High Density Solids Downflow Gas-Solids Reactors Density Solids Downflow Gas-Solids Reactors by ... downflow operation is defined as operation in ... hydrodynamic behaviour of fast

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.

Engstrom, F. and Lee, Y. Y. (1991), "Future Chailenges of Circuiating Fluidized Bed

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.,

96,255-261.

Huang, W-X. and Zhu, J X , '%xperimental Study on Solids Acceleration Length in a

Long CFB Risd', Chm. Eng. Sci., submitted, April 1998

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.

Sci., in press.

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York.

Kullendorff, A. and Andersson, S. (1986), "A General Review of Combustion in

Circulahg Fluidized Beds", Circulating F7uidired Bed Technologv. (ed. P. Basu),

pp. 83-96, Pergamon Press, Toronto.

Shingles, T. and McDonald, A F. (1988), "Commercial Experience with Synthol CFB

Reactors", CirniatingFIuz?i&ed Bed Technooloogy D, (eds. P. Basu and 1. F. Large),

pp. 43-50, Pergamon Press, Toronto.

Steynberg, A. P., Shingies, T., Dry, M. E, Jager, B. and Yukawa, Y (1991). "Sasol

Commercial Scale Experience with Synthol FFB and CFB Catalytic Fischer-Tropsch

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Hasatani), pp. 527-532, Pergamon Press, Toronto.

Wang, Z., Bai, D. and Jin, Y. (1992), b'EZydrodynami~~ of Cocunent Downflow

Circulating Fluidized Bed (CDCFB)", Powder Technul-. 27 1-275.

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the Circulating Fluidized Bed", AIC7i.E Symp- Ser., @(296), 81-90.

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Solids Downer Fluidized Bed", Cun. J. Chem. Eng., a ( 2 ) , in press.

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Fluidized Beds (1): Solids Concentration Profies and Pressure Gradient

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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)