fluidization solids handling and processing industrial applications

20 3.3 Influence of Bed Diameter on Circulating Fluidized Beds .... For example, the two phase theory equating the excess gas velocity above minimum fluidization to the visible bubble fl

HANDLING, AND PROCESSING

tronic or mechanical, including photocopying,

recording or by any information storage and

retrieval system, without permission in writing

from the Publisher.

Library of Congress Catalog Card Number: 98-18924

ISBN: 0-8155-1427-1

Printed in the United States

Published in the United States of America by

Noyes Publications

369 Fairview Avenue, Westwood, New Jersey 07675

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Library of Congress Cataloging-in-Publication Data

Fluidization, solids handling, and processing : industrial

applications / edited by Wen-Ching Yang.

Series Editor: Liang-Shih Fan, Ohio State University

FLUIDIZATION, SOLIDS HANDLING, AND PROCESSING: Edited by Wen-Ching Yang INSTRUMENTATION FOR FLUID-PARTICLE FLOWS: by S L Soo

University of RochesterRochester, NY

Jack Reese

Department of ChemicalEngineering

Ohio State UniversityColumbus, OH

Ohio State UniversityColumbus, OH

Gabriel I Tardos

Department of ChemicalEngineering

City College of City University ofNew York

New York, NY

Richard Turton

Department of ChemicalEngineering

West Virginia UniversityMorgantown, WV

University of WollongongWollongong, NSW, Australia

Shang-Tian Yang

Department of ChemicalEngineering

Ohio State UniversityColumbus, OH

Wen-Ching Yang

Science and Technology CenterSiemens Westinghouse PowerCorporation

This volume, Fluidization, Solids Handling, and Processing, is the first of a series of volumes on “Particle Technology” to be published

by Noyes Publications with L S Fan of Ohio State University as theconsulting editor Particles are important products of chemical processindustries spanning the basic and specialty chemicals, agricultural products,pharmaceuticals, paints, dyestuffs and pigments, cement, ceramics, and

electronic materials Solids handling and processing technologies are

thus essential to the operation and competitiveness of these industries

Fluidization technology is employed not only in chemical production, it also

is applied in coal gasification and combustion for power generation, mineralprocessing, food processing, soil washing and other related waste treatment,environmental remediation, and resource recovery processes The FCC(Fluid Catalytic Cracking) technology commonly employed in the modernpetroleum refineries is also based on the fluidization principles

There are already many books published on the subjects of tion, solids handling, and processing On first thought, I was skeptical aboutthe wisdom and necessity of one more book on these subjects On closerexamination, however, I found that some industrially important subjectswere either not covered in those books or were skimpily rendered It would

fluidiza-be a good service to the profession and the engineering community toassemble all these topics in one volume In this book, I have invitedrecognized experts in their respective areas to provide a detailed treatment

of those industrially important subjects The subject areas covered in this

book were selected based on two criteria: (i) the subjects are of industrial

importance, and (ii) the subjects have not been covered extensively in

books published to date

The chapter on fluidized bed scaleup provides a stimulating

approach to scale up fluidized beds Although the scaleup issues are by no

means resolved, the discussion improves the understanding of the issues and

provides reassessments of current approaches The pressure and

tem-perature effects and heat transfer in fluidized beds are covered in

separate chapters They provide important information to quantify the

effects of pressure and temperature The gas distributor and plenum

design, critical and always neglected in other books, are discussed in detail.

For some applications, the conventional fluidized beds are not necessarily the

best Special design features can usually achieve the objective cheaper and

be more forgiving Two of the non-conventional fluidized beds,

recirculat-ing fluidized beds with a draft tube and jettrecirculat-ing fluidized beds, are

intro-duced and their design approaches discussed Fluidized bed coating and

granulation, applied primarily in the pharmaceutical industry, is treated

from the fluidization and chemical engineering point of view Attrition,

which is critical in design and operation of fluidized beds and pneumatic

transport lines, is discussed in detail in a separate chapter Fluidization with

no bubbles to minimize bypassing, bubbleless fluidization, points to

potential areas of application of this technology The industrial

applica-tions of the ever-increasingly important three-phase fluidization systems

are included as well The developments in dense phase conveying and in

long distance pneumatic transport with pipe branching are treated

sepa-rately in two chapters The cyclone, the most common component

em-ployed in plants handling solids and often misunderstood, is elucidated by

an experienced practitioner in the industry The book is concluded with a

discussion on electrostatics and dust explosion by an electrical engineer

This book is not supposed to be all things to all engineers The

primary emphasis of the book is for industrial applications and the primary

audience is expected to be the practitioners of the art of fluidization, solids

handling, and processing It will be particularly beneficial for engineers who

operate or design plants where solids are handled, transported, and

pro-cessed using fluidization technology The book, however, can also be useful

as a reference book for students, teachers, and managers who study particle

technology, especially in the areas of application of fluidization technology

and pneumatic transport

I’d like to take this opportunity to thank Professor Fan who showedconfidence in me to take up this task and was always supportive I’d alsolike to thank the authors who contributed to this book despite their busyschedules All of them are recognized and respected experts in the areasthey wrote about The most appreciation goes to my wife, Rae, whoendured many missing weekends while I worked alone in the office.

Final determination of the suitability of any information or product for use contemplated by any user, and the manner of that use, is the sole responsibility of the user We recommend that anyone intending

to rely on any recommendation of materials or procedures mentioned

in this publication should satisfy himself as to such suitability, and that he can meet all applicable safety and health standards.

1 Fluidized Bed Scale-up 1

Leon R Glicksman 1.0 INTRODUCTION 1

2.0 REACTOR MODELING: BED DIAMETER INFLUENCE 4

3.0 INFLUENCE OF BED DIAMETER ON HYDRODYNAMICS 10

3.1 Bubbling Beds 10

3.2 Mixing 20

3.3 Influence of Bed Diameter on Circulating Fluidized Beds 22

3.4 Flow Transition 25

4.0 EXPERIMENTAL MEANS TO ACCOUNT FOR SCALE-UP: USE OF SCALE MODELS 26

4.1 Development of Scaling Parameters 27

4.2 Governing Equations 29

4.3 Fluid-Solid Forces 35

4.4 Spouting and Slugging Beds 38

5.0 SIMPLIFIED SCALING RELATIONSHIPS 39

5.1 Low Reynolds Number 39

5.2 High Reynolds Numbers 41

5.3 Low Slip Velocity 42

5.4 General Case 43

5.5 Range of Validity of Simplified Scaling 44

6.0 FURTHER SIMPLIFICATIONS IN THE SCALING

RELATIONSHIP 51

6.1 Viscous Limit 51

6.2 Other Derivations for Circulating Fluidized Beds 54

6.3 Deterministic Chaos 55

7.0 DESIGN OF SCALE MODELS 56

7.1 Full Set of Scaling Relationships 56

7.2 Design of Scale Models Using the Simplified Set of Scaling Relationships 61

8.0 EXPERIMENTAL VERIFICATION OF SCALING LAWS FOR BUBBLING BEDS 65

8.1 Hydrodynamic Scaling of Bubbling Beds 65

8.2 Verification of Scaling Relationships for Bubbling and Slugging Beds 69

8.3 Verification of Scaling Laws for Spouting Beds 75

8.4 Verification of Scaling Relationships for Pressurized Bubbling Beds 76

9.0 APPLICATIONS OF SCALING TO COMMERCIAL BUBBLING FLUIDIZED BED UNITS 80

10.0 HYDRODYNAMIC SCALING OF CIRCULATING BEDS 91

11.0 CONCLUSIONS 100

ACKNOWLEDGMENTS 102

NOTATIONS 103

REFERENCES 104

2 Pressure and Temperature Effects in Fluid-Particle Systems 111

Ted M Knowlton 1.0 INTRODUCTION 111

1.1 Minimum Fluidization Velocity 113

1.2 Bed Voidage and Bed Expansion 120

1.3 Bubbles in Fluidized Beds 124

1.4 Bubble Size and Frequency 125

1.5 Bed-to-Surface Heat Transfer Coefficient 129

1.6 Entrainment and Transport Disengaging Height 131

1.7 Particle Attrition at Grids 134

1.8 Particle Attrition in Cyclones 136

1.9 Jet Penetration 137

1.10 Regime Transitions 139

1.11 Cyclone Efficiency 146

NOTATIONS 147

REFERENCES 149

3 Heat Transfer in Fluidized Beds 153

John C Chen 1.0 INTRODUCTION 153

2.0 BUBBLING DENSE FLUIDIZATION 154

2.1 Hydrodynamic Characteristic 154

2.2 Heat Transfer to Submerged Surfaces 155

3.0 CIRCULATING FAST FLUIDIZATION 173

3.1 Hydrodynamic Characteristics 173

3.2 Heat Transfer 178

NOTATIONS 201

Subscripts 202

REFERENCES 202

4 Gas Distributor and Plenum Design in Fluidized Beds 209

S.B Reddy Karri and Ted M Knowlton 1.0 INTRODUCTION 209

2.0 TYPES OF GRIDS 210

2.1 Perforated Plates (Upwardly-Directed Flow) 210

2.2 Bubble Cap (Laterally-Directed Flow) 210

2.3 Sparger (Laterally or Downwardly-Directed Flow) 211

2.4 Conical Grids (Laterally-Directed Flow) 211

3.0 GRID DESIGN CRITERIA 212

3.1 Jet Penetration 212

3.2 Grid Pressure-Drop Criteria 214

3.3 Design Equations 215

3.4 Additional Criteria for Sparger Grids 218

3.5 Port Shrouding or Nozzle Sizing 219

4.0 PARTICLE ATTRITION AT GRIDS 220

4.1 Attrition Correlation 222

5.0 EROSION 223

6.0 EFFECTS OF TEMPERATURE AND PRESSURE 223

7.0 PLENUM DESIGN 223

8.0 DESIGN EXAMPLES 225

8.1 FCC Grid Design 225

8.2 Polyethylene Reactor Grid Design 230

NOTATIONS 233

REFERENCES 235

5 Engineering and Applications of Recirculating and Jetting Fluidized Beds 236

Wen-Ching Yang 1.0 INTRODUCTION 236

2.0 RECIRCULATING FLUIDIZED BEDS WITH A DRAFT TUBE 237

2.1 Draft Tube Operated As A Fluidized Bed 240

2.2 Draft Tube Operated As A Pneumatic Transport Tube 242

2.3 Design Example for a Recirculating Fluidized Bed with a Draft Tube 257

2.4 Industrial Applications 263

3.0 JETTING FLUIDIZED BEDS 264

3.1 Jet Penetration and Bubble Dynamics 265

3.2 Gas Mixing Around the Jetting Region 281

3.3 Solids Circulation in Jetting Fluidized Beds 295

3.4 Fines Residence Time in Jetting Fluidized Beds 315

3.5 Scale-up Considerations 317

3.6 Applications 319

NOTATIONS 319

Greek Letters 322

REFERENCES 323

6 Fluidized Bed Coating and Granulation 331

Richard Turton, Gabriel I Tardos, and Bryan J Ennis 1.0 INTRODUCTION 331

2.0 COATING OF PARTICLES IN FLUIDIZED BEDS 333

2.1 Introduction 333

2.2 Overview of Coating Process 335

2.3 Microscopic Phenomena 339

2.4 Modelling 344

2.5 Design Criteria 355

3.0 GRANULATION OF FINE POWDERS IN FLUIDIZED BEDS 365

3.1 Introduction 365

3.2 Microscopic Phenomena 366

3.3 Granule Growth Kinetics 380

3.4 Experimental Support and Theoretical Predictions 387

3.5 Granule Consolidation, Attrition and Breakage 398

3.6 Modeling of Granulation Processes 406

3.7 Unwanted Aggregation in Fluidized Beds 418

ACKNOWLEDGMENT 424

NOTATIONS 424

REFERENCES 429

7 Attrition in Fluidized Beds and Pneumatic Conveying Lines 435

Joachim Werther and Jens Reppenhagen 1.0 INTRODUCTION 435

2.0 FACTORS AFFECTING ATTRITION 437

2.1 Material Properties 438

2.2 Process Conditions 440

3.0 ASSESSMENT OF ATTRITION 444

3.1 Breakage and Selection Functions 444

3.2 Attrition Rate 445

3.3 Friability Indices 446

3.4 Grindability Indices 446

4.0 ATTRITION TESTS 447

4.1 Friability Tests 447

4.2 Experiments to Study Attrition Mechanisms 448

4.3 Test Equipment and Procedures 449

5.0 ATTRITION IN FLUIDIZED BED SYSTEMS 455

5.1 Sources of Attrition 455

5.2 Attrition in the Overall Fluidized Bed System, Continuous Processes 473

5.3 Steps to Minimize Attrition in Fluidized Beds 475

6.0 ATTRITION IN PNEUMATIC CONVEYING LINES 478

6.1 Modeling 480

6.2 Parameter Effects 480

6.3 Steps to Minimize Attrition in Pneumatic Conveying Lines 482

NOTATIONS 484

Subscripts 485

Greek Symbols 486

REFERENCES 486

8 Bubbleless Fluidization 492

Mooson Kwauk 1.0 INTRODUCTION 492

2.0 FLUIDIZED LEACHING AND WASHING 492

2.1 Uniform Particles 496

2.2 Mixed Particles 500

2.3 Staged Fluidized Leaching (SFL) 502

3.0 BUBBLELESS GAS/SOLID CONTACTING 502

3.1 Bubbling Fluidization and G/S Contacting Efficiency 502

3.2 Species of Bubbleless G/S Contacting 507

4.0 DILUTE RAINING FLUIDIZATION 508

4.1 Raining Particles Heat Exchanger 508

4.2 Experimental Verification 512

4.3 Baffling and Particles Distribution 515

4.4 Pilot Plant Demonstration 519

5.0 FAST FLUIDIZATION 523

5.1 Longitudinal Voidage Distribution 525

5.2 Regimes for Vertical G/S Systems 529

5.3 Radial Voidage Distribution 533

5.4 Modeling Fast Fluid-bed Reactors 533

6.0 SHALLOW FLUID BEDS 537

6.1 Dynamics for the Distributor Zone 537

6.2 Activated Solids Shallow Fluid Bed Heat Exchanger 537

6.3 Cocurrent Multistage Shallow Fluid Bed 541

6.4 The Co-MSFB as a Chemical Reactor 545

7.0 FLUIDIZATION WITH NO NET FLUID FLOW 546

7.1 Levitation of Discrete Particles 547

7.2 Semi-Fluidization through Oscillatory Flow 551

7.3 Application to Pseudo Solid-Solid Reactions 553

8.0 PARTICLES WHICH QUALIFY FOR BUBBLELESS OPERATION 556

8.1 Powder Characterization 556

8.2 Improving Fluidization by Particle Size Adjustment 562

9.0 WHY BUBBLING AND NOT PARTICULATE FLUIDIZATION 569 9.1 The Energy-Minimized Multiscale (EMMS) Model 570

9.2 Reconciling L/S and G/S Systems 573

10.0 EPILOGUE 576

NOTATIONS 576

REFERENCES 578

9 Industrial Applications of Three-Phase Fluidization Systems 582

Jack Reese, Ellen M Silva, Shang-Tian Yang, and Liang-Shih Fan 1.0 INTRODUCTION 582

Part I: Smelting Reduction, Paper Processing, and Chemical Processing 588

2.0 SMELTING REDUCTION 588

2.1 Introduction 588

2.2 Principles of Smelting Reduction 590

2.3 Post-Combustion and Heat Transfer in SRF 593

2.4 Slag Layer Behavior 599

2.5 Future of Smelting Reduction of Iron Ore 603

3.0 PAPER PROCESSING 604

3.1 Introduction 604

3.2 Chemical Pulping of Wood Chips 605

3.3 Pulp Bleaching and Flotation De-inking 609

4.0 CHEMICAL PROCESSING 614

4.1 Introduction 614

4.2 Hydrotreating/Hydrocracking Petroleum Intermediates 614

4.3 Fischer-Tropsch Synthesis 619

4.4 Methanol Synthesis 621

Part II: Three-Phase Biofluidization 623

5.0 BIOLOGICAL APPLICATIONS OF THREE-PHASE FLUIDIZATION 623

5.1 Introduction 623

5.2 Applications 629

5.3 Bioparticles 637

5.4 Hydrodynamics 643

5.5 Phase Mixing in a Three-Phase Reactor 647

5.6 Mass Transfer 648

5.7 Modeling 651

5.8 Scale Up 653

5.9 Process Strategy 655

5.10 Novel Reactors 657

5.11 Economics 661

5.12 Summary 662

ACKNOWLEDGMENT 663

NOTATIONS 663

REFERENCES 664

10 Dense Phase Conveying 683

George E Klinzing 1.0 INTRODUCTION 683

2.0 ADVANTAGES OF DENSE PHASE CONVEYING 693

3.0 BASIC PHYSICS 695

4.0 PULSED PISTON FLOWS 698

5.0 VERTICAL FLOW SYSTEMS 706

6.0 BOOSTERS 708

NOTATIONS 709

Greek 709

Subscripts 710

REFERENCES 710

11 Design Considerations of Long-Distance Pneumatic Transport and Pipe Branching 712

Peter W Wypych 1.0 INTRODUCTION 712

2.0 LONG-DISTANCE PNEUMATIC CONVEYING 713

2.1 Product Characterization and Classification 714

2.2 Blow Tank Design 733

2.3 Conveying Characteristics 738

2.4 Pressure Drop Prediction 741

2.5 Stepped-Diameter Pipelines 747

2.6 Valves 748

2.7 Pipeline Unblocking Techniques 751

2.8 General Considerations 752

3.0 PIPE BRANCHING 753

3.1 Dust Extraction 754

3.2 Flow Splitting 760

3.3 Pressure Loss 766

NOTATIONS 767

REFERENCES 769

12 Cyclone Design 773

Frederick A Zenz 1.0 INTRODUCTION 773

2.0 REQUIRED DESIGN DATA 774

3.0 CORRELATING FRACTIONAL COLLECTION EFFICIENCY 775

4.0 EFFECT OF SOLIDS LOADING 778

5.0 CYCLONE LENGTH 778

6.0 CONES, DUST HOPPERS AND EROSION 780

7.0 CYCLONE INLET AND OUTLET CONFIGURATIONS 781

8.0 THE COUPLING EFFECT 785

9.0 PRESSURE DROP 787

10.0 SPECIAL CASES 788

11.0 BED PARTICLE SIZE DISTRIBUTION AND CYCLONE DESIGN 791

12.0 CENTRIFUGAL VERSUS CENTRIPETAL CUT POINT PARTICLE SIZE 793

13.0 CYCLONE DESIGN EXAMPLES 794

14.0 ALTERNATE APPROACH TO SOLVING EXAMPLE B 804

15.0 ALTERNATE APPROACH TO SOLVING EXAMPLE C 809

16.0 DIPLEG SIZING AND CYCLONE PRESSURE BALANCE 812

NOTATIONS 814

REFERENCES 815

13 Electrostatics and Dust Explosions in Powder Handling 817

Thomas B Jones 1.0 INTRODUCTION 817

2.0 CHARGING OF SOLID PARTICLES 818

2.1 Triboelectrification 819

2.2 Charge Relaxation 823

2.3 Induction Charging of Particles 824

2.4 Electrostatic Fields and Potentials 825

3.0 FLUIDIZED BED ELECTRIFICATION 829

3.1 Background 829

3.2 More Recent Work 832

3.3 Beneficial Effects of Electric Charge 836

4.0 ESD DUST IGNITION HAZARDS 836

4.1 Basics of Suspended Solids Ignition 837

4.2 Types of Discharges 841

4.3 Charge Dissipation 850

5.0 ESD HAZARDS IN FLUIDIZED BED SYSTEMS 854

5.1 Hazards Associated with Fluidization 855

5.2 Hazards in Peripheral Equipment and Processes 857

5.3 Other Nuisances and Hazards 863

6.0 CONCLUSION 864

ACKNOWLEDGMENT 866

REFERENCES 867

Index 872

Although fluidized beds have been used extensively in cial operations such as fluidized bed combustors and fluid catalytic crack-ing, engineers are still faced with uncertainties when developing newcommercial designs Typically, the development process involves alaboratory bench scale unit, a larger pilot plant, and a still larger demon-stration unit Many of the important operating characteristics can changebetween the different size units There is a critical problem of scale-up:how to accurately account for the performance changes with plant size toinsure that a full size commercial unit will achieve satisfactory perfor-mance In addition, it would be helpful if the smaller units could be used

commer-to optimize the commercial plant or solve existing problems

One discouraging problem is the decrease in reactor or combustorperformance when a pilot plant is scaled up to a larger commercial plant.These problems can be related to poor gas flow patterns, undesirable solidmixing patterns and physical operating problems (Matsen, 1985) In thesynthol CFB reactors constructed in South Africa, first scale-up from thepilot plant increased the gas throughput by a factor of 500 Shingles andMcDonald (1988) describe the severe problems initially encountered andtheir resolution

Fluidized Bed Scale-up

Leon R Glicksman

In some scaled up fluidized bed combustors, the lower tion zone has been divided into two separate subsections, sometimesreferred to as a “pant leg” design, to provide better mixing of fuel andsorbent in a smaller effective cross section and reduce the potentialmaldistribution problems in the scaled up plant.

combus-Matsen (1985) pointed out a number of additional problem areas

in scale-up such as consideration of particle size balances which changeover time due to reaction, attrition and agglomeration Erosion of cy-clones, slide valves and other components due to abrasive particles areimportant design considerations for commercial units which may not beresolved in pilot plants

If mixing rates and gas-solid contacting efficiencies are keptconstant between beds of different size, then thermal characteristics andchemical reaction rates should be similar However, in general, the bedhydrodynamics will not remain similar In some instances, the flowregime may change between small and large beds even when using thesame particles, superficial gas velocity and particle circulation rate perunit area The issue of scale-up involves an understanding of thesehydrodynamic changes and how they, in turn, influence chemical andthermal conditions by variations in gas-solid contact, residence time, solidcirculation and mixing and gas distribution

There are several avenues open to deal with scale-up Numericalmodels have been developed based on fundamental principals Themodels range from simple one-dimensional calculations to complex mul-tidimensional computational fluid dynamics solutions There is no doubtthat such first principal models are a great aid in synthesizing test data andguiding the development of rational correlations In a recent modelevaluation, modelers where given the geometry and operating parametersfor several different circulating beds and asked to predict the hydrody-namic characteristics without prior knowledge of the test results (Knowlton

et al 1995) None of the analytical or numerical models could reliablypredict all of the test conditions Few of the models could come close topredicting the correct vertical distribution of solid density in the riser andnone could do it for all of the test cases! Although it is tempting to thinkthat these problems can be solved with the “next generation of comput-ers,” until there is general agreement and thorough verification of thefundamental equations used to describe the hydrodynamics, the numericalmodels will not stand alone as reliable scale-up tools

On the other hand, there is a blizzard of empirical and empirical correlations which exist in the fluidized bed literature to predict

semi-fluid dynamic behavior In addition there are probably a large number of

proprietary correlations used by individual companies The danger lies in

extrapolating these relations to new geometric configurations of the riser

or inlet, to flow conditions outside the range of previous data, or to beds of

much different sizes Avidan and coauthors in a 1990 review of FCC

summed up the state of the art: “basic understanding of complex

fluidiza-tion phenomena is almost completely lacking While many FCC licensors

and operators have a large body of in-house proprietary data and

correla-tions, some of these are not adequate, and fail when extrapolated beyond

their data base.” (Avidan, et al., 1990.)

As a example, consider the influence of mean particle size In the

early work on bubbling fluidized bed combustors, attempts were made to

use relations from the classic fluidization literature which had

concen-trated on FCC applications with much smaller particles In many cases, it

was discovered that the relationships for small particles gave erroneous

results for combustors with much larger particles For example, the two phase

theory equating the excess gas velocity above minimum fluidization to the

visible bubble flow was shown to be severely distorted for large particle

systems Jones and Glicksman (1985) showed that the visible bubble flow

in a bubbling bed combustor was less than one fifth of u o -u mf In other

cases even the trends of the parametric behavior were changed Heat transfer

to immersed surfaces in fine particle bubbling beds increased strongly

with a decrease in the mean particle size For large particle beds, the heat

transfer, in some instances, decreased with a decreased particle diameter

Another approach to scale-up is the use of simplified models with

key parameters or lumped coefficients found by experiments in large

beds For example, May (1959) used a large scale cold reactor model

during the scale-up of the fluid hydroforming process When using the

large cold models, one must be sure that the cold model properly simulates

the hydrodynamics of the real process which operates at elevated pressure

and temperature

Johnsson, Grace and Graham (1987) have shown one example of

verification of a model for 2.13 m diameter industrial phthalic anhydride

reactor Several bubbling bed models gave good overall prediction of

conversion and selectivity when proper reaction kinetics were used along

with a good estimate of the bubble size The results were shown to be

quite sensitive to the bubble diameter The comparison is a good check of

the models but the models are incomplete without the key hydrodynamic

data In this case, the bubble size estimates were obtained from

measure-ments of overall bed density in the reactor

As Matsen expresses it, after over a half a century of scale-upactivity in the chemical process industry, “such scale-up is still not anexact science but is rather a mix of physics, mathematics, witchcraft,history and common sense which we call engineering.” (Matsen, 1995.)

A complete treatment of scale-up should include the models,numerical calculation procedures and experimental data designers need tocarry out successful scale-up from small size beds to commercial units.This would involve a large measure of the existing fluidized bed researchand development effort; clearly, such a task is beyond the scope of a singlechapter Since changes in the bed size primarily influence scale-upthrough changes in the bed hydrodynamics, one focus of this chapter is onexperimental results and models which deal explicitly with the influence

of bed diameter on hydrodynamic performance for both bubbling andcirculating fluidized beds The changes in the bed dynamics will, in turn,impact the overall chemical conversion or combustion efficiency throughchanges in the particle-to-gas mass transfer and the heat transfer from thebed to immersed surfaces or the bed wall Several examples of thisinfluence are also reviewed

The second focus of this chapter is on the use of small scaleexperimental models which permit the direct simulation of the hydrody-namics of a hot, possibly pressurized, pilot plant or commercial bed Byuse of this modeling technique, beds of different diameters, as well asdifferent geometries and operating conditions, can be simulated in thelaboratory To date, this technique has been successfully applied tofluidized bed combustors and gasifiers Derivation of the scale modelingrules is presented for a variety of situations for gas solid fluidized beds.Verification experiments and comparisons to large scale commercialsystems are shown Rules for the use of this experimental modelingtechnique for FCC operations as well as for the simulation of bed-to-solidsurface heat transfer are also given

In this section, representative results are reviewed to provide aprospective of reactor modeling techniques which deal with bed size.There probably is additional unpublished proprietary material in this area.Early studies of fluidized reactors recognized the influence of bed diam-eter on conversion due to less efficient gas-solid contacting Experimentalstudies were used to predict reactor performance Frye et al (1958) used

a substitute reaction of ozone decomposition to study hydrocarbon

synthe-sis The ozone decomposition can be run at low pressures and

tempera-tures and can be rate-controlled in the same way and by the same catalyst

as the reaction under development Frye and coworkers used three beds of

2 inch, 8 inch and 30 inch diameter, respectively, to study the size

influence We should interject a caution that the use of pressures and

temperatures different than the actual reaction may mean that the

hydro-dynamics of the substitute reaction model will differ from the actual

application; this is illustrated later in the chapter Figure 1 shows the

apparent reaction rate constant for the different bed diameters at two

different bed heights with the other parameters held constant Note that

the rate constant decreased by roughly a factor of three between the 2 inch

and 30 inch beds

Figure 1 Apparent reaction-rate constant vs reactor diameter and bed height (From

May (1959) reports results of tests done in cold models used tosimulate the flow through large reactors whose performance had beenfound to be inferior to that of smaller pilot units The importance of thisproblem can be appreciated from the scale of the equipment used Figure

2 shows the 5 foot diameter unit used for the scale-up tests This unit wasfluidized with compressed air at 27 to 38°C (80 to 100°F) and pressures up

to 689 KPa (100 psi) Gas residence time in the bed was determined by theuse of tracer gas Radioactive solid tracers were introduced into the bed todetermine solid mixing The data obtained in the larger units are muchmore erratic with evidence of large scale mixing patterns Figure 3 showsthe axial mixing coefficients obtained in experiments with different sizebeds Mixing in the larger diameter bed is an order of magnitude largerthan that in a small laboratory unit The measured hydrodynamic behavior

of the gas and solid was combined with a reaction model to predict thereactor behavior Here again, there should be concern about the accuracywith the air experiments done at ambient temperature Use of identicalbed geometry and bed solid material does not guarantee identical hydro-dynamics The shift in gas properties from the cold model to the hotreactor may cause a marked difference in behavior Additional scalingparameters must be maintained constant between the reactor and the coldmodel to insure identical hydrodynamics, and in some cases just toguarantee identical flow regimes!

Volk et al (1962) show the effect of bed diameter on the sion of CO in the “Hydrocol” reaction in which hydrogen and carbondioxide are converted over an iron catalyst to hydrocarbons and oxygen-ated hydrocarbons in a bubbling or possibly slugging bed Figure 4 showsthe CO conversion It is seen that the conversion rate is reduced as thereactor diameter increases Volk used vertical tubes within the reactor toreduce the equivalent diameter of the system, equal to the hydraulicdiameter, four times the free cross sectional area divided by the wettedperimeters of all surfaces in the cross section The performance was found

conver-to be correlated by the equivalent diameter It was also found that bedexpansion was correlated with bed diameter In their process, larger bedswere built with internals which kept the equivalent diameter the same asthat of smaller units The large units with internals appeared to givecomparable gas-to-solid contacting The use of vertical internals may not

be feasible for a number of reasons, such as tube erosion The use of theequivalent diameter approach may not be universally valid

Figure 3 Solid diffusivity in axial direction for large units (From May, 1959.)

12 'r'! 5 in i O

Key Symbol Description

Van Swaaij and Zuiderweg (1972) used the ozone decomposition

reaction to study the conversion characteristics in a bubbling bed Studies

were made with beds of 5, 10, 23, 30 and 60 cm diameter and up to 300 cm

bed heights The results were compared with predictions using a

two-phase flow model with the mass transfer coefficient between the bubble

and dense phase derived from residence time distribution results of

gas-tracer pulse response tests Figure 5 shows the height of the mass transfer

unit H α , which is equivalent to u/ α where α is the mass transfer

coeffi-cient, as a function of the bed diameter The results from the ozone

conversion and the residence time distribution interpreted by the two

phase model gave reasonably similar results In these cases, the mass

transfer between phases is the limiting resistance for the reaction Note

that for larger bed diameters the mass transfer coefficient decreases Van

Swaaij and Zuiderweg (1973) showed that the inclusion of vertical tubes

in a bed gave bubble to dense phase mass transfer results which were

roughly equivalent to a smaller open bed with the same hydraulic diameter

while the solids axial mixing was higher than that predicted using the

hydraulic diameter

Figure 5 Mass transfer unit for ozone conversion for different bed diameters.

(From Van Swaaij and Zuiderweg, 1972.)

6 FROM CONVERSION DATA O C o FR)M RTD TESTS ( PRESENT INVESTIGATION )

u

«{H

D,m

Bauer et al (1981) measured the influence of bed diameter on thecatalytic decomposition of ozone Figure 6 shows the decrease of theconversion with bed diameter for Bauer’s data This figure also shows theinfluence of distributor design on conversion In many small scaleexperiments, a porous plate is used which will give better performancethan the distributors used in large shallow bed commercial designs.

Avidan and Edwards (1986) successfully scaled up from benchscale to demonstration plant from 0.04 m to 0.6 m diameter while main-taining nearly 100% conversion for a fluid bed methanol to gasolineprocess In this case, they ran at a superficial gas velocity which was highenough to be in the turbulent flow regime suppressing bubbles By thistechnique they eliminated the losses associated with gas bypassing inbubbles

Figure 6 Conversion catalytic decomposition of ozone for different bed

diam-eters and distributors (From Werther, 1992.)

3.0 INFLUENCE OF BED DIAMETER ON

HYDRODYNAMICS 3.1 Bubbling Beds

In the studies mentioned above, the major objective was theexperimental determination of conversion as a function of overall design

parameters and particle properties There have also been studies which

have looked at the influence of bed diameter on the hydrodynamics in an

effort to understand the cause of the conversion loss with bed size

increase These studies have aided in the development of physical

models of reactor performance

De Groot (1967) measured gas residence time, bed expansion and

solid axial mixing in a series of beds at different diameter fluidized with

air at ambient conditions He used a narrow size range and broad size

range of crushed silica with sizes below 250 µm Beds with diameters of

0.1, 0.15, 0.3, 0.6, and 1.5 m were used in the tests There was a substantial

decrease in bed expansion and bubble fraction for narrow size range

particles at large bed diameters, indicating the possibility of gas bypassing

in bubble channels (Fig 7) The axial diffusivity also increased with bed

diameter and was a strong function of particle size distribution (Fig 8)

!0

Weather (1974) measured the bubble characteristics in cylindrical

beds of diameters 100, 200, 450 and 1000 mm, respectively, for fine

particles with a mean particle diameter of 83 µm He showed that for beds

200 mm or smaller, common sizes used for laboratory experiments, the

bed diameter had a strong effect on the bed hydrodynamics There was a

zone of preferred bubble flow near the bed walls at lower elevations (Fig

9) The bubbles grew in size and moved toward the centerline,

presum-ably by coalescence, higher up in the bed (Fig 10) The transition to

slugging occurred higher up in the larger bed at the same superficial gas

velocity The bubble velocity increased with height until slug flow was

approached, after which the velocity decreases For the only case shown,

the 100 mm bed, the maximum velocity occurred when the bubble

dimen-sion was about one third of the bed diameter In larger beds, the bubble

rise velocity was higher for the same bubble volume (Fig 11) Hovmand

and Davidson (1971) reviewed data on bubble rise velocity and concluded

that the bubble rise velocity is governed by the bed diameter rather than

the bubble diameter when the bubble diameter exceeded 1/3 of the bed

diameter Note that Weather’s results at different superficial velocity are

well correlated by the drift flux form,

Eq (1) u b~Φ( u o - u mf ) + c gd v

Figure 9 Bubble gas flow V b as a function of the distance r from the vessel center

line in a height of 30 cm above the distributor in beds of different diameters D (u =

Figure 10 The flow pattern of the bubble phase in a bed of 200 mm diameter

(u o = 9 cm/sec, H = 50 cm) (From Werther, 1974.)

,

<.J QJ

~(I)I5

~

5 , ~

,

~ 0

\::

(I) 0 0)

~ Q Q :3 Q - 0 C.,J

Figure 11 Relationship between local mean bubble size and local mean bubble

rise velocity in beds of different diameters (u o = 9 cm/sec) (From Werther, 1974.)

It is curious that the influence of bed size appeared to hold, even

when the bubble is much smaller than the bed diameter This may be tied

to the local concentration of bubbles within certain sections of the bed

The increased local bubble flow led to higher coalescence rates and higher

local bubble velocities There was a distinction in bubble velocity

be-tween Geldart group A and B powder while the ratio of visible bubble

flow to u-u mf seems to be independent of the group Werther also found

that the visible bubble flow, the product of the number of bubbles per unit

time crossing a given surface and their respective volumes was

consider-ably less than u-u mf especially in the lower regions of the smaller bed and

throughout large diameter beds The residence time of bubbles was

significantly higher for beds 200 mm or smaller, than it was for larger

diameter beds Werther concludes that the smallest bed diameter that

appears suitable for obtaining good scale-up results is 500 mm It should

be pointed out that this criterion will probably vary with flow conditions,

diameter of a sphereof volume equal

to local mean bubbe volume I dv I cm

diameter of a sphereof volume equal

to local mean bubbe volume I dv I cm

10 I

diameter of a sphereof volume equal

to local mean bubbe volume I dv I cm

bed depth and particle size Also, Werther’s experiments were carried outwith air at ambient conditions Although the trends and physical picturemay be similar for beds at elevated temperature and pressure, there may besome changes in the criteria.

Whitehead (1978) found patterns of bubble tracks in a large 1.2 msquare bed similar to Werther: preferred bubble tracks near the walls andcorners of a shallow open bed and merging of bubbles toward the bedcenter at higher elevations Nguyen, Potter, and Whitehead (1979) alsofound that a horizontal tube bank in the large bed caused smaller bubbleswhich appeared in more random locations across the upper surface of thebed This work was carried out for fine solids at low superficial velocity,

15 cm/sec, and modest bed depths

Geldart (1970) showed a substantial distinction between bubblesizes in two dimensional and three dimensional beds He used 128 µmriver sand in a 30.8 cm round bed and a 68 × 1.27 cm rectangular crosssection bed The results, shown in Fig 12, show that the bubbles in thethree dimensional bed are larger There were differences in the visiblebubble flow rate at the same superficial velocity Geldart ascribes thedifferences in bubble diameter to differences in visible bubble flow rate aswell as to out-of-line coalescence in the three dimensional bed

Figure 12 Variation of bubble diameter and concentration with initial bed

height; u - u o = 6 cm/sec for two dimensional and three dimensional beds (From Geldart, 1970/71.)

Ho Ccm)

Glicksman and McAndrews (1985) determined the effect of bed

width on the hydrodynamics of large particle bubbling beds Sand

par-ticles with a mean diameter of 1 mm were fluidized by air at ambient

conditions The bed width ranged from 7.6 cm to 122 cm while the other

cross sectional dimension remained constant at 122 cm Most

experi-ments were carried out with an open bed The bubble rise velocity

increased with the bed width, in the representation of bubble velocity as

the value of φ varied from 0.4 in a two-dimensional bed to 0.6 in the three

dimensional bed (Fig 13) The mean vertical chord length of bubbles

decreased with bed width (Fig 14) The visible bubble flow decreased

dramatically with an increase of bed width at a fixed superficial velocity

(Fig 15) The gas throughflow coefficient m, used to represent the total

gas flow,

Eq (3) u o = Q b + (1 -δ )u mf + mδ u mf

increased from a mean value of 3.6 for the two dimensional bed to a mean

value of 11.7 for the largest, 1.22 × 1.22 m bed cross section The bed

depths at minimum fluidization were 46 cm and 76 cm in the tests For

these rather shallow beds there were no observable preferred bubble

tracks; the location of erupting bubbles was random across the bed

sur-face For the cases observed, with u o u mf varying from 1.3 to 1.8, the

influence of the wall was absent when the bubble diameter was

roughly 1/5 the bed width or less For deeper beds or higher gas velocity,

the ratio of bubble diameter to bed width is expected to be the best

criterion for determining when wall effects will be negligible For one test

series, five staggered rows of horizontal tubes with a horizontal

center-to-center spacing of 15.2 cm were placed in the 1.22 m square bed The tube

bed results are shown in Figs 13–15 with a T symbol The behavior of

bubbles in the bed with the tube bank resembled that of an open bed with

smaller width

Figure 13 Variation of bubble rise velocity with wall spacing; 1 standard

: tjJ = Ub- (U -Um,J/(gD~J/2 (From Glicksman and McAndrews, 1985.)

WALL SPACING Dr CH

Figure 15 Visible bubble flow Qb.lmf=76 cm, probeheight=46 cm x= 1.3 Umf;

O = 1.5Umf; V = 1.8Umf; T denotes with tubes Experiment uncertainty = 20% (From Glicksman and McAndrews, 1985.)

3.2 Mixing

Van Deemter (1980) surveyed data on solid mixing in fluidized

The differences in behavior between small laboratory beds and

larger demonstration units can, in part, be attributed to a switch from

porous plate distributors in the small bed to discrete hole or bubble caps in

Figure 16 Mixing coefficients for different vessel diameters; M L: longitudinal

dispersion of fluid, M B : gas backmixing, M S : longitudinal solid dispersion (From

§e

i0000

&

O O 8

-00' o 0 O 0 O

Figure 17 The effect of bed diameter on solid mixing (From Yerushalmi &

Avidan, 1985.)

the larger beds The porous plates give a better quality of fluidization, e.g.,smaller bubbles, for shallow beds and beds of moderate depth (Rowe andStapleton, 1961)

Yerushalmi and Avidan (1985) suggest that the axial dispersioncoefficient of solids in slugging and turbulent flow varies approximatelylinearly with the bed diameter, similar to Thiel and Potter (1978) Thedata are shown in Fig 17 although May’s results are probably in thebubbling fluidization regime rather than turbulent flow

3.3 Influence of Bed Diameter on Circulating Fluidized Beds

Arena et al (1988) measured the hydrodynamic behavior of twocirculating fluidized beds (CFB) with riser diameters of 0.041 m and 0.12

m ID, respectively, of roughly the same height At the same superficialgas velocity and solid recirculation rate, the larger diameter column had ahigher solids fraction The average slip velocities derived from this dataare also higher for the large diameter riser (see Hartge et al 1985).Yerushalmi and Avidan (1985) found a similar trend when comparing

15.2 cm and 7.6 cm columns Noymer et al (1995) also compared two

columns of 5.08 cm and 7.68 cm diameter of the same height which were

used to simulate larger pressurized fluidized bed combustors They found

higher solids loading for the larger diameter riser at equal gas velocity and

solid recirculation In addition, the fraction of the wall covered by clusters

was higher for the larger diameter column when the two beds had equal

solids flow and when the two beds had equal cross section averaged solids

concentration

Rhodes et al (1992) compared the solids flux profiles across the

cross sections of a 0.152 m and 0.305 m diameter circulating bed riser

They found a region where the solid profile, given by the ratio of local flux

to average flux, had a similar variation over the cross section, which was

insensitive to the level of solid flux The variation of the local solids flux

over the radius was a function of the gas velocity and the riser diameter In

the larger riser, the profiles were somewhat flatter and the thickness of the

downflowing region relative to the bed radius was smaller The

compari-sons were not exact since the cross sections compared for the two beds

were at different heights

Zhang, Tung and Johnsson (1991) carried out investigations with

three different fast bed systems with diameters of 32, 90 and 300 mm,

respectively They found that the radial voidage distribution, as a ratio of

the cross sectional average, was independent of bed diameter and solids

recycle rate (Fig 18) The similarity does not hold at transition to the

turbulent regime The results are for the center of the riser excluding the

entrance and exit regions It would be interesting to determine if the

similar voidage profiles hold for larger diameter risers

Figure 18 Comparison of radial voidage profiles calculated by correlation with

experimental data in the three different beds used (FCC/air) (From Zhang et al.,

The thickness of the downflowing layers at the wall of the CFB istypically defined as the distance from the wall to the position of zerovertical solid flux Measurements of the layer thickness were made on a

12 MW and 165 MW CFB boiler by Zhang, Johnsson and Leckner (1995).They found that the thickness increased for the larger bed They relateddata from many different beds (Fig 19), with the equivalent bed diameter,taken as the hydraulic diameter, using the following form

Eq (4) δ = 0.05De0.74

Figure 19 Empirical correlation and experimental data of thickness of

downflowing layer at the wall of a CFB as a function of the equivalent bed

diameter (From Zhang et al., 1995.)

The thickness, δ, was found to be insensitive to particle

concen-tration, gas velocity and height within the furnace That suggests that thethickness results from a balance of solids internal circulation which isgenerally much higher than net throughflow If the local solids flux profile,

.0

*

~08).x0

-ar - cc al 1969 0a;dc8 & Bi8t 1971

~ccaL.l986

88118 -ai., 1988 RJxx- I m W-,, al.I991 c-.i8-aL.1991 ZI-.&~1991 W - 1992

12 MW CfB dIiI-W.

165 MW CfB d8

-*-I '() (/) (/) w z

~

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~ 0.1