 10 MWe CFB boiler.
Solid Volume Fraction
Figure 2- 1 1 . Cornparison of heat transfer mechanisms at di fferent temperatures and solids concentrations (Gliskman et al., 1997).
Couturier (1989) used a 25.4 mm horizontal cylinder to measure the local heat transfer
coefficient inside a 22 MWe CFB boiler. Werdermann and Werther (1994) observed fiom
their expenments in two different CFB boiler at Duisburg (226 MWth) and Bensburg
( 1 09 Mwth).
2.2.4. The Influence Factors on Heat Transfer in Circulating Fluidized Bed
From several previous studies in both laboratory and commercial units, it is
generally recognized that the suspension density of bed is the most important influencing
factor on the heat transfer in a CFB. The residence time of particles on the surface, which
is influenced by the particle velocities, is another major factor afEecting the heat transfer
coefficient (Basu and Nag, 1987). Ln addition, the operating conditions, particle
properties and geornetry of the bed also evidently affect the heat transfer in circulating
fluidized beds. Al1 of the above are related to the hydrodynamics of the solid and gas
mixture in the vicinity of the heat transfer surface.
(2) ïhe suspension densiry
From several previous studies, it is generally recognized that solid concentration
of the bed is the dominant factor influencing heat transfer in CFBs. Heat transfer
coefficients increase with suspension density. This is expected because the thermal
capacity of solids is much higher than that of a gas. Expenmental data gathered in
laboratory units by different investigators (Mickley and Trilling, 1949, Kiang et al., 1976,
Fraley et al., 1983, Kobro and Brereton, 1986) are plotted in Figures 2-12 and 2-1 3 in
order to compare the effect of suspension densities. It has been confirmed by many
researchers that heat transfer coefficients increase with the suspension density nearly
linearly. Glicksman (1988) suggested that heat transfer coefficients Vary as the square
root of the cross-sectional average suspension density. Using data from cold mode1 beds.
Divillo and Boyd (1 994) obtained the following fwictional relationship:
(21 Temperature
The heat transfer coefficient increases with bed temperature due to higher thermal
conductivity of gas and higher radiation at higher temperature (shown in Figure 2-14).
Above 400°C, heat transfer coefficients increase with temperature predominantly due to
radiation (Wu et al.. 1989b, Grace 1990a). In very dilute beds, radiation can become the
dominant heat transfer process. Iestin et al. (1 992) observed the effect of bed temperature
on heat transfer coefficients 60m their measurements in the 125MWe boiler at Carling,
and correlated their data in the following fom:
where dP is the pressure drop across the entire fumace (a rneasure of the suspension
density), Tb is the temperature of the furnace and k, a, ,O are empirical constants.
(3) Superficid gas velocity
In the riser reactors, most of the researchers believe that the superficial gas
velocity does not have great influence on the heat transfer coefficient if solids suspension
density is kept the same except in very dilute beds (Basu and Nag 1987; Wu et al., 1987;
Ebert et al., 1993). This is because the solids suspension density is relatively higher in
risers, resulting in a relatively low contribution of the gas convection component. So that
the effect of gas velocity on the heat transfer in risers seems not very significant.
WU ol (1987) wu et O 109 89) wu rt si (iYW S E K T H ~ d tniu CHI e l (ni81 suesreiro ami ruu taw MAC-AIID YCIRM (19901 UAniilINclur rno AaAa(rn1) ~ I n o A w Am muRtlY9ll UIHALlNGAM *)(b ItOCIAII99a
WirlRfîi (ma1
Figure 2- 12. The relationship between the heat transfer coefficient and the suspension density reporteci b y di fferent research groups
(Basu and Nag, 1996)
O K o k o rnd Brcrelon (1906). 2 5 0 ~ . 850°C ' O Buu (1 990). 2 9 6 ~ . 89SeC
BUU (1990). 296pm. 815% . v ûaw (1990). 2 9 6 ~ . 730% Wu, Orace et al. (19896). 250-30ûp1n. 880%
- - I - .. I
b
- M
r
b
e
h
CI
I l i 1 1 I I I I . S 10 20 30 60 50 60
Suspension Dcnsity, ~ v m '
Figure 2- 13. The relationship between the heat transfer coefficient and bed suspension density (Basu and Nag, 1996)
Figure 2- 14. The relationship between the heat transfer coefficient and bed temperature (Wu sr al., 1989a)
(4) Soli& circulating rate
Heat trmsfer coefficients were found to increase with increasing solids circulation
rate at a given superficial gas velocity (Feugier et al. 1987). This is because increasing
the solids circulating rate increases suspension density in the fluidized bed, which in tum
increases solids concentration on the heat transfer surface.
(5) Particle size
An effect of particle size on the heat transfer coefficient is observed on the short
heat &ansferring surfaces. There is considerable evidence (Subbarao and Basu 1 986, Wu
et al., I989a, Basu et al., 1990) that heat transfer coefficients increase with a decrease in
mean particle diameter (Figure 2-15). Finer particles can cause higher heat transfer
coefficients due to several effects: (1) Small particles would irnprove heat transfer
effectively because of its shorter average distances for conduction between the wall and
adjacent particles and the small thermal conductive resistance. (2) Small particles result
in higher heat transfer coefficients because smaller particles cari increase the effective
heat transfer area covered by particles .
The influence of particle diameter is significant oniy for results obtained with
short heat transfer surfaces. If the heat transfer surface is long, the particles contact with
the surface for too long time and reach thermal equilibriurn as the particles faIl down
along the wall. Both large and small particles have sufficient time to equilibrate in
temperature with the heat transfer surface. Therefore, the dominant factor gives way to
the length of heat transfer surface, and the effect of particle size seems to diminish
(Glicksman, 1988).
Figure 2- 1 5 . The relationship between the heat trans fer coefficient and the particle size (Wu et aL, 1989)
(6) nie Zength of heat tramfer surfiace
Heat transfer coefficients for a longer surface were found to be usually lower than
those for a short surface. When particles move down dong a longer heat transfer surface,
their temperature tends to equilibrate with the surface temperature. Therefore, there is a
decrease of temperature ciifference between the particle and heat transfer surface,
resulting in lower overall heat transfer rate. Nag and Moral (1991) plotted the particle
Nusselt nurnber against a dimensionless probe height at two different superficial
velocities. It is observed that Nu decreases with an increase of W D as well as an increase
of superficial velocity. They used their own experimental data to obtain the Functional
relationship
where Lb is the length of the heat transfer probe, and D is the column diameter.
The local solid concentration at the wall has been reported to depend on the
geometric configuration of the reactor (Lorkhart et ai., 1995). The walls of most CFB
combustion units are generally not smooth. The walls are usually made of membrane
tubes where two adjacent tubes are welded to a fin, foming a cavity (Figure 2-16). It is
expected that the presence of the membrane tubes on the wall will affect the local solids
concentration in the wall region, hence affecting heat transfer behavior. Visual
observations made on the wall of a I2MWe boiler (Golriz, 1994) show that particles
concentrate over the fins and tend to stay there longer than those traveling on the crest of
the tubes. Lockhart et al. (1995) measured the local solids concentration and heat transfer
coefficient in a 152 mm diameter 9.3 m ta11 CFB. It was found that solid holdup is higher
in the fin region. Therefore, for small heat transfer surfaces, the heat transfer coefficient is
higher in the fin rather than in the crest region (Figure 2- 17). However this may not be
Fut m u i d i d Suspension
Figure 2- 16. Schematic of the membrane heat transfer surface
400.
Sensor Tube Position
4 #1 32 mm fin
O 20 40 60 80 1 O0
Suspension Density (kg/& )
300
200
100
Average Solids Concentration (%)
- + #2 32 mm crest + #3 19mm crest
-
-
1 Tube Cmst Cwner Rn
O l I I 1 1 O 20 40 60 80 1
Suspension Density (kg/& )
Figure 2- 17. Heat transfer coefficient measured in different regions of a membrane wall (Lockhart et al., 1995)
true for long heat transfer surface, because particles are relatively protected and sweep
much further along the fins than along the crests, so that particles in contact with the fins
are more likely to approach thermal equilibrium with the surface than the particles near
the crests. Therefore, long membrane surfaces may have higher heat transfer coefficients
for the crests than for the fins.
2.3.5. Heat Transfer Modelling
For proper design of circulating fluidized bed reactors, it is important to know the
effect of design and operating parameten on the bed to waIYsurfaces heat transfer
process. The mechanistic models can predict this effect in some ways.
One useful model for heat transfer fiorn the bed to the wall in a CFB is developed
from the model proposed by Mickley and Fairbanks (1955). The heat transfer was
represented as a transient process from the "packet", a group of particles, to the wall with
the resulting heat flux related to the residence time of the 'packet" at the surface. The
cluster of particles at the wall could be modeled as a homogeneous matenal with a
constant effective conductivity and heat capacity. If the cluster is thicker than the thermal
penetration depth, the transient heat transfer solution for a semi-infinite body can be used.
The instantaneous heat transfer coefficient and time-rnean heat transfer coefficients are,
respectively :
Where k, is thermal conductivity of the "packet" and t is the contact time. This mode1
gives reasonable values for relatively long residence times, but it leads to unrealistically
hi& values of the heat transfer rate at low residence times.
The circulating fluidized bed norrnally operates in the fast bed regime. Visual
observations and video tapes show that CFB beds comprise of dense clusten or strands
and a gas phase continuum with dispersed solids. The agglomeration of solids into
clusters (or strands) is a major characteristic of the circulating fluidized bed. When
clusters slide over the wall an unsteady-state heat conduction into the serni-infinite
clusters takes place. In addition, there is a gas film resistance on the wall between layers
of particles and the wall. The downward travel of these clusters takes place primarily in
the gas-solid boundary layer (annulus) at the bed wall. Experimental (Wu et al., 1989a, b)
as well as visual studies on the walls of a fast bed show that the wall is swept by discrete
clusters instead of a continuous film of solids. The intermittent nature of the strands on
the wall was also observed in large commercial CFB boilers (Leckner, 199 1b; Couturier
et al., 1993). The clusters, after travelling a certain distance, dissolve or detach
themselves fiom the wall, and then replaced by new clusters. Using the above
information, the time-average heat transfer coefficient due to particle convection is
written as (Mahalingam and Kola. 1991):
Where kg is thermal conductivity of gas, k, is thermal conductivity of cluster, cc is cluster
concentration.
To determine the gas convection heat transfer coefficient which is contributed by the gas
phase in contact with the wall; the correlation of Wen and Miller (1 96 1 ) for heat tram fer
fiom the dilute phase c m be used:
where pd, is the suspension density in the up-flow of gas-solid mixture.
For determination of radiative heat transfer, h, the clusters, which are away from the wall
or in contact with it, are assumed to be at the bed temperature Tb. The radiant heat
exchange behveen the cluster and the wall, both being considered gray, is given by:
Qr = h , ( d 4 ) D' (T-T,)
wheref,., is the cluster-to-wall view factor, which depends on the shape, disposition and
emissivities of the two bodies:
1
The average heat transfer coefficient in a circulating fluidized bed is given by
where 6, is the fiaction of the heat transfer surface exposed to the cluster and (1-dC) the
bction to voids or the dispened phase at any instant.
Sustracthg the corresponding equations into eqn. (2-13), we have:
These models usually can successfûlly predict the most of effects of physical variables
and operating conditions on heat transfer processes and provide some important
information for the design of circulating fluidized bed, but are still not very satisfactory
for quantitative calculation
2.3. Cas and Solids Mixing in Circulating Fluidized Beds
The contact efficiency between gas and solids is closely related to hydrodynarnics, mass
and heat transfer behavion in circulating fluidized beds and has significantly influence to
the overall system performance. However, the contact efficiency of gas-solids in a
circulating fluidized bed is less understood, in part, due to the difficulty of measurement
and the lack of a uniform definition for gas-solids contact efficiency.
A nurnber of investigators have measured the contact efficiency between gas and
solids in CFB riser by indirect methods (Sun and Grace, 1990; Kagawa et al.. 199 1 ), but
these required several assumptions which are difficult to veriS. Wei et al. (1994) and
Rogues et al. (1 994) used the phosphorescent particle tracer technique to measure the
solids residence tirne distribution (RTD) in the downer, but this experirnent process was
comparatively cornplex. In general, the gas-solids mixing is related to gas and solids
interphase heat transfer behavior in the reactor, and is the combination of flow
hydrodynarnics and reaction efficiency. Previous studies on the m a s transfer and heat
transfer in fluidized beds (Kunii and Levenspiel 1991, Richardson and Szekely 1961)
have recognized that the mechanisms for gas-solids mass and heat transfers are related,
and mass transfer may occur simultaneously with the transfer of heat. The gas-solids
interphase heat transfer process is controlled by the gas and solids mixing behavior, and
effective gas-solids mixing c m normally improve the heat exchange between the two
phases. Thus, the results of heat transfer between gas and solids can be used to estimate
gas-solids contact efficiency and m a s transfer coefficients. Therefore, it is possible to
use the thermal method to investigate the gas-solids mixing phenomena in fluidized beds
and this is a more direct and sirnpler method to study the gas and solids contact
efficiency.
Dry et al. (1987, 1992), Dry and White (1989) employed a thermal method to
measure the gas residence time and estimate contact efficiency between gas and solids in
a high-velocity upflow fluidized bed of FCC catalyst. Their technique involved the use of
heated air as a gas tracer and the detection of the air temperature in the bed by a rapid-
response thermocouple. A smail flow of auxiliary air, previously heated to 550 OC in an
electric Fumace, was adrnitted into the riser bottom for a period of tirne. Wherever gas-
solid mixing in the riser was intimate, local gas and solids temperatures would
equilibrate. Where mixing is less complete, the gas would retain a temperature higher
than that expected at gas-solid suspension equilibnum. The temperature of the gas in the
bed was monitored by an ultrafine thermocouple mounted in an aspirating probe. Solids
were prevented from entering the probe by a thin porous filter. The thermal response was
computed to obtain the contact efficiency beiween the gas and solids (see section 3.4.2.
for details).
In their senes of studies, it was found that the effects of reactor geometry, gas
inlet structure, operating condition and particle type al1 have the significant effect on the
gas-solids contact efficiency. Their results showed that fine particles could give better
contacting than coarser particles, and the overall contact efficiency decreased with
decreasing inlet diameter. It is believed that the results fiom these experiments can
provide very important information for riser reactor design and operation. A similar but
improved approach was employed in our study to obtain the gas-solids mixing efficiency
in the entrance region of the downer, which has evident effect on the industrial
applications and developments of the downer reactors.
CHAPTER 3. EXPERIMENTAL EQUIPMENT
3.1. The Circulating Fluidized Bed Unit
The nser/downer circulating fluidized bed was made of plexiglass with a downer
of 100 mm in inner diameter and 9.3 m in height, and an accompanying riser of diameter
100 mm and height of 15.1 m (Figure 3-1). There are also a 3.15 m high, 0.25 m i.d. steel
solids storage tank, a 3.35 m hi&, 0.10 m i.d. acrylic solids circulation measuring vessel,
an inertial downer separator, a primary riser cyclone, two secondary and one combined
tertiary cyclones. Solids fa11 £iom the storage tank through a butterfly valve into the
bottom of the riser and are entrained upwards fiom the bottom of the riser with the high
velocity riser fluidization air through a rnulti-tube air distributor. When the gas and solids
flow upward to the top of the riser, the gas is separated by the prùnary cyclone and then
M e r cleaned by secondary and tertiary cyclones, before finally entering a baghouse
filter. The sotids are fed into the downer through a solids distributor. A separate air
Stream is added to the downer through a gas disûibutor. At the bottom of the downer, the
gas-solids suspension enten a fast separator where most of the entrained solids are
recovered and retunied to the storage tank. The remaining solids are then removed by
additional secondary and tertiary cyclones. The solids circulating rate is regulated by the
solids control valve and c m be measured by deflecting the collected solids into the
measuring tank.
The pas-solids separator (1.07 m x 0.38 m x 0.75 m) in the present research was
designed sirnilarly to a successful unit developed at Tsinghua University. It is a simple
inertial separator in which gas and solids suspension fust pass through a specially
designed nozzle and then irnpinge on a c w e d guiding plate with a gradually increasing
radius where more than 99 % of the solids are separated from the gas phase.
Riscr primary cyclone downer disributor
Riser
(O. I rn i . d i I S. 1 rn)
Downer rncrtial fast scparator
%
Riser distributor
Riser disaibutor air - :
Downer disnibutor air
+ Downer main air
1 Tcniary cyclone
Storage tank (0.25 rn i .b/ 2 m)
' Solids flow rontrol valve Riser main air
Air out t
