EPA-625/7-76-001
ENVIRONMENTAL POLLUTION CONTROL
PULP AND PAPER INDUSTRY
PARTI
AIR
U.S. ENVIRONMENTAL PROTECTION AGENCY
Technology Transfer
October 1976
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ACKNOWLEDGMENTS
This process design manual was prepared for the Office of Technology Transfer of the U.S.
Environmental Protection Agency. Coordination and preparation of Part I of this manual,
Air Pollution Control, was carried out by EKONO, inc., Bellevue, Washington, and EKONO,
Oy, Helsinki, Finland.
• . . U.S. EPA reviewers were James C. Herlihv of the U.S. EPA Division of Stationary
Source Enforcement, Washington, D.C., James A. Eddinger of the U.S. EPA Emission Stand-
ards and Engineering Division, Research Triangle Park, N.C., Gene Tucker of the U.S. EPA
Industrial Environmental Research Laboratory, Research Triangle Park, N.C., and George
S. Thompson, Jr., of the U.S. EPA Office of Technology Transfer, Cincinnati, Ohio.
NOTICE
The mention of trade names of commercial products in this publication is for illustration
purposes and does not constitute endorsement or recommendation for usc by the U.S.
Environmental Protection Agency. This manual is presented as a helpful guide to the user
and should in no way be construed as a regulatory document.
ii
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PART I: AIR
CONTENTS
Chapter Page
ACKNOWLEDGMENTS ii
CONTENTS iii
LIST OF FIGURES ix
LIST OF TABLES xv
FOREWORD xix
INTRODUCTION 1-1
1.1 Gaseous and Particulate Emissions from Kraft Pulp and Paper Mill
Process Sources 1-2
1.2 Gaseous and Particulate Emissions from Sulfite Pulp and Paper
Mill Process Sources 1-9
1.3 Power Boilers 1-10
1 .4 References 1-13
2 DIGESTER GASES 2-I
2.1 Batch Digesters 2-I
2.2 Continuous Digester Gases 2-li
2.3 References 2-14
3 EVAPORATION GASES 3-1
3.1 Black Liquor Properties 3-1
3.2 Evaporator Types 3-2
3.3 Evaporator Gas Scrubbing 3-8
3.4 General Evaporator Air Pollution Abatement Programs 3-] i
3.5 In-plant Controls 3-] 3
3.6 References 314
‘U
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CONTENTS — Continued
Chapter Page
4 NONCONDENSABLE GAS TREATh’IENT 4-1
4.1 Gas Stream Characteristics 4-1
4.2 Gas Handling Systems 4-6
4.3 Gas Treatment Systems 4-14
4.4 System Economics 4-19
4.5 References 4-20
5 CONDENSATE TREATMENT 5-1
5.1 Condensate Components 5-1
5.2 Digester Condensates 5-2
5.3 Evaporator Condensates 5-4
5.4 Condensate Chlorination 5-9
55 Condensate Stripping 5-9
5.6 References 5-20
6 BROWN STOCK WASHER GASES 6-1
6.1 Displacement Washing 6-1
6.2 Diffusion Washers 6-2
6.3 References 6-5
7 STORAGE TANK VENT GASES 7-1
7.1 Storage Tank Vent Gas Composition 7-]
7.2 Storage Tank Vent Gas Treatment 7-1
7.3 References 7-1
8 TALL OIL RECOVERY GASES 8-1
8.1 Batch Tall Oil Recovery 8-1
8.2 Continuous Tall Oil Recovery 8-2
8.3 References 8-3
iv
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CONTENTS — Continued
Chapter Page
9 BLACK LIQUOR OXIDATION 9-I
9.1 Wcak Black Liquor Oxidation—Air 9-I
9.2 Strong Black Liquor Oxidation—Air 9-I I
9.3 Agitated AirSparging 9-16
9.4 Combination Systems 9-16
9.5 Molecular Oxygen Systems 9-lB
9.6 Process Effects 9-27
9.7 Air Pollution Effects 9-30
9.8 Oxidation Tower Emissions 9-33
9.9 Process Economics 9-34
9.10 References 9-43
10 RECOVERY BOILER DESiGN AND OPERATION 10-I
10.1 General Conditions [ 0-I
10.2 Combustion of Black Liquor Dry Solids 10-15
10.3 Different Recovery Boiler Designs 10-28
10.4 Process Variables 10-40
10.5 Diverse Obnoxious Compounds 10-51
10.6 Direct Contact Evaporation 10-5!
10.7 Flue Gas Scrubbing for Gaseous Emissions 10-56
10.8 Collection of Particulate Matter from Recovery Boiler Flue Gas 10-61
10.9 Economy of Recovery Boiler Operation 10-74
10.10 References 10-77
Appendix 10-1 10-81
Appendix 10-2 10-82
11 LIME BURNING AND LIME DUST HANDLING 11-1
11.1 Rotary Lime Kilns 11-1
11.2 Fluid izcd Bed Calciners 11-2
11.3 Particulate Emission Control 11-3
11.4 Gaseous Emission Control 11-6
11.5 Oxygen Addition 11-9
11.6 Process Economics 11-9
11.7 Lime Dust 1-landling 11-10
1]..B References 11-11
v
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CONTENTS — Continued
Chapter Page
12 SMELT DISSOLVING TANK 12-i
12.1 Smelt Dissolving Tank Particulate Matter Emissions 12-1
12.2 Smelt Dissolving Tank IRS Emissions 12-2
12.3 References 12-4
13 EMISSIONS OF OXIDES OF NITROGEN, HYDROCARBONS, AND
WATER VAPOR 13-1
13.1 Nitrogen Oxides 13-1
13.2 Water Vapor [ 3-4
13.3 Organic Compounds 13-6
13.4 References i3-6
14 AIR POLLUTiON CONTROL IN SULFITE PULP MILLS 14-1
14.1 Sulfite Pulping Processes 14-1
14.2 Digester Gases 14-4
14.3 Washer Gases 14-7
14.4 Evaporator Gases 14-7
14.5 Combustion Gases 14-JO
14.6 Acid Preparation Gases 14-12
14.7 SSL Recovery Boilers 14-16
14.8 SSL Recovery Systems 14-20
14.9 Problem of Nitrogen Compounds for Ammonium-Bascd Pulping 14-27
14. LO References 14-30
15 OTHER PROCESS SOURCES 15-i
15.1 Bleach Plant Gases 15-1
15.2 Wastcwater Treatment 15-3
15.3 Odor Problems from l)iffusc Sources 15-6
15.4 References 15-8
16 PO VER BOILERS 16-1
16.1 Supply Patterns 16-1
L6.2 Combustion Parameters 16-2
vi
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CONTENTS — Continued
Chapter Pap
16.3 Boiler Types 16.6
164 Particulate Emissions 16 .14
16.5 ltdc e , ,s 1620
17 PROCESS MONITORING 17 - 1
17.1 Source Measurements 17-I
17.2 Gaseous Monitoring 17-I
17.3 Particulate Monitoring 17.24
174 Odor Meaeuremcnts 17-31
17.5 Mobile Laboratories 17-35
17.6 Economics 17-39
17.7 References 1741
APPENDIX A — GLOSSARY OF SYMBOLS A-i
APPENDiX B — CHEMICAL FORMULAS B-I
V I I
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LIST OF FIGURES
Figure No. Page
2.1 Batch Digester Flow Shcet 2-2
2-2 Kraft Batch Digester Blow Steam Flow 2-3
2-3 Kraft Batch Digester Blow Gas Flow After Condensing and Without
Equalization 2-4
2-4 Vapor Pressure of 0.01 M H 2 S vs. pH at Various Temperatures 2.5
2-5 Vapor Pressure of OM 11V CH 3 SH vs. pH at Various Temperatures 2.6
2-6 Blow I-teat Recovery Control System 2-7
2-7 Odor Compounds in Relief Gas After Turpentine Condenser as a
Function of Condenser Outlet Temperature 2.8
2-8 Vapor Pressures of Methanol, Water, a-Pinene & -Pinene 2-12
2-9 Continuous Digester Flow Sheet 2-13
3-1 Vapor-Liquor Equilibrium for FI 2 S Over Black Liquor 3-2
3-2 Multi-Effect Vacuum Evaporation Plant Flow Sheet 3-3
3-3 Evaporation Plant Direct Contact Condenser With Water Ring Pump 3-5
34 Evaporation Plant Two Stage Surface Condenser With Water Ring
Pump 3-7
3-5 Multiple Effect Back Pressure Evaporation Plant Flow Sheet 3-9
3-6 Flash Evaporation Plant Flow Sheet 3-JO
3-7 Multiple Effect Single Stage Thermocomprcssor Evaporation Plant
Flow Sheet 3-11
3-8 Hotwell Gas Scrubber for 100 Metric Tons Per Hour (40 gpm) Evapo-
ration Plant for H 2 S—Separation of 95% or More 3-12
4-1 Vaporsphcre Flow Equalization Gas Holders 4-7
4-2 Floating Cover Flow Equalization Gas Holders 4-8
4-3 Packed Bed Scrubber For Noncondensable Gas Handling System 4-11
4-4 Liquid Condensate Trap for Noncondensable Gas Handling System 4-13
4-5 Safety Devices for Noncondensable Gas Handling Systems 4-14
4-6 Unsteady State Flow System for Batch Digester Noncondensable Gas
Incineration 4-15
4-7 Steady State Flow System for Continuous Digester Noncondensable
Gas Incineration 4-16
4-8 Noncondensable Gas Incineration System 4-17
5-1 Evaporation Plant Surface Condenser With Water Jet Condenser and
Water Ring Pump 55
ix
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LIST OF FIGURES — Continued
Figure No. Page
5.2 Evaporation Plant Surface Condenser With Recirculated Water Jet
Condenser and Water Ring Pump 5-6
5.3 Contaminated Condensates Air Stripping Plant Flow Sheet 5-1 1
5.4 Contaminated Condensates Steam Stripping Plant Flow Sheet 5.1 3
5-5 Stripping Efficiency for Different Steam-Condensate Ratios With 10
Theoretical Plates 5-15
5-6 Condensate Stripping in an Evaporation Plant 5-16
5-7 Simplified Flow Sheet for Kraft Process With Condensate Segregation,
Stripping, & Reuse 5.19
6-1 Vacuum Washers Flow Sheet 6-2
6-2 Pressure Washers Flow Sheet 6-3
6.3 Batch Diffusion Washers Flow Sheet 6-4
6.4 Continuous Diffusion Washers Flow Sheet 6.5
8-1 Batch Tall Oil Plant Flow Sheet 8-1
8-2 Continuous Tall Oil Plant Flow Sheet 8.3
9-1 Collins Porous Plate Diffuser Weak Black Liquor Oxidation System 9.2
9-2 Trobeck-Ahlen Multiple Sieve Tray Weak Black Liquor Oxidation
System 9.5
9-3 Packed Tower Systems for Weak Black Liquor Oxidation 9-8
9-4 Agitated Air Sparging System for Black Liquor Oxidation 9-10
9-5 Champion Unagitatcd Air Sparge Strong Black Liquor Oxidation
System 9-13
9.6 Operating & Performance Data for Single Stage Strong Black Liquor
Oxidation System 9-14
9-7 Champion Two Stage Unagitated Strong Black Liquor Oxidation
System 9-16
9-8 Western Kraft Pipeline Reactor Strong Black Liquor Oxidation System 9-17
9-9 Two Stage Combination Weak & Strong Black Liquor Oxidation With
Oxygen 9-20
9-10 Owens-illinois System for Two Stage Weak & Strong Black Liquor
Oxidation With Oxygen 9-23
9-LI Effect of Production Rate on Capital & Operating Costs for Weak &
Strong l3lack Liquor Oxidation With Oxygen 9-36
9-12 Operating Costs for Weak Black Liquor Oxidation With Oxygen 9-39
x
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LIST OF FIGURES — Continued
Figure No. Page
10-I 1-teat Values vs. Oxygen Demand for CompleLe Combustion of Lignin
and Carbohydrates 10-4
10-2 Heat Values vs. Oxygen Demand for Complete Combustion for Some
North American Dry Solids 10-5
10-3 Flow Diagram for Black Liquor Through Recovery Boiler-North
American System ]0-8
10-4 Flow Diagram for Black Liquor Through Fluidized Bed Reactor With
Waste Heat Recovery Boiler ]0.9
10-5 Flow Diagram for Black Liquor Through Recovery Boiler Scancli-
navian System—Low Odor System 10-] 1
10-6 Principle Design of Air Distribution to Recovery Boiler Furnaces 10-15
10-7 Air Supply According to Gotaverken Angtcknik AB Design 10.17
10-8 Modern Kraft Recovery Unit from Babcock & Willcox 10-18
10-9 Modern Kraft Recovery Unit from Combustion Engineering 10-]9
10-10 Equilibrium Diagram for Condensed Phases and Gases for a Sodium
Based Black Liquor 10-21
10-1 1 Distribution of Sodium and Sulfur in a Black Liquor Recovery
Furnace as a Function of Temperature 10-23
10-12 Equilibrium Diagram for a Na 2 C0 3 -Na 2 S System 10-25
10-13 Effect of NaCI Addition on the Melting Point of a Synthetic Pulp [ 0-26
10-14 Babcock & Wilcox Recovery Boiler With Cyclone Evaporator 10-30
LO- [ 5 Cyclone Evaporator 10-31
10-16 Combustion Engineering Recovery Boiler With Cascade Evaporator 10-32
10-17 Open View of Cascade Evaporator 10-33
10-18 Combustion Engineering Recovery Boiler With Cascade Evaporator
Ace System 10-34
10-19 Combustion Engineering Recovery Boiler With Laminaire Air Heater
& Complete Multiple Effect Evaporation L.A.H. System 10-35
10-20 Recirculation Air Heater for Scandinavian Recovery Boiler 10-37
10-21 Typical Götaverken Angteknik Recovery Boiler 10-38
10-22 Heat Value vs. Oxygen Demand at Complete Combustion 10-45
10-23 Viscosity of Black Liquors 10-47
10-24 Adjustable Air Ports 10-50
10-25 British Columbia Research Council Design for H 2 S Absorption Scrub-
ber 10-58
10-26 Glitsch-Weyerhaeuscr Design for a TRS Scrubbing System 10-59
10-27 Formation of Visible Plume Through Condensation of Water Vapor 10-60
10-28 Electrostatic Precipitator Size as a Function of Collecting Efficiency 10-67
xi
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LIST OF FIGURES — Continued
Figure No. Page
10-29 Capital Cost for Recovery Boiler Department 10-75
10-30 Flue Gas Energy Losses 10-76
14-1 Principal Flow Diagram for Spent Liquor Collection in Blow Pits, 2
Stage Displacement Wash 14-6
14-2 Principal Flow Diagram for Spent Liquor Collection in Rotary Washer
Plant (3 2-Zone Filters) 14-8
14-3 Principal Flow Diagram for Spent Liquor Evaporation in Multiple
Effect Vacuum Plant, 5 Effect, 7 Bodies 14-9
14-4 Flow Sheet for Calcium Base Raw Acid Preparation 14-13
]4-5 Acid Bisulfite Fortification System Flow Sheet 14-14
14-6 Flow Sheet for Magnesium Base Raw Acid Preparation 14-15
14-7 Babcock & Wilcox Process for Magnesia Base Sulfite Chemicals
Recovery 14-2 1
14-8 STORA Process for Sodium Base Sulfite Chemicals Recovery 14-23
14-9 SCA Process for Sodium Base Sulfite Chemicals Recovery 14-25
14-10 Tampella Process for Sodium Base Sulfite Chemicals Recovery 14-26
15-1 Equilibrium Solubiity of Chlorine Dioxoide in Water 15-3
15-2 Equilibrium Solubility of Sulfur Dioxide in Water 154
17-1 Gas Sample Handling & Conditioning System for Externally Located
Continuous Gaseous Monitoring System 17-3
17-2 Continuous Source Monitoring System for Rcduced Sulfur Emissions
With Coulometric Titration 17-9
17-3 Total Reduced Sulfur Monitoring With an Electrochemical Membrane
Cell Detector 17-10
17-4 Continuous Conductivity Monitor for Measuring Sulfur Oxide Emis-
sions from Sulfite Mill Sources 17-12
17-5 Electrochemical Transducer Membrane Cell for Continuous Sulfur
Dioxide Monitoring 17-13
17-6 Internally Located Ultraviolet Spectrometer for Sulfur Dioxide Moni-
toring in Flue Gas Streams 17-14
17-7 lnternally Located Ultraviolet Spectrophotometer for Sulfur Dioxide
Monitoring in Flue Gas Streams 17-15
17-8 Rotating Syringe Instrument Calibration Procedure 17-22
17-9 Permeation Tube Instrument Calibration Procedure 17-23
1 7-10 ASME Batch Particulate Sampling Train 17-26
xii
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LIST OF FIGURES — Continued
Figure No. Page
17-11 EPA Batch Particulate Sampling Train 17-26
17-12 Multistage Cascade Impactor for Particle Size Distribution Determina-
tion 17-28
17-13 Membrane Filter System for Particle Size Distribution Determination 17-28
17-14 Continuous Monitoring of Particulate Emissions With a Conductivity
Cell Detector 17-29
17-15 Continuous Monitoring of Particulate Emissions With a Sodium Ion-
Specific Electrode 17-30
17-16 Swedish Dynamic Dilution System for Odor Level Evaluation 17-35
17-17 Scentometer Dilution System for Odor Threshold Evaluation 17-36
17-18 Gaseous Sampling System for Sulfur Gas Analysis in NCASI Mobile
Laboratory [ 7-37
17-19 Particulate and Gaseous Sample Handling Systems for ITT-Rayonier
Mobile Laboratory 17-39
XIII
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LIST OF TABLES
Table No. Page
1-1 External Control Techniques for Gaseous and Particulate Matter Emis-
sions From Kraft Pulp Mill Sources 1-3
1-2 Typical Gas Characteristics for Kraft Pulp Mill Processes 1-4
1-3 Typical Reduced Sulfur Gas Concentrations from Kraft Pulp Mill
Sources 1-5
14 Typical Reduced Sulfur Gas Emission Rates from Kraft Pulp Mill
Sources 1-6
1-5 Typical Concentrations and Rates for SOx and NO Emission from
Kraft Pulp Mill Combustion Sources 1-7
1-6 Typical Concentrations and Emission Rates for Particulate Matter
from Kraft Pulp Mill Sources (After Control Devices) 1-9
1-7 Typical SO 2 Emission Rates from Sulfite Pulp Mill Sources 1-10
1-8 Uncontrolled Air Pollutant Emissions from Fuel Combustion in
Auxiliary Power Boilers 1-12
3-1 Effect of Condenser Type on Reduced Sulfur Gas Emission from
Evaporator Noncondensable Gases 3-14
4-1 Gas Flow Rates from a Batch Digester 4.3
4-2 Typical Ranges in Digester and Evaporator Noncondensable Gas Flow
Rates 4-3
4-3 Flammability Limits in Air for Compounds Present in Kraft Noncon-
densable Gas Stream 4-5
4-4 Dilution Rcquircments With Air to Avoid Explosions for Digester
Noncondensable Gas Streams 4-5
4-5 Flame Propagation Speeds for Air-Mercaptan Mixtures 4-6
4-6 Dimensions of Flow Equalization Devices in Kraft Noncondensable
Gas Handling Systems 4-9
4-7 Batch Digester Blow Gas Flow Rates for Sizing Noncondensable Gas
Flow Equalization Devices 4-10
4-8 Piping Systems for Kraft Noncondensable Gas Handling Systems 4-12
4-9 Gas Flow Rates to Burning Devices from Noncondensable Gas Han-
dling Systems 4-16
4-10 Capital Costs for Installation of Noncondensable Gas Handling
Systems 4-20
5-1 Main Components of Typical Kraft Mill Condensates 5-2
5-2 Typical Kraft Mill Condensate Compositions, Mean Values for 10 Mills 5-3
xv
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LIST OF TABLES — Continued
Table No. Page
5.3 Typical Kraft Mill Condensate Characteristics for 17 Mills 5-4
54 Calculated Evaporator Condensate Flow, Sulfur, Methanol, and BOD
Distribution for Liquor Sequence 3-4.5.2.1 5-7
5.5 Chlorine Demand of Reduced Sulfur Compounds for Oxidation to
Sulfur 5-9
6-1 Flows, Composition and Sulfur Release of Vacuum Washer Foam
Tank Vent Gases 6-3
8-1 Bath Tall Oil Recovery Plant TRS Components and Typical Concen-
trations 8-2
8-2 Tall Oil Recovery Noncondensable Gas Flows and Reduced Sulfur
Emissions 8-2
9-1 Design and Operating Parameters for Porous Plate Diffuser Black
Liquor Oxidation Systems 9-3
9-2 Design and Operating Parameters for Trobeck-Ahlen Multiple Sieve
Tray Weak Black Liquor Oxidation Units 9-6
9-3 Design and Operating Parameters for Weyerhaeuser Concurrent Flow
Packed Tower Weak Black Liquor Oxidation System 9-9
94 Operating and Performance Data for Agitated Air Sparged Weak Black
Liquid Oxidation Systems 9-11
9-5 Design Criteria of Plug Flow Reactor Systems for Black Liquor Oxida-
tion With Molecular Oxygen 9-19
9-6 Oxygen Requirements for Weak Black Liquor Oxidation 9-21
9-7 Effect of Weak Black Liquor Oxidation on Liquid Chemical Composi-
tion 9-24
9-8 Effect of Weak Black Liquor Oxidation on Sulfur Gas Emissions from
Kraft Mill Process Sources 9-25
9-9 Effect of Black Liquor Oxidation on Sulfur Gas Emissions During
Direct Contact Evaporation 9-31
9-10 Effect of Weak Black Liquor Oxidation on Malodorous Sulfur Gas
Emissions from Evaporator Noncondensable Gases 9-32
9-11 Reduced Sulfur Emissions from Black Liquor Oxidation Tower Vents
Using Air 9-34
9-12 Estimated Capital Costs for Black Liquor Oxidation Systems 9-35
9-13 Approximate Annual Operating Costs for Black Liquor Oxidation Sys-
tems Using Air 9-37
xvi
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LIST OF TABLES — Continued
Table No. Page
9-14 Effect of Number of Stages on Annual Operating Costs for Strong
Black Liquor Oxidation With Air 937
9-15 Estimated Capital Costs for Black Liquor Oxidation Systems Using
Molecular Oxygen 9-40
9-16 Operating Costs and Operating Credits for Black Liquor Oxidation
With Oxygen 9 -41
9-17 Typical Ranges in Operating Variables, Reduced Sulfur Emissions and
Cost Factors for Black Liquor Oxidation Systems 942
10-1 Black Liquor Dry Solids Content 10-2
10-2 Black Liquor Combustion Products in Caloric Bomb Using Modified
Technique 10-6
10-3 Black Liquor Combustion Products in Recovery Furnace 10-6
104 Operating Conditions for American and Scandinavian Black Liquor
Concentration 10-12
10-5 Capital Cost Comparison of American and Scandinavian Black Liquor
Concentration Practices 10-13
10-6 Annual Operating Cost Comparison of American and Scandinavian
Black Liquor Concentration Practices 10-13
10-7 Recovery Furnace Exhaust Gas Properties for Direct and Indirect Con-
tact Evaporation Systems 10-52
10-8 Effect of Black Liquor Oxidation on Sulfur Gas Emissions During
Direct Contact Evaporation 10-54
10-9 Effect of Black Liquor pH on 1-12 S Emissions During Direct Contact
Evaporation 10-55
10-10 Average Particulate Emissions from Recovery Boiler Electrostatic
Precipitators in the United States 10-71
11-1 Energy Requirements for Lime Mud Calcining Systems 11-2
11-2 Operating Characteristics for Particulate Liquid Scrubbers Employed
on Kraft Lime Kilns 11-4
11-3 Particulate Collection Efficiencies for Liquid Scrubbers on Kraft Pulp
Mill Lime Kilns 11-5
11-4 Gaseous Emissions from Kraft Pulp Mill Lime Kilns 11-6
11-5 Capital and Operating Costs for Lime Kiln Particulate Scrubbers 11-10
12-1 Smelt Dissolving Tank Particulate Matter Control Devices 12-2
12-2 TRS Emissions from Smelt Dissolving Tanks 12-3
xvii
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LIST OF TABLES — Continued
Table No. Page
13-1 Effect of Flame Temperature on Nitric Oxide Equilibrium Concentra-
tion and Reaction Time 13-2
13-2 Nitrogen Oxide Emissions from Kraft Pulp Mill Process Sources 13-3
13-3 Nitrogen Oxide Emissions from Power Boilers 13-4
134 Water Vapor Emissions to the Atmosphere from Kraft Pulp Mill
Sources 13-5
14-] Main Sulfite Pulping Processes 14 -2
14-2 Sulfite Process Chemicals, Price, Combustion and Recovery 14-2
14-3 Scandinavian Sulfite Pulp Mill Emissions 14-3
144 Typical Potential American Sulfite Pulp Mill Emissions 14-4
14-5 Nitrogen Oxides Emissions from an Ammonium Base Sulfite Recovery
Furnace 14-29
15-1 Odorous Gases from Diffuse Kraft Pulp Mill Sources ] 5-7
16-1 Overall National Distribution of Energy Sources for the Pulp and
Paper Industry 16-1
16-2 Characteristics of Fuel Burned in Power Boilers at Pulp and Paper
Mills in the United States 16-3
16-3 Uncontrolled Air Pollutant Emission Factors for Fuel Combustion in
Power Boilers for the Pulp and Paper industry 16-4
164 Particulate Emission Characteristics from Selected U.S. Power Boilers ]6-15
16-5 Typical Particle Size Distribution of Fly Ash from Coal- and Wood-
Fired Power Boilers 16 -16
17-1 Selective Preserubbing Solutions for Sulfur Gas Separation 17 -6
17-2 Approximate Ranges in Calibration Factors for Sulfur Gases With
Coulometric Titrator 17-8
1 7-3 Operating Characteristics of Gas Chromatographic Detectors 17-21
1.74 Odor Threshold Levels for Malodorous Sulfur Compounds 17-32
1 7-5 Odor Intensity Level Evaluation Scale 17-33
17-6 Approximate Capital Costs for Continuous Gaseous Stack Monitoring
Instrumentation 1 7-40
1 7-7 Approximate Capital Costs for Continuous Particulate Stack Monitor-
ing Instrumentation 17-40
xviii
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FOREWORD
The formation of the United States Environmental Protection Agency marks a new era of
environmental awareness in America. The Agency’s goals are national in scope and encom-
pass broad responsibility in the area of air and water pollution, solid wastes, pesticides, and
radiation. A vital part of EPA’s national pollution control effort is the constant develop.
ment and dissemination of new technology.
It is now clear that only the most effective design and operation of air, water, and solids
waste control facilities, using the latest available techniques, will be adequate to meet the
future air and water quality objectives and to ensure continued protection of the nation’s
environment. It is essential that this new technology be incorporated into the contempo-
rary design of pollution control facilities to achieve maximum benefit of our pollution con-
trol expenditures.
The purpose of this manual is to provide the pulp and paper industry engineering commu-
nity with a new source of information for use in the planning, design, and operation of pres-
ent and future control facilities. It is recognized that there are a number of design manuals,
manuals of standard practice, and design guidelines currently available in the field that ade-
quately describe and interpret current engineering practices as related to traditional environ-
mental control design concepts. It is the intent of this manual to supplement this existing
body of knowledge by describing new pollution control methods and by discussing the
application of new techniques for more effectively removing a broad spectrum of contami-
nants from air and water discharges. This manual contains two parts; the first describes air
pollution control, while the second presents water and solid waste pollution control for the
pulp and paper industry.
Much of the information presented is based on the evaluation and operation of pilot,
demonstration, and full-scale plants. The design criteria thus generated represent typical
values. These values should be used as a guide and should be tempered with sound engineer-
ing judgment based on a complete analysis of the specific application.
This manual will be updated as warranted by the advancing state-of-the-art to include new
data as they become available and to refine design criteria as additional full-scale operational
information is generated. Part I of this manual, Air Pollution Control, is presented herein.
Part II, Water and Solids Pollution Control, is currently in preparation and will shortly be
available for inclusion into this manual.
xix
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CHAPTER 1
INTRODUCTION
From about 1965 to 1975, air pollution technology in the United States puip and paper
industry has undergone major advancements in the design of pollution abatement systems for
controlling gaseous and particulate emissions. These technological advances are particularly
significant in the kraft pulping segment of the pulp and paper industry because
reduced-sulfur gas and particulate emissions have been reduced by changes within the
production cycle itself.
A major difficulty confronting the design engineer has been the unavailability of the bulk of
this emerging practical process design information in a centralized source, such as
manual. The design approach emphasized in this manual combines process control
techniques to minimize the formation of air pollutants with treatment methods for removal
of air pollutants from process streams and flue gases.
The information presented in this manual is not limited to North America technology, but
also includes technology on internal process control in the Scandinavian pulp and paper
industry. Particular stress is placed on explanation of the chemical and physical processes
that generate air pollutants in specific unit operations so that the advantages and limitations
of both internal and external process control methods can be understood.
Most of the concepts presented in this manual have been demonstrated in actual field
installations. Some of these design concepts may be superseded by a rapidily advancing
technology; others will endure. Future design approaches must properly consider all aspects
of the relationship between the process and air pollution. in particular, the design engineer
must be cognizant of the rapidly increasing cost of energy and its direct relationship to
certain air pollution control measures. Additionally, deviations from design performance,
which become more important as emission limitations become more stringent, will be less
tolerable.
Chapter 2 through 13 emphasize the air pollution problems of the kraft or sulfate pulping
process. Chapter 14 emphasizes the sulfite process. Chapter 15 and 16 discuss air pollution
sources in pulp mills that are not a direct part of the pulping process. Chapter I 7 discuSSes
process monitoring of air pollutants.
i-I
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1.1 Gaseous and Particulate Matter Emissions from Kraft Pulp and Paper Mill Process
Sources
The United States pulp and paper industr includes more than 360 mechanical and chemical
pulp mills of all types. Further, the industry plans to build approximately 40 new mills
during the L970’s.
The industry has made a major contribution to our country’s effort to control excessive air
pollution from its facilities and has in the past cooperated extensively with government
regulatory agencies in dissemination of constantly emerging new control technology (1, 2,
3,4,5,6,7).
The atmospheric emissions from the kraft process include both gaseous and particulate
materials. The major gaseous emissions are malodorous reduced sulfur compounds, such as
hydrogen sulfide (H 2 S), methyl mercaptan (CH 3 SH), dimethyl sulfide (CH 3 SCI-1 3 ), and
dimethyl disulfide (CH 3 SSCH 3 ); oxides of sulfur (SO,j; and oxides of nitrogen (NOx). The
particulate matter emissions arc primarily sodium sulfate (Na 2 SO 4 ) and sodium carbonate
(Na 2 C0 3 ) from the recovery furnace, and sodium compounds from the lime kiln and smelt
tanks.
Both 1 l2 S and the organic sulfides arc extremely odorous and arc detectable at a
concentration of only a few parts per hillion. Thus odor control is one of the principal air
pollution problems in a kraft pulp mill. Gas volumes released per unit of production vary
considerably between individual process units. Most kraft pulp mill flue gas streams contain
appreciable amounts of water vapor.
A summary of the major external control techniques for gaseous and particulate emissions
from specific kraft pulp mill sources is presented in Table 1-1. Specific applications are
described in appropriate sections of the manual.
1.1.1 Reduced Sulfur
The major gaseous emissions from kraft pulp mill sources are the malodorous reduced sulfur
compounds, organic nonsulfur compounds, oxides of sulfur and oxides of nitrogen. The
malodorous sulfur gases emitted from kraft pulp mill sources aLl have extremely low odor
threshold levels of between I and 10 parts per billion (ppb) by volume (8). The most
common reduced sulfur compounds emitted from kraft pulp mill sources arc H 2 S,
CU 3 SCH 3 , and CH 3 SSCH 3 ; other alkyl sulfur compounds can be emitted in small quantities
from certain wood species.
The major potential sources for the reduced sulfur gas emissions to the atmosphere include
digester blow and relief gases, vacuum washer hood and seal tank vents, mnimltiple.effcct
1-2
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TABLE [ -I
EXTERNAL CONTROL TECHNIQUES FOR GASEOUS AND
PARTICULATE MATTER EMISSiONS FROM KRAFT PULP
MILL SOURCES
Particulate
Emission Source Gaseous Control Control
Digestcr gascs Incineration NA (not applicable)
Condensation
Washer vents incineration NA
Evaporator gases incineration NA
Scrubbing
Condensation
Condensate water Steam stripping NA
Air stripping
Condensatc Stripper Vent Incineration NA
Black Liquor Oxidation Incineration NA
Tower Vent
Tall Oil Vent Scrubbing NA
Recovery Furnancc Scrubbing Prccipi tators
Scrub bing
Fil Ira tion
Smelt Tank Scrubbing Scrubbing
Lime Kiln Scrubbing Scrubbing
Precipitators
Slaker Vent NA Scrubbing
Bleach Plant Scrubbing NA
Papcr Machines Incineration NA
Adsorption
Condensation
Power Boilers NA Cyclones
Precipitators
Scrubbing
evaporation hotwell vents, recovery furnace flue gases following direct contact vents, smelt
dissolving tanks, slaker vents, black liquor oxidation Lank vents, lime kiln exit vents and
wastewater treatment operations. Summaries of values on variations in gas flow rates,
malodorous sulfur gas concentrations, and emission rates per unit production for the kraft
process units are presented in Tables 1-2 to 1-4. These values are based on a variety of
sources.
1-3
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TABLE 1-2
TYPICAL GAS CHARACTERISTICS FOR KRAFT PULP MILL PROCESSES
Process Offgas Characteristic
Emission Source Flow Ratc Temperature Moisture Content
m 3 /t °C
(f 1 3 /ton) (°F)
D igesLer, Batch:
Blow Gases 3-6,000 65-100 30 -99
(96 .190,000) (150-210)
Relief Gases 0.3-100 25-60 3-20
(10-3,200) (80 .140)
Digester, Continuous 0.6-6 75- 150 35-70
(20-200) (170-300)
Washer Hood Vent 1,500-6,000 20-45 2-10
(48 000- 190,000) (70-110)
Vashcr Seal Tank 300-1,000 55-75 15-35
(9,600-32,000) ( 130-170)
Evaporator Hotwell 0.3-12 80-145 50-90
(10-400) (180-290)
BLO Tower Exhaust 500-1,500 70-80 30-40
(16,000-48,000) (160-180)
Itceovery Furnace 6,000-12 000 120-180 25 -35
(190,000-380,000) (250-360)
Smelt Dissolving Tank 500-1,000 70-110 35-45
(16,000-32,000) (160-230)
Lime Kiln Exhaust 1,000-1,600 65-95 25-35
(32 ,000-51,000) (150-200)
Lime Slaker Vent 12-30 65-75 20-25
(400-1,000) (150-170)
ilAt standard conditions of dry gas (211°C & 7ôOmml lg (70°F & 2992 in Hg))
Flow in cubic meters per metnc ton and (cubic feet per short ton)
1-4
-------
TABLE 1-3
TYPICAL REDUCED SULFUR GAS CONCENTRATIONS FROM KRAFT
PULP MILL SOURCES
Concentration (ppm by volume)
Emission Source 1-12 S CH 3 SH CH 3 SCH 3 Cl - I 3 SSCH 3
Digester, Batch:
Blow Gases 0-1,000 0-10,000 100-45,000 10-10,000
Relief Gases 0-2,000 10-5,000 100-60,000 100-60,000
Digester, Continuous 10-300 500-10,000 1,500-7,500 500-3,000
Washer Hood Vent 0-5 0-5 0-15 0-3
Washer Seal Tank 0-2 10-50 10-700 1-150
Evaporator Hotwell 600-9,000 300-3,000 500-5,000 500-6,000
BLO Tower Exhaust 0-10 0-25 10-500 2-95
Recovery Furnace 0-1,500 0-200 0-100 2-95
(after direct contact
evaporator)
Smelt Dissolving Tank 0-75 0-2 0-4 0-3
Lime Kiln Exhaust 0-250 0-100 0-50 0-20
Lime Slaker Vent 0-20 0-1 0-1 0-1
Both oxides of sulfur ( 50 x) and oxides of nitrogen (NOx) can be emitted in varying
quantities from specific sources in the kraft chemical recovery system. The major source of
sulfur dioxide (SO 2 ) emissions is the kraft chemical recovery furnace, because of
combustion of sulfur-containing black liquor fuel. Under certain conditions, somewhat
similar quantities of sulfur trioxide (SO 3 ) can be released to the atmosphere, particularly
when residual fuel oil is added as an auxiliary fuel (9). Lesser quantities of SO 2 can also be
released from the lime kiln and smelt dissolving tank. Trace quantities of sulfur oxides may
also be released from other kraft mill sources. Oxides of nitrogen can be formed in any fuel
combustion process by the reaction between oxygen and nitrogen at elevated temperatures.
1-5
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TABLE 1-4
TYPICAL REDUCED SULFUR GAS EMISSION RATES FROM KRAFT PULP
MiLL SOURCES
Emission Rate, kg sulfur pcr metric ton of air dried pulp
Emission Source H 2 S CH 3 SH CH 3 SCH 3 CH 3 SSCH 3
Digester, Batch:
BlowGases 0-0_i 0-1.0 0-2.5 0-1.0
Relief Gases 0-0.05 0-0.3 0.05-0.8 0.05-1.0
Digester, Continuous 0-0.1 0.5-1.0 0.05-0.5 0.05-0.4
Washer Hood Vent 0-0.1 0.05-1.0 0.05-0.5 0.05-0.4
Washer Seal Tank 0-0.01 0-0.01 0-0.05 0-0.03
Evaporator Hotwell 0.05-1.5 0.05-0.8 0.05-1.0 0.05-1.0
BLO Tower Exhaust 0-0.01 0-0.1 0-0.4 0-0.3
Recovery Furnace 0-25 0.2 0-1 0-0.3
(after direct contact
evaporator)
Smelt Dissolving Tank 0-1 0-0.8 0-0.5 0-0.3
Lime Kiln Exhaust 0-0.5 0-0.2 0-0.1 0-0.05
Lime Slaker Vent 0-0.01 0-0.01 0.0.01 0-0.01
The major constituent formed in nitric oxide (NO), a small portion of which can be oxidized
to form nitrogen dioxide (NO 2 ): together they arc classified as total oxides of nitrogen.
Nitrogen oxide emissions from kraft pulp mill process sources, such as the recovery furnace
and lime kiln, are normally lower than for most other fuel combustion processes. This is
primarily due to the large quantities of water present in black liquor and lime and which act
as a heat sink to suppress the flame temperature. Larger quantities of oxides of nitrogen can
be formed, however, when auxiliary fuels such as natural gas or fuel oil are added to the
recovery furnace.
1-6
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TABLE 1-5
TYPICAL EMISSION CONCENTRATIONS AND RATES FOR SO, AND NO
FROM KRAFT PULP MiLL COMBUSTION SOURCES
Concentration (ppm by vol.) Emission Rate, kg/t*
Emission Source SO 2 SO 3 NOx (as NO 2 )
Recovery Furnace:
No Auxiliary Fuel
Auxiliary Fuel
Added
Lime Kiln Exhaust
Smelt Dissolving
Tank
SO 2 SO 3 NOx(asNO2)
0-1,200 0-100 10-70 0-40 0-4 0.7-5
0-1,500 0-150 50-400 0-50 0-6 [ .2-10
0-200 ---- 100-260 0-1.4 ---- 10-25
0-100 ---- ---- 0-0.2
*kilograms per metric ton of air dried pulp.
A summary of concentrations and emission rates for oxides of sulfur and oxides of nitrogen
for specific kraft pulp mill sources is presented in Table 1-5. The information included is
based upon a variety of industry sources. The extreme variations in operating conditions
that occur in the industry, including operating combustion temperature and type of fuel,
account for the broad ranges in these data.
1.1.3 Organic Compounds
Organic compounds other than those containing sulfur can also be emitted in varying
quantities from several different kraft pulp mill process sources. The major types of
materials that can be released to the atmosphere include terpenes, hydrocarbons, alcohols,
phenols and other organic compounds liberated from wood. Additional organic compounds
can be produced when organic materials are applied as coatings to paper sheet or can be
induced when spent caustic solutions are used as chemical make-up for the process.
The primary significance of these materials is that they may either act directly as odorant
gases or as liquid particulate carriers for odorous sulfur gas molecules, particularly the
terpene compounds. The olefinic hydrocarbons or terpenes may undergo photochemical
reactions in polluted atmospheres.
1-7
-------
The major potential process sources of organic- nonsulfur compound emissions to the
atmosphere include digester blow and relief gases, multiple-effect evaporator noncon-
densable gases, brown stock washer hood and seal tank vents, black liquor oxidation tower
vents, black liquor storagc tank vents, direct-contact evaporator exhaust gases, digester arid
evaporator condensate water vents and wastewater treatment facilitics, and paper machine
coating and dricr vents.
Major process variables that affect emissions of these compounds to the atmosphere include
the wood species being pulped or the type of coating material applied, the respective organic
or condensate stream temperature, the volatility of the respective organic compounds, and
the type and effectiveness of any air pollution control device.
1.1.4 Particulate Matter Emissions
The major potential process sources of particulate emissions from the kraft chemical
recovery system are the recovery furnace, the smelt dissolving tank, arid the lime kiln. The
recovery furnace is the largest potential particulate emission source. The major chemical
constituent in the recovery boiler particulate emissions is Na 2 SO 4 , with smaller quantities
of Na 2 CO 3 and sodium chloride (NaC I) also present. The smelt dissolving tank vents and
lime kiln exhaust gases are also sources of varying quantities of particulate matter consisting
primarily of carbonate, hydroxide, sulfate and chloride salts of calcium and sodium. Particle
sizes from these sources can range from 0_i pm (4 X iO in) to greater than 1000 p (4 X
102 in) in diameter for uncontrolled emissions and from 0_i to 10 pm (4 X i0 6 to 4 X
icr in) in diameter where these sources have high efficiency particulate control devices.
The two major types of particulate matter control devices employed for kraft recovery
furnaces are electrostatic precipitators (ESP) following cyclone or cascade-type direct
contact evaporators, and venturi-type evaporator-scrubbers in a one-or two-stage configura-
tion. Low pressure drop secondary wet serubbers have been employed to supplement older
and less efficient primary particulate collection devices at several existing mills to alleviate
particle fallout in areas adjacent to the plant premises. Packed tower or showered mesh
demister scrubbers are employed for particulate control on smelt dissolving tank exhaust
gases, while venturi or eyclonic scrubbers are normally used for particulate control on lime
kiln or fluorosolid calciner exhaust gases. The amount of particulate matter emitted from
kraft pulp mill process sources depends both on the process operating conditions and on the
types and collection efficiencies of any control devices employed.
A summary of typical ranges in particulate concentrations and emissions rates from kraft
pulp mill process sources is presented in Table i-6.
1-8
-------
TABLE 1-6
TYPiCAL CONCENTRATIONS AND EMISSION RATES FOR
PARTiCULATE MATTER FROM KRAFT PULP MILL SOURCES
(AFTER CONTROL DEVICES)
Emission Source Concentration Emission Rate
g/scm ( gr/scf) kg/t ( lb/ton )
Recovery Furnace:
After Electrostatic
Precipitator 0.06-1.1 (0.03-0.5) 0.5-12 (1.0-24)
After Venturi
Evaporator 0.9-2.3 (0.4-1.0) 7-25 (14-50)
Lime Kiln 0.07-1.1 (0.03-0.5) 0.15-2.5 (0.3-5.0)
Smelt Dissolving Tank 0.04-2.3 (0.02-1.0) 0.01-0.5 (0.02-1.0)
1.2 Gaseous and Particulate Emissions from Sulfite Pulp and Paper Mill Process Sources
The primary emissions from the sulfite pulping process are SO 2 and particulate matter. In
special cases of burning alkaline sulfite liquor in recovery furnaces under reducing
conditions, H 2 S emissions may also occur. Otherwise, there are practically no organic
reduced sulfur compounds produced in the sulfite process. Nitrogen oxides are emitted from
various combustion sources, particularly from the recovery furnace of ammonium-based
mills.
1.2.1 Sulfur Dioxide
Various process sources within the sulfite mill can emit SO 2 . The main sources are the
digester blow pits, multiple-effect evaporators, and liquid burning or chemical recovery
systems. Minor process sources include pulp washers and the acid preparation plant. Typical
values of SO 2 emission rates are listed in Table 1-7.
1.2.2 Particulate Matter
The recovery furnace is the significant process source of particulate matter in a sulfite pulp
mill. Potential particulate matter emissions depend greatly on the degree of recovery of
sulfite waste liquor, as well as on the degree of control of particulate matter.
1-9
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TABLE 1-7
TYPICAL SO 2 EMISSION RATES FROM
SULFITE PULP MiLL SOURCES
Emission Source Emission Rate, kglt (lb/ton)*
Uncontrolled Controlled 0
Blow Pit:
Hot blow 30-75 (60-150) 1-2.5 (1-5)
Cold blow 2-10 (4-20) 0.05-0.3 (0.1-0.6)
Evaporators 1-30 (2-60) 0.025-1 (0.05-2)
Recovery Process 80-250 (1.60-500) 6-20 (12-40)
Washers 0.5-1 (1-2)
Acid Preparation 0.5-1 (1-2)
*Per mass I (ton), of air dried pulp.
**A1kal ne scrubbing of gases.
1.2.3 Nitrogen Oxides (NO )
In ammonium-bascd sulfite pulp mills, coml ustion of the spent sulfite liquor will result in
emissions of nitrogen oxides from the recovery furnace. Emissions from one such system
ranged from 4.7 kgft to 11.8 kg/t (9.4 to 23.6 lb/ton).
1.3 Power Boilers
The pulp and paper industry is a major energy consumer in the United States, accounting
for about 2.2 percent of the total national energy consumption. This amounts to
approximately [ .6 X 101 8 j per year (1 5 X 1015 BTU/yr) of which approximately one-half
is associated with the manufacture of kraft pulp and paper (10). Typical energy
consumption requirements for a kraft pulp mill are about 28.7 CJ per metric ton of air dried
pulp (30 million BTU/ton) of which 50 to 60 percent can normally be supplied by
combustion of the black liquor solids (11).
For mills employing on site debarking, an additional 20-30 percent of the energy
requirement can be supplied by the burning of waste wood in bark boilers. As a result, it is
normally necessary for kraft pulp mills to obtain about 5 to 30 percent of their energy
requirements by burning supplementary fuels such as coal, fuel oil, and natural gas. The
1-10
-------
total energy that must be supplied by supplementary combustion of coal, oil, gas, or wood
in kraft pulp and paper mill power boilers can range from less than 0.9 million to more than
5.4 GJ per metric ton of pulp (0.75 to 4.5 million BTU per ton of pulp).
The exact energy requirements for the auxiliary fuel burning of coal, oil, gas, or wood will
vary between individual mills depending on their respective energy balances, physical
characteristics, and availability of fuels in each local area. The major air pollutants of
possible concern from auxiliary fuel burning operations include particulate matter from coal
and wood, sulfur oxides from coal and fuel oil, and nitrogen oxides from coal, oil, gas, and
wood. Available particulate matter control devices for coal and wood-fired power boilers
include electrostatic precipitators, liquid scrubbers, fabric filters, and mechanical cyclones.
A summary of uncontrolled air pollutant emissions from auxiliary fuel combustion in power
boilers in the pulp and paper industry is presented in Table 1.8 (12).
1-11
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TABLE 1-8
UNCONTROLLED AIR POLLUTANT EMISSIONS FROM FUEL
COMBUSTION IN AUXILIARY POWER BOILERS
Air Pollutant Emission Rate; kg /b 6 kJ (lb/b 6 BTIJ)
Air Pollutant Bituminous Coal* Residual Fuel Oil** Natural Gas Waste Wood
Particulate Matter 0.38 (0.95) 0.024 (0.060) 0.005 (0.01) 1.50 (3.75)
Sulfur Oxides
(asSO 2 ) 0.84(2.1) 0.46(1.1) 0.16(0.40)
Nitrogen Oxides
(as NO 2 ) 0.39 (0.98) 0.23 (0.58) 0.16 (0.40) 0.43 (1.1)
Hydrocarbons
(as CH 4 ) 0.007 (0.02) 0.17 (0.43) 0_li (0.28)
Carbon Monoxide 0.021 (0.053) 0.11 (0.28)
wBased on average heating value of 25.7 Mi/kg coal (11,000 BTU/Ib)
# Ba d on average heating value of 41 9 GJ/m 3 oil (150,000 BTU/gal)
+Based on average heating value of 39.1 Mj/m 3 natural gas (1,050 BTIJ/ft 3 )
++Bascd on waste wood heating values as follows:
[ tern Units Dry Basis Wet Basis
Moisture Content % by mass 0.0 50.0
Heating Value Mi/kg 18.6 9.3
Heattng Value BTU/lb 8000 4000
1-12
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1.4 References
1.4.1 Cited References
1. Hendrickson, E. R., Robertson, J. E., and Koogler, J. B., Control of Atmospheric
Emissions in the iVood Pulping Industry, Volumes I, ii, 111. Final Report, Contract No.
CPA 22-69-18, U.s. Department of Health, Education, and Welfare, National Air
Pollution Control Administration, Raleigh, North Carolina, March 15, 1970.
2. Proceedings of the Symposium on Recovery of Pulping Chemicals: Helsinki, Finland,
May 13-17, 1968, Finnish Pulp and Paper Research Institute, EKONO, Helsinki,
Finland, 1969.
3. Cooper, H. B. H., and Rossano, A. T., Jr., Odor Control Technology for Kraft Pulp Mills.
Report prepared for U.S. Environmental Protection Agency Odor Control Technology
Manual, University of Washington, Seattle, Washington, August, 1970.
4. Proceeding of the International Conference on Atmospheric Emissions from Sulfate
Pulping, April 28, 1966, Sanibel Island, Florida. Hendrickson, E. R. (ed.) Sponsored by
USHPS, University of Florida, and National Council for Air and Stream Improvement.
Deland, Florida, E. 0. Painter and Printing Co., 1966.
5. Atmospheric Emissions from the Pulp and Paper Manufacturing industry. EPA-450/].-
73-002. September 1973. (Also published as NCASI Technical Bulletin No. 69,
February, 1974.)
6. Field Surveillance and Enforcement Guide—Jl’ood Pulping Industry, US. EPA Contract
No. 68-02-0618. Prepared for EPA, Research Triangle Park, N.C., October 15, 1973
(Revised Draft).
7. Galeano, S. F., and Leopold, K. M.,A Study of Emissions of Nitrogen Oxides in the Pulp
Mill. Tappi, 56:74-76, March 1973.
8. Wilby, F.V., Variations in Recognition Odor Threshold of a Pond. Journal of Air
Pollution Control Association, 19:96-100, February 1969.
9. Maksimov, V. F., Bushmelav, V. A., Torf, A. I., and Lesohhin, V. B., Testing the
Turbulent Flow Venturi Apparatus, Bumazhnaya Proyshlennost, 40:14-15, May 1965.
10. Personal communication with Dr. Ronald Slinn, American Paper Institute, New York,
New York, November, 1973.
11. Miller, R. R., One Pulp and Paper Company’s View of the Energy Crisis. Tappi, 57:
62-64, February, 1974.
1-13
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12. Compilation of Air Pollutant Emission Factors. U.S. Environmental Protection
Agency, Office of Air Programs, Research rfriangle Park, N.C. Publication No. AP.42
(Revised). February 1972.
1.4.2 Additional Reading
1. Britt, K. W. (ed.), Handbook of Pulp and Paper Technology. New York, Reinhold
Publishing Company, 1964.
2. Wcnzyl, I-I., Kraft Pulping: Theory and Practice. New York, Lockwood Publishing
Company, 1967.
3. Whitney, R. P. (cd.), Chemical Recovery in Alkaline Pulping Processes, Tappi
Monograph Series No. 32. Technical Association of the Pulp and Paper industry, New
York, New York, 1968.
4. Turpentine Recovery Systems. Pulp Chemicals Association, New York, 1972.
5. Tall Oil Recovery. Pulp Chemicals Association, New York, 1968.
].J4
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CHAPTER 2
DIGESTER GASES
The digestion process is the third most important source of odor pollution in the kraft
process. Black liquor combustion and weak black liquor evaporation are first and second,
respectively. Gases from the digester contain organic sulfur compounds from reactions
between components of the wood and the sulfide of the white liquor. These gases also
contain some H 2 S, turpentine, and traces of methanol (CH 3 OI-l), ethanol (CH 3 CH 2 OH),
and acetone (CH 3 COCH 3 ), as well as displaced air.
The composition and quantity of digester gases will differ between batch digesters and
continuous digesters. Variations will also occur with the type of wood, the sulfide
concentration in the white liquor, the final cooking temperature, and the cooking time.
Most of these factors are determined by production requirements and vary widely with the
production schedule.
The basic method for minimizing odor pollution in the digester area is to effect adequate
condensation and to contain the relief and blow gases. As a further step, contained gases
may be incinerated. In the following sections important factors in using these methods are
discussed.
2.1 Batch Digesters
Batch digesters often present air pollution problems because of surges of gas flow produced
during blowing. These surges can temporarily overload the condensing or heat recovery
system. Economics of production demand that a plant be operated at full capacity. Often,
the capacity of the batch digesters is limited by their condensing systems, so that full-scale
production can overload these systems. This problem can be compounded since liquid
carryover from short term overloads can increase fouling of the condensing systems: thus,
further lowering the capacity and making the systems inadequate for even less than design
flows. The condensing system for batch digestcrs must be designed for peak flows and all
components in the heat recovery system must be kept in good operating condition. To allow
proper operation, the system must have enough temperature and pressure difference
measurements to enable the operator to judge the condition of the system. Figure 2.1 shows
a typical batch digester and its blow heat recovery system. Point sources of odor release are
indicated in this figure.
2-1
-------
SECONDARY
CONDENSER
HEAT
EXCHANGER
11 HOT WATER
10 COOLING WATER
RELIEF GAS
TURPENTINE
FOUL CONDENSATE
BLOW GAS
5 HOT WATER
4 WARM WATER
BLOW CONDENSTATE BLEED
2 FRESH WATER
0 PULP—LIQUOR
POINTS OF POSSIBLE ORDER RELEASE ARE ENCIRCLED BY 0
FIGURE 2-1
BATCH DIGESTER FLOW SHEET
SEPERATOR
LIQUOR
4-
CHIPS
12
13
14
15
-------
2.1.1 Blow Gases
Batch digester blow gases are a major cause of air pollution. As a rough guide to the volumes
that can be expected, a single blow will produce approximately one ton of steam per ton of
air dry pulp (90 percent absolute dry fiber). Blowing a 200 m 3 (7100 ft 3 ) digester to
atmospheric pressure from 0.64 MPa (78 psig) produces about 20 t (22 tons) of pulp and
34,000 m 3 (1,200,000 ft 3 ) of steam within a 20 minute period (Figure 2-2). Effective
condensation of this volume of steam requires a blow heat recover s stem with adequate
capacity and a control system which reacts quickly but retains stability.
The vent from the blow heat accumulator (flow 6 in Figure 2-1) might typically show flow
histories as illustrated in Figure 2-3. These flow histories were recorded with a pitot tube.
Normally, about 90 percent of the volume is steam, and the noncondensable portion i
about 3 m 3 per metric ton of pulp (96 ft 3 /ton). Typical flow and composition ranges were
given in Tables 1-2 and 1-3. A successful blow gas treatment program starts with an efficient
blow heat recovery system (section 2.2.1.1) followed by a flow equalization system and a
gas incineration system.
Blow gas flows, such as those shown in Figure
reasonable flow equalization system and thus
treatment program. Loss of recoverable heat is also
100
90
80
70
60
50
40
30
20
I0
0
2-3, cases 2 and 3, will overload any
seriously degrade the entire blow gas
probable in these two cases.
c
2500
2000
1500 u-
U i
IV A1 Cl)
500
10 20 30
TIME, MINUTES
FIGURE 2-2
KRAFT BATCH DIGESTER BLOW STEAM FLOW (1)
-C
C
0
IL
U i
I —
C,)
0
-J
0
2-3
-------
A
0
dt 550m 3 /blow
MEAN - 1400 m 3 /h
20
E
‘ Q 10
9
Li
40
-c 30
Q20
0
I0
0
CASE 2. MALFUNCTION OF BLOW HEAT RECOVERY HEAT EXCHANGERS
50
40
30
20
I0
0
23
V(t) dt 4500m 3 /btow
MEAN - 11,700 m 3 /h
0 5 0 15 20 25
TIME, MINUTES
CASE 3. INCREASING MALFUNCTION OF HEAT EXCHANGERS
FIGURE 2-3
KRAFT BATCH DIGESTER BLOW GAS FLOW AFTER CONDENSING
AND WITHOUT EQUALIZATION (2)
5 10 15
TIME, MINUTES
CASE I. NORMAL OPERATIONS
20 25
23
V( t ) dt 300m 3 /b low
MEAN - 8200 m 3 /h
0
10 15
TIME, MINUTES
20 25
N,
E
Nb
2-4
-------
Final odor gas treatment by white liquor scrubbing is not very efficient because of the
dominance of nonionizable organic sulfur compounds. Only H 2 S and CH 3 SH can be
efficiently recovered by alkaline scrubbing (Figures 2-4 and 2-5).
Therefore, incineration of the blow gases (section 4.3) after proper gas flow equalization
(section 4.2.2) is recommended.
2.1.1.1 Blow Heat Recovery
The blow heat recovery system, as shown in Figure 2-1, and its proper operation will
significantly affect the further treatment of both blow gases and blow condensates. Factors
that may affect the batch digester air pollution abatement program are final blow pressure,
blow tank drop separation, primary blow steam condenser, secondary blow steam con-
denser, blow heat accumulator and blow heat recovery heat exchanger.
The blow gas flow is directly proportional to the final blow pressure (i.e., the higher the
final blow pressure, the more violent the blow gas flow, and the more difficult it is to
condense and collect the blow gases for treatment). Decreasing the final pressure, however,
pH at 25°C
FIGURE 2-4
VAPOR PRESSURE OF 0.01 M H 2 S VS. pH AT
VARIOUS TEMPERATURES (4)
I
E
C,)
I
0
o
U-
0
LU
C))
C l)
LU
a-
700
600
500
400
300
200
100
0
L
185°C
Slt\\
—
140°C
2 4 68101214
2-5
-------
90
185°C
0—
160°C
o.—
140°C
__
10 II 12 13 14
pH AT 25°C
FIGURE 2-5
VAPOR PRESSURE OF O.O1N CH 3 SH VS. pH
AT VARIOUS TEMPERATURES (4)
may require relief or other operations that take time and consequently decrease production
rates. A violent blow also means more fiber and liquor carryover to the heat recovery system
with possible fouling of heat exchanger surfaces. Lower blow pressure also gives better puip
quality. Final blow pressure is usually 0.4-0.5 MPa (58-72 psig).
An efficient blow tank drop separator stops fiber and liquor carryover (3). Fiber and liquid
carryover fouls filters and heat exchanger surfaces, thereby decreasing heat transfer and
increasing pressure drops in the system. A reduction in system capacity results in inferior
condensation of blow steam and leads to difficulties in noncondensable blow gas treatment.
The capacity of the blow steam condenser must be sufficient for handling the largest
digester, the highest blow pressure, the shortest blow period, the highest water temperature,
or any combination of these. This applies as well to the main condenser pump as to the
direct contact condenser itself. Usually the addition of cold water to the primary condenser
(flow 2 in Figure 2-1) will improve condensation, but will increase the blow condensate
bleed (flow 3 in Figure 2-1). A most important factor is the condenser control system.
‘80
E
E
i 70
(I)
a)
60
50
0
40
LU
a:
30
C /)
LU
a:
a-
a:
2 10
0
2-6
-------
Usually it is a simple temperature control of the water flow (See Figure 2-6). This control
system is too slow for the abrupt beginning of the blow and, .‘hen made more sensitive to
compensate for speed, it often becomes unstable. No really good solution to this control
problem exists. One solution to minimize blow gas treatment problems is Lo install a feed
forward control. This control can be a simple on-command to open the condenser water
valve at a specified time before the digester blow valve is opened. This results in lowering the
temperature of the hot water produced in the blow heaL recovery sysLem (I). Condenser
outlet temperature should be kept around 90° C ( [ 940 F).
A secondary condenser is very common as a backup for the primary condenser. If properly
controlled it will accommodate temporary overloads on the primary condenser, thus
facilitating treatment of noncondensable blow gases. It may be either a direct contact
condenser or a surface condenser (the latter gives a more concentrated condensate). A
suitable condenser outlet temperature (see Figure 2-7) should be arouiid 50° C (122° F).
For the blow heat accumLilator, the two determining factors are sufficient size to
accommodate large blows and long pauses between blows, and an internal construction that
will cause the zone between hot- and cooled-water to be as sharp and undisturbed as
possible. Frequently, hot water does find its way down to the bottom of the accumulator,
enters the condenser, and thereby drastically reduces the condensing capacity. Subsequent
gas handling problems occur.
PRIMARY
CONDENSER
BLOW
CONDENSATE
HOT WATER
WARM WATER
COLD WATER
FIGURE 2-6
BLOW HEAT RECOVERY CONTROL SYSTEM
BLOW
GAS
B.DW
STEAM
r
2.7
-------
x
S
20
E
I5
cI S
x
z
o 10-
x
I— x
L i i x
C i
0
0
o
.
00
‘ 0
50 60 70 80 90
CONDENSER OUTLET TEMPERATURE, °C
• HYDROGEN SULFIDE
o METHYL MERCAPTAN
x DIMETHYL SULFiDE
o DIMETHYL DISULFIDE
FIGURE 2-7
ODOR COMPOUNDS IN RELIEF GAS AFTER TURPENTINE
CONDENSER AS A FUNCTION OF CONDENSER
OUTLET TEMPERATURE (2)
The internal construction should consist of adequate baffles to disperse the hot condensate
in top layers that are as even as possible. The accumulator should also be evenly insulated to
avoid local cold spots where liquid will cool and flow down, mixing the contents of the
accumulator.
2-8
-------
Experiences from mills show that installed heat exchanger surfaces, which are initially
sufficient, may later prove too small. This change occurs because surfaces gradually get
fouled, especially on the blow heat side, by fiber and liquor carryover. These dirty surfaces
reduce heat transfer and increase pressure drops, thus rcducing flows and furthcr decreasing
heat transfer. Spiral heat exchangers have an advantage over plate exchangers in decreasing
fouling by fiber and liquor carryover, but they are more difficult to clean. Another factor,
which will reduce heat transfer, is a reduced hot water demand (flow 5 in Figure 2-1) or a
higher warm water temperature (flow 4 in Figure 24). Reduced heat transfer means a higher
temperature in the accumulator bottom and reduced condensation of the blow gases.
Injecting fresh cool water (flow 2 in Figure 2-i) to the primary condenser with controlled
temperature counteracts the reduction in condensation. This method is used extensively,
but it has the great drawback of increasing substantially the blow condensate bleed (flow 3
in Figure 2.]).
The blow condensate hlecd has to be treated, for instance, by steam stripping, and the
stripping equipment investment and the stripping steam consumption will directly affect the
size of the bleed. Thus, if the blow heat recovery capacity is insufficient, a trade-off
between air pollution abatement and water pollution abatement must be made. A successful
odor abatement program should include monitoring of temperature and pressure drops over
the blow heat recovery heat exchangers to cheek their performance.
2.1.1.2 improving Blow Heat Recovery
The most common approach to blow gas treatment is to use a blow gas collection and flow
equalization system followcd by gas incineration. Any malfunction in the blow heat
recovery chain will adversely affect the blow gas treatment, primarily the blow gas
collection and flow equalization system, as clearly demonstrated in Figure 2-3.
Mills with odor abatement problems caused by inadequate blow heat recovery can improve
their situation by:
1. improving the blow heat recovery control system through:
a. More extensive instrumentation of condensing, heat accumulating, and heat
exchanging systems, and
b. Better tuning of the control system.
2. Increasing the blow heat recovery capacity through installation of:
a. More heat exchanger surface,
b. More pump capacity for the primary blow steam condenser,
2-9
-------
c. More pump capacity for the heat exchanger circulation,
d. Secondary blow-steam condenser,
e. BctLer baffles in the accumulator, and
f. More accumulator volume.
3. Increasing the flow equalization capacity through installation of more gas (up to
10 times more) accumulator volume.
4. Decreasing the blow heat recovery system load through prolonging the blow
period, thereby decreasing pulp production.
An odor control program for the blow heat recovery system should consider recommenda-
tion 1, above, first, and then the others. Recommendatwn 4 should only be used as a last
resort.
2.1.2 Relief Cases
The purpose of the digester relief is to remove air and other noncondensable gases during
operation and to reduce digester pressure before blowing. Relief takes place more or less
continuously during digestion, as well as during the final deliberate blow pressure relief
before blowing. For softwood, the amount of steam in the continuous relief is about 180 kg
(400 Ib), and in the final relief around 90 kg (200 Ib) (5). In modern batch digcsters, the
final blow pressure is more efficiently and swiftly reduced by introducing cooler weak black
liquor into the upper part of the digester.
The relief is usually passed through a surface condenser, thus producing hot water. When
pulping softwood, the relief condensate will contain turpentine, which is recovered in a
separating vessel. The noncondensable relief gases flows and compositions are presented in
Table 1-2 and Table 1-3. Hardwood usually produces more relief gas than softwood. A
normal softwood value is around 1 m 3 per metric ton of pulp (32 ft 3 /ton).
The relief gases do not present a major problem because of the relatively small volume and
even flow as compared to blow gases. As with the blow gases, the recommended treatment
is incineration (section 4.3).
2.1.3 Turpentine Recovery
Turpentine recovery takes place with turpentine-containing softwoods. Recovery is through
gas relief to the turpcntine recovery system. Typically about 270 kg steam/t (540 lb/ton) of
pulp is relieved to and condensed in the recovery system. The condensate is separated into
2-10
-------
one turpentine fractton and one underilow of contaminated condensate in the decanter.
This condensate is one source of odor in a kraft mill and can be treated by steam stripping
(section 5.5).
The amount of odor compounds in the relief gas can be decreased by decreasing the
condensate outlet temperature (Figure 2-7), which also produces a lower hot water
temperature from the condenser.
Decreasing condenser outlet temperature also means greater safety in collecting and
handling relief gases when there is a possibility of contact with air, since lower condensing
temperature correspondingly produces less turpentine in the relief gas (Figure 2.8).
2.2 Continuous Digester Gases
Continuous digesters present a much smaller pollution problem than batch digesters because
contaminated condensates and odorous gases flow at a regular rate. Treatment capacity can
be designed for the mean flow, without the need for peak flow equalization as with batch
digesters. Continuous digester odor gases do not differ significantly in composition from
batch digester odor gases (see Table 1-3).
The amount of noncondensable gases released from the digester itself varies according to
how the flash steam is used (Table 1-2). A rather typical downflow of a continuous digester
arrangement is presented in Figure 2-9. Other types of continuous digesters exist, but they
do not differ very much with respect to air pollution generation. Some types may have a
relief vent from the top of the digester to the turpentine recovery system. Most existing
units have a countercurrent wash zone in the bottom of the digester. The wash liquor (flow
17 in Figure 2-9) temperature is 75.800 C (167-176° F). This gives a pulp-liquor (flow 1 in
Figure 2-9) temperature of 80-85° C (176-185° F). This so-called “cold blow” produces
very minor odor emissions from the blow tank. Thus, the major gaseous odor release from
the digester will be the flash steam. This steam can be used and condensed in many different
ways yielding noncondensable gases, which may be collected and incinerated. The total
amount of flash steam is about 0.8 ton per ton of pulp.
2.2.1 Flash Steam
The spent liquor from the continuous digester is drawn off and expanded, or flashed,
usually in two stages. The flash steam from the primary flash tank is usually used to
impregnate the chips in the presteaming vessel. The presteaming vessel relief, which contains
the noncondensable gases from the primary flash steam and from the presteamed chips, then
passes to a turpentine recovery system. The amount of primary flash steam to the
presteaming vessel is 0.5-0.6 ton per ton of pulp. The secondary flash steam amount is
0.2-0.3 ton per ton of pulp, and it may be used for various purposes.
21 1
-------
0 20
ii
/
II
a,
I,’
I’
I,
II
I
I
II
‘I
III
I,
I,,
I,
I,
-PINENE !
I
Il
/
‘I
I,,
II
,,
I,
I,,
1/
I,
/
I
I
I
I
/
/
/
/
/
/
60 80
TEMPERATURE, °C
FIGURE 2-8
VAPOR PRESSURES FOR METHANOL, WATER,
a-PINENE & 13-PINENE (6)
METHANOL
o
I
E
E
w
800
700
600
500
400
300
200
100
0
WATER
ft -PINENE
./
/
40
100
120
140
2-12
-------
13 WHITE LIQUOR
12 CHIPS
VENT
HOT WATER
COOLING WATER
14 STEAM
TURPENTINE
15 COND
FOUL CONDENSATE
FLASH STEAM
WEAK LIQUOR
PULP + LIQUOR
WASH LIQUOR
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY 0
FIGURE 2-9
CONTINUOUS DIGESTER FLOW SHEET
2.2.1.1 Turpentine Recovery
The relief from the presteaming vessel is brought to a condenser. The amount of condensate
is about 0.1 ton per ton of pulp. When pulping soft woods, the condensate is separated in
the decanter into a turpentine fraction (flow 8 in Figure 2-9) and an odorous water fraction
(flow 7 in Figure 2-9), which is typically treated by steam stripping. The amount and
composition of the noncondensable gases (flow 9 in Figure 2-9) varies between values given
in Tables 1-2 and 1-3. The amount of odorous compounds can be decreased by decreasing
the condenser outlet temperature, as earlier shown in Figure 2-7. The remaining odorous
gases may be collected and incinerated (section 4.3). For safety aspects, see section
4.1.3.
11
10
0
2-13
-------
2.2.1.2 Flash Heat Recovery
Primary flash heat recovery takes place in the presteaming vessel, as described previously.
The secondary flash steam (flow 16 in Figure 2.9) may be utilized in different parts of the
kraft mill. Some possibilities are, using it as additional impregnating steam, routing it to
the turpentine recover)’ system (increasing the turpentine and hot water output),
condensing it in a separate flash steam condenser for hot water production, using it as a
partial steam source for a black-liquor evaporation plant, and using it as a partial steam
source for a contaminated condensate steam.stripping column.
Wherever the flash steam is condensed, its noncondcnsable components will remain and
must be collected and treated, preferably by incineration. If treatment does not occur, part
of the digester odor is simply transferred to another release point, such as the evaporation
plant.
2.3 References
1. Kock, P. A., Treating Kraft Digester Waste Gases. M.S. Thesis. Chemical Engineering
Department, Helsinki Technical University, Finland. September 12, 1972. (Swedish).
2. Kekki, R., Kraft Mill Odor Abatement by Condensate Stripping and Waste Gas
Incineration. M.S. Thesis, Wood Industry Department, Helsinki Technical University,
Finland. September 18, 1969. (Finnish).
3. Sarkanen, K. V., Hrutifiord, B. F., Johanson, L. N., Gardner, H. S., Kraft Odor. Tappi,
53:776.783, May 1970.
4. Martin, C. C., Fiber Carryover with Blow Tank Exhaust. Tappi, 52 :2360.2362, Decem.
ber 1969.
5. The Finnish Paper Engineers’ Association (SPY). The Pulping of Wood. Helsinki,
Frenckellin Kirjapaino Oy, 1968 (Finnish).
6. Weast, P. C. (ed.), Handbook of Chemistry and Physics, 47th edition. Cleveland, The
Chemical Rubber Co., 1966. p. D105.D 138.
2.14
-------
CHAPTER 3
EVAPORATiON GASES
The evaporation of black liquor is one of the three major malodorous gas producing
processes in a kraft mill; the other two are black liquor combustion and wood digestion.
For the final evaporation of black liquor to combustion strength, there are two different
methods in current use. One method is indirect evaporation with steam; the other is direct
contact evaporation with hot flue gases from the recovery boiler. This latter method causes
an air pollution problem in the combustion of black liquor and is discussed in Chapter 10.
The evaporator gases discussed here are only the noncondensable gases generated from
indirect steam evaporation of black liquor. Evaporator gases will contain sulfur and organic
compounds boiled off from the black liquor, plus displaced and leaked air.
The amount and composition of the evaporator gases vary widely depending on black liquor
properties, such as originating wood, pH, and sulfide concentration, and evaporation plant
properties, such as temperature level, equipment type, and plant condition.
The most significant difference between evaporator and digester gases is that the dominating
compounds in the evaporator gases are H 2 S and CH 3 SH instead of organic sulfur
compounds. This feature makes odor abatement and sulfur recovery through white liquor
scrubbing entirely feasible. Incineration is another possible control technique.
3.1 Black Liquor Properties
The one black liquor property that has the greatest effect on the evaporation plant odor
release is its sulfide concentration. Together with pH and temperature, the sulfide
concentration determines the quantity of H 2 S liberated from black liquor and eventually
vented. This relationship is shown in Figure 3.1, which also suggests that a black liquor with
a certain sulfidity will generate more H 2 S when it has a higher dry solids concentration (i.e.,
after more evaporation).
Bringing the sulfide concentration down to zero will obviously eliminate the H 2 S release
from the evaporation plant; this reduction in sulfide concentration also is accomplished by
oxidizing the weak, black liquor with air or oxygen before evaporation (see Chapter 9).
Black liquor from rotary drum vacuum filters will evolve less odor from diffusion washers,
since measurable oxidation of the sulfur content occurs in the filters (3).
3.1
-------
2
It
0
a. 0
a,
S
to TEMPERATURE = 90 -130°C
PRESSURE 600- 1200mm Hg
a ’ pH =12
• DIRECT MEASUREMENTS
0 INDIRECTLY CALCULATED
U )
N
I _________________________
o ib 20 30 40
Na 2 S IN LIQUOR, g No 2 S/Kg LIQUOR
FIGURE 3-1
VAPOR-LIQUOR EQUILIBRIUM FOR H 2 S OVER BLACK LIQUOR (1, 2)
The evaporation plant odor release will also depend on black liquor temperatures and pH,
but these factors are fixed by process conditions.
The amount of noncondensable odor gases from the evaporation plant will be greater if the
black liquor originates from hardwood digestion than from softwood (3).
3.2 Evaporator Types
There are different ways of evaporating spent black liquor indirectly with steam. The most
important ones, from an air pollution point of view, are discussed in the following sections.
3.2.1 Multiple-Effect Vacuum Evaporation
This is the dominant evaporation system. It can have many stages, usually 3 to 7. Each stage
may have multiple bodies. It is normally equipped with condensate flashing, liquor
prcheating, hot water generating, degassing, tail steam condensing and vacuum generating
systems. Falling film, rising film, or forced circulation evaporation can be used. The most
important features of condensate and gas treatment are shown in a generalized flow sheet in
Figure 3-2. The flows of condensate from the different stages are shown separately (flows
2-5 in Figure 3-2) to illustrate condensate treatment (section 5.3.2). Actually, they are
flashed in series through stages 3 to 5. Stages 2 to 5 are vented to the vacuum system and
stages 4 and 5 are equipped with liquor preheaters that serve as vent vapor condensers. The
3.2
-------
THICK WEAK SOAP SOAP TAIL CONDENSATES
LIQUOR LIQUOR
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY 0
FIGURE 3-2
* 18 WARM WATER
17 FRESHWATER
16 FRESHWATER
FRESH WATER
STEAM
VENT
STEAM
20
22
14
15
MULTI-EFFECT VACUUM EVAPORATION PLANT FLOW SHEET
-------
tail-end condensing systems is shown with a surface primary condenser, a generalized
secondary condenser (section 3.2.1.2), and a generalized vacuum device (section 3.2.1.3).
There are many different ways of feeding and circulating the liquor, but the one shown is
very common.
The odorous evaporator gases are usually released from the vacuum device (flow 9 in figure
3-2) through a duct from the hotwell.
The odorous evaporator condensates are usually released as secondary condensates (flows 2
to 5 in Figure 3.2) and as tail condensates from the hotwell (flows 6 to 8 in Figure 3-2). To
some extent, a trade-off is achieved between the release of odorous gas and condensate.
The range of flow and composition of the evaporator gases from the hotwell is given in
Tables 1-3 and 1-2, respectively. A discussion of some of the factors that influence the
gaseous emission and the condensates from the hotwell follows.
3.2.1.1 Evaporator Effect—Venting
The evaporation plant noncondensables generated in each effect must be vented so that they
cannot reduce the condensing heat transfer coefficient nor increase corrosion inside the
plant. Venting can be done in several ways, but the two principal systems are single-stage
venting and two-stage venting.
Single-stage venting is usually done by bleeding off a certain adjusted amount of vapor from
each effect to a central vent duct. In a vacuum evaporation plant, the vent duct usually goes
to the secondary condenser, then to the vacuum pump, and the gases are vented out of the
system from the hotwell or from the steam jet ejector.
When the weak liquor is fed into the evaporation plant, there is a large boil-off of
compounds with high vapor pressure in the feed effect (Number 3 in Figure 3-2). This vapor
condenses in the next stage leaving large amounts of noncondensable vapor in that vapor
space. Even in the next liquor evaporation.stage, boil-off is substantial, and noncondensables
are carried over to the following vapor space.
The application of two-stage venting can be advantageous. The vapor spaces of the first one
or two stages arc vented following the liquor feed stage. These vapor flows amount to 5 to
15 percent of the total vapor flow and are vented to separate condensers. These condensers
can serve as preheaters for Liquor or water and will, in their turn, be vented to the common
duet leading to the vacuum system. In this way the heat transfer coefficients of the
evaporator surfaces stay high, and the vapor is condensed in two fractions (for instance,
flows 5a and Sb in Figure 3-2). The fraction “a” will be large (say 85 percent of the vapor),
but will contain only around 40 percent of the biochemical oxygen demand (BOD) and
3-4
-------
sulfur compounds. The fraction “b” will be small (about 15 percent of the total), but will
contain the remaining 60 percent of the BOO and sulfur compounds. This will greatly
facilitate subsequent treatment of the condensates (section 5.3.2)
3.2.1.2 Evaporator Condensers
Evaporator condensers are mainly of two types, surface condensers and direct eon lact
condensers. Both will also be barometric condensers (i.e., they are equipped with barometric
legs or tail pipes).
The direct contact condenser, shown in Figure 3-3, is small in size, efficient, hard to plug or
foul, and inexpensive. It will, however, add water to the condensates and more
noncondensable gases will dissolve in this water depending on the water temperature.
Therefore, there is a larger volume of contaminated water to treat at correspondingly larger
costs (section 5.3.1). High fresh water temperatures in the summer will require large water
8 FRESH WATER
LAST
STAGE
VENT
VAPOR 1
WATER RING
i VACUUM PUMP
COND. 2
SECONDARY
CONDENSATE
POINTS OF POSSIBLE ORDER RELEASE ARE ENCIRCLED BY (3
FIGURE 3-3
EVAPORATION PLANT DIRECT CONTACT CONDENSER WITH
WATER RING PUMP
DIRECT CONTACT
CONDENSER
TAIL
CONDENSATE
3.5
-------
flows to maintain sufficiently low pressure in the evaporator plant tail end. On the other
hand, there will be less noncondensable gases to treat.
One interesting variation of a water ring pump circuit of an indirectly cooled condenser is
one in which white liquor is circulated instead of contaminated condensates. The cooled
white liquor scrubs the noneondensable gases and absorbs H 2 S and Cl-I 3 SH, first in the jet
condenser, and then in the water ring vacuum pump. Fresh white liquor is supplied, and the
overflow is returned to the process. Using this circuit saves the investment in an entire
scrubber (4).
3.2.1.2 Surface Condensers
The surface condenser has some advantages over the direct contact condenser. First, it does
not generate more condensate than the vapor condensed. Furthermore, it can be used to
produce clean warm water. It is also easier to control. Condensation can be controlled at the
desired degree of subeooling of the condensate. Controlled temperature subeooling is used
quite efficiently when the condensation is divided between one primary and one secondary
surface condenser (Figure 3.4). One can design and dimension the condenser and control the
cooling water flows so finely that about 85 percent of the vapor from the last stage will
condense in the primary condenser without any subcooling anywhere on the condenser
surfaces. Lack of subcooling will allow mmtmum dissolving of noncondensable gases and
low boiling organics in the condensate, and the condensate will be quite low in sulfur and
BOD and can be reused without further treatment. The remaining 15 percent of the vapor is
sent to the secondary condenser, where it is condensed and subcooled as much as feasible.
In this way, a substantial enrichment of noneondensable sulfur compounds and low boiling
organies occurs in this condensate, which may then be treated (e.g., by steam stripping). By
applying the concept of partial condensation in two steps (section 3.2.1.1), one condensate
stream can be divided into a large, rather clean part and a small, highly contaminated part.
The small part can be treated at significantly reduced cost. The importance of this design
concept will become more evident in the discussion of condensate treatment (see section
5.3.2).
3.2.1.3 Evaporator Vacuum Pumps
Two main types of vacuum devices arc in common usc, the water ring type of vacuum
pump, which is typical in Scandinavian countries, and the steam jet ejector type, which is
typical in North America. In some installations other types are used, such as water jet
ejector vacuum pumps. All vacuum pumps that allow noncondcnsable gases to contact fresh
cooling water will shift part of the odorous components from the hotwell vent gases to the
hotwell condensates.
3 -6
-------
9 WARM WATER
8 FRESH WATER
LAST
STAGE r* ®vENT
VAPOR 1
WATER RING
VACUMM PUMP
COND. 2
SECONDARY SURFACE SECONDARY SURFACE
CONDENSATE CONDENSATE & PUMP CONDENSATE
POINTS OF POSSIBLE ORDER RELEASE ARE ENCIRCLED BY (3
FIGURE 3-4
EVAPORATION PLANT TWO STAGE SURFACE CONDENSER
WITH WATER RING PUMP
3.2.1.4 Evaporator Flash Steam Feed
One source of primary steam for the evaporation plant is the secondary flash tank of a
continuous digester. Depending on the flash steam capacity and its pressure and
temperature, it may mect all or part of the steam demand of a particular evaporation plant.
In thc latter case, it may be supplied to the first evaporator effcct or to a alater one. From
an air and water pollution point of view, the important thing is that all noncondensable and
low boiling compounds flashed from the spend liquor in the flash tank pass over to the
evaporation plant and add to the noncondensable gas and contaminated condensate released
there. This release will be richer in organic compounds and turpentine than ordinarily is thc
situation in the evaporation plant. Thus, sulfur recovery with white liquor scrubbing will be
less effective, and more caution must be observed when treating the noncondensable gases
(sections 4.1.3 and 4.2.3).
3-7
-------
3.2.1.5 Condition of the Evaporator
According to Table 1-2, the amount of gases from the hotwell may vary between 0.3 and 12
m 3 per metric ton of pulp (10 and 400 ft 3 /ton). A reasonable value for softwood would be
around 1 m 3 of noneondensable gases per metric ton of pulp (32 ft 3 /ton). if much larger
values occur, the vacuum portions of the evaporation plant should be inspected for air leaks.
3.2.2 Other Evaporation Types
3.2.2.!. Multiple-Effect Back Pressure Evaporation Plant
The multiple.effeet back pressure system (Figure 3-5) is similar to the vacuum evaporation
plant, but uses a higher feed steam and tail steam pressure.
The tail steam is then used in some other process in the mill. Most of the noncondensable
gases, both sulfurous and organic, will be carried with the tail steam and vented elsewhere,
but can be collected for treatment at that point. The higher pressures in this evaporator
stage entail higher temperatures that cause more noneondensables to evaporate and also
cause a greater risk of liquor burnout on hot spots. Because of this risk, this type of
evaporator is currently not in widespread usage.
3.2.2.2 Flash Evaporation Plant
The flash evaporation column (Figure 3-6) is a fairly recent type of evaporator for black
liquor evaporation, in the column, the liquor is passed down and flashed in stages under
decreasing pressure. The flash steam is used to preheat the liquor in stages on its way up to a
new cycle of flashing. The evaporator includes a vacuum condensation system and operates
with complete crosscurrent flow. The gases vented from the vacuum system are very similar
to those from an ordinary multistage vacuum evaporation plant.
3.2.2.3 Thermoeompressor Evaporation Plant
A thermocompressor evaporation plant (Figure 3.7) is a single-stage evaporator with
multiple bodies, through which the liquor is passed in stages. Pressures and temperatures are
rather uniform over the whole plant, the compressor furnishing the difference of about 15°
C (27° F). All stages are vented through a liquor preheater. Vented gases are similar to those
of a back-pressure evaporation plant.
3.3 Evaporator Gas Scrubbing
Although noneondensable gas treatment will be extensively reviewed in Chapter 4,
evaporator gas scrubbing is discussed here because this equipment can be integrated with the
evaporation plant in different ways.
3-8
-------
VENT
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY Q
19 BACK PRESSURE
STEAM
SECONDARY CONDENSATES
63 WEAK LIQUOR
..-* ®THICKLIQUOR
FIGURE 3-5
MULTIPLE EFFECT BACK PRESSURE EVAPORATION PLANT FLOW SHEET
4
I I I
I I
STEAM
20
PREHEATER
-------
20 STEAM
22 CONDENSATES
VENT
THICK LIQUOR
WEAK LIQUOR
. © SECONDARY CONDENSATES
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY
FIGURE 3-6
FLASH EVAPORATION PLANT FLOW SHEET
Evaporator gas scrubbing with white liquor in a direct contact condenser of vacuum pump
was discussed previously (see section 3.2.1.2). It is also possible to install a scrubber after
the condenser and before the vacuum pump. The Venemark-design white liquor scrubber
has becn used in such an application (5). A small scrubber can be installed over the hotwell.
Such a scrubber design is shown in Figure 3-8. Noncondensable gases pass through the
packed column which is washed with white liquor sprays.
3-10
-------
COMPRESSOR
21 POWER
20 MAKE-UP STEAM
( j THICK LIQUOR
VENT
WEAK LIQUOR
SECONDARY
CONDENSATES
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY Q
FIGURE 3-7
MULTIPLE EFFECT SINGLE STAGE THE RMOCOMPRESSOR
EVAPORATiON PLANT FLOW SHEET
3.4 General Evaporator Air Pollution Abatement Programs
General programs for abating the air pollution of an evaporation plant include suitable
combinations of the following actions:
1. Checking of plant condition, in general, especially for possible leaks;
2. Installing condensers for two-stage venting;
3. installing surface condensers for two-stage condensing;
4. Collecting contaminated condensates for treatment, such as steam stripping;
5. installing a direct contact condenser or using vacuum pump scrubbing with white
liquor;
6. Installing a suitable white liquor scrubber for hotwcll gases;
PREHEATER
3-li
-------
SCRUBBED GASES
25mm (I”) RASCHIG RING
4--ID. 150mm (6”)
HOT WELL
MAX. FLOW
GASES
- 10 00m 3 1 hr
(590 ft 3 /m n.)
COOLED WHfTE LIQUOR
MAX. FLOW - 50 m 3 /hr
(30tt 3 /min.)
FIGURE 3-8
HOTWELL GAS SCRUBBER FOR 100 METRIC TONS PER HOUR (440 GPM)
EVAPORATION PLANT FOR H 2 S — SEPARATION OF 95% OR MORE
a
I D. 150 mm (6”)
I.D. 600mm (24”)
MIST ELIMINATOR
I.D. 50mm
(2”)
WHITE LIQUOR
RETURN
3-12
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7. Collecting hotwetl gases and incinerating them; and
8. Installing a weak black liquor oxidation system.
3.5 In-Plant Controls
Malodorous sulfur gases can be released from black liquor during multiple-effect evaporation
by liquor heating and by the stripping action of the steam (6). The H 2 S release during
multiple-effect evaporation is influenced by the inlet Na 2 S concentration, the black liquor
pH, and the type and degree of treatment in the weak black liquor upstream of the
evaporators (7). The potential liberation of organic sulfur compounds is influenced by their
concentration in the incoming black liquor which varies with wood species, the liquor
temperature, and the degree and type of treatment upstream of the multiple-effect
evaporators.
Major variables upstream of the multiple-effect evaporators which can affect sulfur gas
emissions include the wood species pulped, the pulping conditions, the type of pulp
washing, and the possible use of weak black liquor oxidation.
The organic sulfur emissions from pulping certain hardwood species are greater than for
most softwoods, particularly at high white liquor sulfidity levels. Vacuum drum washing of
pulp results in a stripping or organic sulfur compounds and in oxidation of a portion of the
Na 2 S. Diffusion washing is done in the absence of air and does not involve oxidation of the
Na 2 S or evolution of the organic sulfur gases. Inlet liquor concentrations of sulfur com-
pounds to the multiple-effect evaporators are expected to be higher following diffusion
washing of pulp than following drum washing.
The pH of the black liquor is an important variable affecting the liberation of H 2 S and to a
lesser extent Cl-I 3 SH. Both gases are slightly acidic in nature, with greater ionic dissociation
in aqueous solution favored by increased pH. Increasing the black liquor pH above 12.0
helps to reduce H 2 S emissions, lignin precipitation as a cause of evaporator plugging, and
the tendency for evaporator scaling and corrosion (8) (9). Addition of caustic soda to weak
black liquor in controlled quantities can raise the pH to the required levels.
Weak black liquor oxidation with either air or oxygen can reduce sulfur gas emissions from
multiple-effect evaporator noncondensable gases. Reid (10) and Galeano (11) report
reductions in H 2 S emissions of 70 and 99 percent from multiple-effect evaporators after
weak black liquor oxidation wjth air and oxygen, respectively.
Malodorous sulfur compounds emitted from the black liquor during multiple-effect
evaporation must end up in either the noncondensable gas stream or the condensate liquid.
The type of condenser employed has a definite effect on the distribution of sulfur
compounds between these two. Because of its scrubbing action, the use of the barometric
3-13
-------
jet condenser results in a greater portion of the sulfur compound emissions ending up in the
condensate rather than in the gas stream, as shown in Table 3-1 (6).
TABLE 3.1
EFFECT OF CONDENSER TYPE ON REDUCED SULFUR GAS EMISSION
FROM EVAPORATOR NONCONDENSABLE GASES (6)*
Condenser Type H 2 S CH 3 SH CH 3 SCH 3 CH 3 SSCH 3 Total
kg S/t kg Sit kg S/t kg Sit kg S/t
Surface 2.28 0.49 0.09 0.21 3.07
Barometric 0.06 0.07 0.05 0.01 0.19
lb S/ton lb S/ton lb S/ton lb S/ton lb S/ton
Surface 4.50 0.97 0.18 0.42 6.07
Barometric 0.12 0.13 0.10 0.02 0.37
*kg S/t = kilograms of sulfur per metric ton of air dried pulp
lb S/ton = pounds of sulfur per short ton of air dried pulp
3.6 References
1. Venemark, E., Black Liquor Evaporation, Part 2. Svensk Paperstidning, 61 (20):
881-887, October 31, 1958. (Swedish).
2. Arhippainen, B. and Jungerstam, B. Kraft Liquor Evaporation. In: Proceedings of the
Symposium on Recovery of Pulping Chemicals. Helsinki, Finland, May 13-17, 1968.
Finnish Pulp and Paper Research Institute and EKONO Oy, Helsinki, Finland, 1969, p.
132.
3. Sarkanen, K. V., Hrutfiorod, B. F., Johanson, L. N., and Gardner, H. S., Kraft Odor.
Tappi, 53: 776-783, May 1970.
4. Rönnholm, A. A. R., Reducing Evaporation Plant Pollution and its Treatment. Paperi
ja Puu, 54 (11): 715-730, 1972.
5. Swedish patent 226 789, Stockholm, Sweden.
3 14
-------
6. Hendrickson, E. R., Robertson, J. E., and Koogler, J. B., Control of Atmospheric
Emissions in the Wood Pulping Industry, Vols. I, II, III. Final Report Contract No.
CPA 22-69-18, U.S. Department of Health, Education, and Welfare, National Air
Pollution Control Administration, Raleigh, North Carolina, March 15, 1970.
7. Douglass, 1. B., Sources of Odor in the Kraft Process: Odor Formation in Black Liquor
Multiple Effect Evaporators. Tappi, 52: 1738-1741, September 1969.
8. Berry, L. R., Black Liquor Scaling in Multiple Effect Evaporators. Tappi, 49: 68A-71A,
April 1966.
9. Cry, M. E., and Harper, A. M., Multiple Effect Evaporator Project. Pulp and Paper
Magazine of Canada, 61: T247-T249, April L960.
10. Reid, H. A., The Odor Pro blem at Maryvale. Appitta, 3(2):479-500, December 1949.
11. Galcano, S. F., and Amsdcn, C. D., Oxidation of Kraft Weak Black Liquor with
Molecular Oxygen. Tappi, 53: 2142-2146, November 1970.
3-15
-------
CHAPTER 4
NONCONDENSABLE GAS TREATMENT
Digester and evaporator noncondensable gases characteristically have relatively low volume
flow rates and high malodorous sulfur compounds concentrations. Organic sulfur
compounds, such as CH 3 SH, CH 3 SCH 3 , and CH 3 SSCH 3 , are emitted from digesters in
varying quantities from 0.25 to 2.5 kg of sulfur per metric ton of pulp (0.5-5.0 lb sulfur per
ton of pulp). The primary sulfur compounds that are emitted from multiple-effect evapora-
tors are H 2 S and CH 3 SH in quantities ranging from 0.1 to 1.5 kg of sulfur per metric ton of
pulp (0.2-3.0 lb sulfur per ton of pulp). Unless properly controlled, these gas streams can
cause intense odor pollution at low elevations near the mill.
4.1 Gas Stream Characteristics
The noncondensable gases from digesters and evaporators are relatively high concentration,
low volume streams. Their odor levels are easily reduced by chemical or thermal oxidation.
Major parameters of the gas stream which must be considered in the design of
noncondensable gas handling and treatment systems include temperature, moisture content,
flow rate and variability, sulfur gas concentrations, organic material concentrations, and
flammability limits.
4.1.1 Process Sources
The major noncondensable gas streams collected and treated in kraft pulp mills are the
blow and relief gases from batch and continuous digesters, and the hotwell and condenser
vents from multiple-effect evaporators. The gas collection systems for kraft pulp mills must
be individually designed to connect all the various gas sources.
Batch digesters normally make up the largest single volume source and give rise to the
greatest variations in flow rates of any of the noncondensable gas streams. Batch digesters
are normally vented through the relief system at the condenser vent and the turpentine
decanter vent. The relief gases are low volume gas streams that flow on a more or less
continuous basis during the 3- to 5-hour cooking period. The relief gases normally contain
large amounts of terpenes, in addition to the sulfur compounds released during the cook,
and can pose an explosion hazard.
The batch digester is normally vented to the blow tank at the end of each kraft cook over a
10- to 20-minute period, with the resultant release of large quantities of steam, inert gas,
organic compounds, and malodorous sulfur compounds. The volume of gas released depends
on the digester volume, the gas temperature and moisture content, the degree of vapor
4-1
-------
condensation effected, the amount of air volume in the system, and the amount of inert
gases present in the digester. The volume of gas to be handled occurs in a large single surge
at the end of each cook. This feature normally requires the use of a flow equalization device
to avoid upsetting the operation of the burning device and to assure maximum thermal
operating efficiency. The use of sufficiently-sized blow heat condensers may obviate the
need for an equalization device.
Continuous digesters normally are nearly constant flow rate devices (except during periods
of process upset), requiring no flow equalization devices. Continuous digesters arc normally
vented at the steaming vessel relief, at both condenser and turpentine decanter vents, from
the blow tank vent after the condenser system and sometimes from the top of the digester
unit itself. The exact venting system arrangement varies among individual digesters. The
total amounts of gas and the terpenes and organic sulfur compounds liberated are not
normally as large for continuous as for batch digesters. Also, the amounts are somewhat
dependent on the digester blow temperatures. The result is that the point of emission for
these materials is often transferred from the digester system to the brown stock washer and
multiple-effect evaporator sections.
Noncondensable gases from multiple-effect evaporators differ in character from vent gases
from digesters in that they contain smaller amounts of terpenes and organic sulfur
compounds. The gases are made up primarily of H 2 S and CH 3 SH liberated from the black
liquor during the evaporation process. Evaporator noncondensable gases are normally
collected from the hotwell and condenser vents and can vary between mills. Normally, the
noncondensable gas flow rates are considerably larger from indirect contact (surface)
condensers than from direct contact (jet) condensers.
4.1.2 Flow Rates
The gas flow rates for individual process streams are subject to wide variations among
individual mills, depending on production rate, process operating variables, and the degree
of condensation for heat recovery. The single most important flow rate variable for
noncondensable gas streams is the peak flow rate for batch digester blow gases, particularly
during periods of condenser malfunction. Some type of pressure relief system for batch
digester blow gas systems should be provided during periods of condenser malfunction,
which are usually indicated by gas temperatures well above normal. A summary of one
mill’s gas flow rates from a batch digester to a gas holder under conditions of average flow,
maximum flow, and condenser upset is presented in Table 4-1.
Major variables affecting the overall gas flow rates from noncondensable gas streams are the
process unit type and operating conditions, the production capacity for the particular unit
from which the stream is vented, and the degree and type of vapor condensation employed.
A summary of typical ranges for noncondensable gas flow rates is presented in Table 4-2.
4-2
-------
TABLE 4- i
GAS FLOW’ RATES FROM A BATCH
DIGESTER* (1)**
Operating Digester Blow Pulp Produc-
Condition Basis tion Basis
m 3 /blow (ft 3 /blow) m 3 /t (ft 3 /ton)
Average 45 (1,600) 4 (128)
Maximum 113 (4,000) 10 (320)
Upset 3,540 (125,000) 320 (10,250)
*Gas flows at actual stack conditions.
**Pulp cooking capacity is 11.4 t/cook (12.6 ton/cook).
TABLE 4-2
TYPICAL RANGES IN DIGESTER AND
EVAPORATOR NONCONDENSABLE GAS FLOW
RATES
Source Process Stream Flow Rate
m 3 /t
(ft 3 /ton)
Digester Batch Blow 475-6,350
(15,200-203,500)
Batch Relief 0.3-95
(10-3040)
Contthuous 0.6-6
(20-200)
Evaporator Surface Condenser 0.6-13
(20-420)
Jet Condenser 0.3-3
(10-100)
4-3
-------
Flow rates of noncondensable gas streams to subsequent treatment devices following flow
equalization normally range from 5 to 35 m 3 /h (3 to 21 cfm) with average values of 10 to
15 m 3 /h (6 to 9 cfm) reported (2). Higher flow rates are reported when there is insufficient
heat exchanger condensation capacity, when flow equalization is not employed, or when
there is significant air leakage into the gas handling system.
4.1.3 Safety Considerations
The design and operation of a system for thermal oxidation of the noncondensable gases
require measures to prevent explosions. Four factors that must be considered are variations
in gas flow rate, passage of entrained moisture droplets into the burning device,
flammability limits of sulfur and organic compounds in the gas stream in the inlet piping,
and flame propagation speeds as compared to gas flow velocities in the piping. Of particular
importance is the presence of terpene compounds in the noncondensable gas handling and
burning systems.
Excessive variations in flow rates of noncondensable gases entering the air inlet of a burning
device can blow out the flame. An additional factor which should be considered is that the
net retention time of the gas stream in the burning device is reduced during flow surges.
Solutions to this problem are to employ large surge tanks with large condenser capacities, to
use flow equalization devices, and to provide pressure relief vent systems.
Entrained moisture droplets can create hazards when entering the combustion zone. First,
the droplets can cool the flame and may even extinguish it. Second, the evaporation of
water can result in a large increase in gas volume as the water changes to steam. These surges
in volume can cause an unstable operation in the burning device and lead to a flameout. A
flameout may allow an explosive mixture of flammable materials to accumulate in the
combustion unit, resulting in an explosion when the system is reignited.
Explosive limits must be considered in thc design of noncondensable gas handling systems,
both before and after dilution with primary air. A summary of flammability limits for
materials commonly present in noncondensable gas streams is listed in Table 4-3 (3).
Terpenes have the lowest explosive limits of any of the compounds listed and are normally
the most critical component in noncondensablc gas streams for purposes of design for
explosive safety.
Work by Ginodman (4), Coleman (5), and Del-laas and Hansen (6) indicates that it is
necessary to dilute noncondensable gas streams by a sufficient amount in the primary air
inlet before they enter combustion devices. Dilution must occur at a fast enough rate so that
the nearly oxygen-free noncondensable gas stream can pass the lower explosive limits of the
most critical material (normally terpcncs) without an explosion. A summary of the
44
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TABLE 4-3
FLAMMABILITY LIMITS IN AIR FOR COMPOUNDS
PRESENT IN KRAFT NONCONDENSABLE GAS
STREAM (3)
Flammability Limits
Material Lower Upper
Concentration, % by vol.
H 2 S 4.3 45.0
CH 3 SH 3.9 21.8
CH 3 SCH 3 2.2 19.7
Terpene 0.8
necessary dilution requirements for batch digester relief gas alone, and for combined
digester blow and relief gas is presented in Table 44.
The primary danger of explosions exists in the primary air inlet of the combustion device
immediately following introduction of the noncondensablc gases. The presence of large
quantities of terpene compounds tends to make digester relief gases a greater explosive
hazard than blow gases.
TABLE 44
DILUTION REQUIREMENTS WITH AIR TO AVOID
EXPLOSIONS FOR DIGESTER NONCONDENSABLE
GAS STREAMS (4) (5) (6)
Gas Stream Volume Dilution Required
Air/Gas Ratio
Relief only 50/1
Relief & blow 20/1
An additional consideration regarding potential explosions involves flame propagation
speeds of air-gas mixtures. Data collected by Ghisoni (7) on flame propagation speeds for
air-mercaptan mixtures are listed in Table 4-5.
4.5
-------
TABLE 4-5
FLAi IE PROPAGATION SPEEDS FOR AIR-
MERCAPTAN MIXTURES (7)
Mercaptan Concentration Flame \‘elocity
% by Vol. rn/s (ft/sec)
18.9 0.55 (1.8)
22.8 0.46 (1.5)
23.1 0.40 (1.3)
23.7 0.37 (1.2)
25.5 0.18 (0.6)
25.7 0.15 (0.5)
Gas velocities in the noncondensable gas piping and the primary air inlet must be greater
than any flame propagation speeds to prevent damage to process units from possible
explosions. Maintenance of gas velocities of at least 1 rn/s (3 ft/see) at all times and the use
of flame arrester devices in the noncondensable gas line should minimize the danger of
explosions from excessive flame speeds.
4.2 Gas Handling Systems
The major parts of a noncondensable gas handling system are condensation, flow
equalization, liquid scrubbing, piping, safety control, and air inlet sections.
Removing moisture and turpentine upstream of the burning device is particularly important
in preventing possible flameouts. Sulfur compounds can be absorbed into alkaline liquids for
white liquor makeup. The noncondensable gas piping system must be designed with safety
devices to prevent explosions and must also have design features to permit the rapid dilution
of the noncondensable gas stream before it enters the combustion unit.
4.2.1 Vapor Condensation
The primary purpose of the condensers in digester and evaporator gas handling systems is
heat recovery from the gas stream by water vapor condensation.
The condenser systems also remove a portion of the organic vapors, such as terpcncs and
sulfur compounds. The vapors so removed also require, in turn, treatment of the condensate
waters for odor removal. The terpenes can be recovered by flotation and decantation from
the condensed water. After collection, they can be sold as byproduct turpentine or used as
an auxiliary fuel in the lime kiln.
4-6
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The condensers also serve to reducc the volumc of gas to be treated by cooling the gas
stream and thereby condensing out a portion of the contained water vapor. Cooling the gas
stream to between 75 and 85° C (167 and 185° F) is normally desirable to remove the
major portion of the water and reduce the gas volLime. Where no flow equalization devices
are employed, the use of large heat transfer surface areas in condensers is particularly
important for blow gas streams from batch digesters. Surface condensers are more desirable
from a water pollution reduction standpoint than comparable spray or jet condensers since
they produce a lower volume of more highly concentrated contaminated condensate. This
type of condenser is normally larger in surface area and, therefore, has a higher capital cost.
4.2.2 Flow Equalization
The two major devices employed for noncondensable gas flow equalization are the
vaporsphere and the floating cover gas holder. DeHaas and Hansen (6) describe the use of a
vaporsphcre for collection of batch digester noncondcnsablc gases. The vaporsphere is a
spherical device with a flexible fabric diaphragm attached around the epicenter of the sphere
as shown in Figure 4.1 (2). The diaphragm consists of a mylar film sandwiched between two
layers of cotton canvas that can be fabricated by local tent and awning manufacturers. The
useful life of such mylar canvas diaphragms is as long as 27 months (8).
Dia phragm
FIGURE 4-1
VAPORSPHERE FLOW EQUALIZATION GAS HOLDERS (2)
Vacuum
Pressure
Relief
Blow
Relief
And
Sliding
Weight
Flow
Control
Gases
4.7
-------
Several operating and safety features are necessary to assure safe and reliable operation of
the vaporsphere. The system contains a counterweight connected to the diaphragm which
moves up and down as the gas flows in and out of the vaporsphere. The diaphragm is
weighted to provide a slight positive pressure on the system at all times. Automatic flow
controls on the system are used to prevent damage to the diaphragm and to prevent
excessive air leakage. A pressure and vacuum relief system prevents damage to the
diaphragm from either excessive gas flows from condenser malfunctions or from excessive
suctions. The condenser section and all associated piping to the vaporsphere must he sealed
to prevent the possibility of air leakage and a potential explosive mixture from forming (6).
Also, the inlet and outlet gas streams must be vented separately for the full flow
equalization effect to be obtained from the vaporsphere.
Floating cover gas holders also are used for flow equalization in kraft pulp mills. They
consist of two vertical cylindrical tanks with the upper located inside the lower and a water
seal to prevent gas leakage, as shown in Figure 4-2. Gas enters the inside cover cylinder
above the water seal, causing the cover to be displaced upward when gas enters and
downward as it exits. Gas is withdrawn from a separate exit pipe to achieve the full degree
of flow equalization for the system.
The floating cover gas holder must have several operating and safety features to assure its
reliable operation. A -pressure and vacuum relief system is added to prevent the shell
cylinder from being damaged or dislodged by excessive pressures or vacuums. These are
Vacuum
Pressure
Relief
System
TI
-J
FIGURE 4-2
FLOATING COVER FLOW EQUALIZATION GAS HOLDERS (2)
Blow By- pass
Vent
4-8
-------
connected to a hydrostatic water seal with an overflow to the drain to prevent the water scat
from being lost. A bypass vent allows venting of excessive gas surges such as those caused by
condenser malfunctions. Ping pong balls are placed in the space above the water seal
between the shell and cover cylinders to maintain alignment and smooth operation.
Vaporsphcrc and floaLing cover gas holders can normally be constructed of mild steel.
Capital costs for these systems, including alt appurtenances, are about S20,000 to $35,000.
A summary of the dimensions and construction materials for flow eqLialization devices is
presented in Table 4-6 (2).
TABLE 4-6
DIMENSIONS OF FLOW EQUALIZATION DEVICES* IN
KRAFT NONCONDENSABLE GAS HANDLING
SYSTEMS (2)
Type of Gas Holding Dimensions
Mill Unit** Diameter Height Volume
m (ft) m (ft) m 3 (cu ft)
A VS 8.2 (27) .. - - 170 (6,000)
B VS 8.2 (27) .- -. 170 (6,000)
G VS 8.2 (27) - - 170 (6,000)
H VS 6.6 (21) -. .. 142 (5,000)
E FC 8.5 (28) 4.6 (15) 283 (10,000)
*Mild steci used at all five mills as material of construction.
= vaporsphere, FC = floating cover
Values for digester blow gas design flows for sizing flow equalization devices are reported by
Blosser and Cooper (2) and DeHaas and Hansen (6). A summary of digester blow gas volume
flows for varying operating conditions is listed in Table 4.7.
4.2.3 Liquid Scrubbing
Liquid scrubbing of the noncondensable gas stream is added for purposes of organic mist
removal, gas stream cooling, and sulfur recovery. DeHaas and Hansen (6) report that a liquid
scrubber is needed to prevent turpentine mist droplets from reaching the burning device and
causing periodic flameouts. The contact of the noncondensablc gas stream with sufficiently
cool scrubber liquor results in additional cooling of the gas stream, reduces its volume, and
removes additional water vapor.
4.9
-------
TABLE 4-7
BATCH DIGESTER BLOW GAS FLOW RATES FOR SIZING NONCONDENSABLE GAS FLOW EQUALIZATION I)EVICES
Gas Flow/Digester Blow* Gas Flow/Unit Production*
Operating Condition Average Range Average Range
m 3 /blow (cu ft/blow) m 3 /t (Cu ft/ton)
Normal 42 (1,500) 14-113 (500-4,000) 4 (125) 2-5 (70-160)
Condenser Malfunction 425 (15,000) 285-570 (10,000-20,000) 47 (1,500) 3 L-62 (1,000.2,000)
No Heat Recovery 1,130 (40,000) 5,660-18,400 (200,000-650,000) 1,250 (40,000) 625-2,000 (20,000-64,000)
*Gas flows are at stack conditions.
-------
The use of alkaline scrubbing liquids, such as sodium hydroxide (NaOH) solution or white
liquor, results in removal of the acidic sulfur compounds, such as 11 2 S and CH 3 SH, from the
noncondensable gas for return to the chemical makeup system. It is not possible, however,
to remove significant quantities of organic sulfur compounds, such as GH 3 SCH 3 and
C l 3 SSCH 3 , from the noncondensable gas stream under normal circumstances. Reduced
sulfur gas emissions from digester gas sources can range from below 0.25 to above 2.5 kg per
metric ton of pulp (0.5 to 5.0 lb per ton of pulp); evaporator noncondensable gas sulfur
emissions are normally 0.025 to 0.25 kg p r metric ton of pulp (0.05 to 0.5 lb per ton of
pulp). Sulfur compounds from the digesters are primarily Gil 3 S d 3 and Gil 3 SSCH 3 ; while
sulfur compounds from the evaporators are primarily H 2 S with lesser amounts of CH 3 SH.
The alkaline scrubbers generally used are packed bed scrubbers employing a countercurrent
flow of liquid and gas. The usual packing for these devices is gravel, stone, or one-inch thick
packing rings (2). The liquid solution employed for scrubbing is either caustic soda or white
liquor for the return to the chemical makeup system, or water for subsequent discharge to
the sewer. A typical scrubber system is illustrated in Figure 4-3 (2).
FIGURE 4-3
PACKED BED SCRUBBER FOR NONCONDENSABLE
GAS HANDLING SYSTEM
ing
30”
4-11
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4.2.4 Piping Systems
Design of the piping for a noncondensabic gas handling system requires consideration of
such items as materials for construction, explosion hazard safety, and gas flo %’ pressure
drop. Noncondensable gas handling systems are normally constructed of mild steel, but 304
or 316 stainless steel has been used in some applications to inhibit corrosion. Constructing
noncondensable gas piping systems to obtain a minimum velocity of 1 rn/s (3 ft/scc) is
usually necessary to minimize the likelihood of flame propagation through the pipe. High
velocities also result in increased pressure drops through the piping, particularly for lines
longer than 30 m (100 ft). An auxiliary fan or a large diameter pipe should be installed to
minimize pressure drop across the system and to allow use of the inlet draft of combustion
air systems. Most mills have installed 10 cm (4 in) diameter pipe for noncondensable gas
piping systems; while 7 cm (3 in) and 20 cm (8 in) have been used in a few mills. The larger
diameter piping has found its greatest use at mills employing continuous digesters. A
summary of piping system gas velocities, pipe diameters, and materials of construction is
presented in Table 4-8.
TABLE 4-8
PIPING SYSTEMS FOR KRAFT NONCONDENSABLE GAS
HANDLING SYSTEMS (2)
Mill Diameter Gas Velocity Materials of Construction
cm (in) rn/s (ft/see)
A 10 (4) L5 (4.8) Mild steel
B 10 (4) 1.6 (5.3) Mild steel
C 20 (8) 3 (9.9) 316 Stainless
D 20 (8) 13 (41.7) 304 Stainless
E 10 (4) 1.7 (5.7) Mild steel
F 10 (4) — — Mild steel
C 20 (8) 7 (23.1) Mild steel
H 7 (3) 1 (3.3) Mild steel
4.2.5 Safety Considerations
In designing noncondensable gas handling systems, specific safety features arc needed to
assure a minimum explosion hazard, prevent liquid entrainment, and assure stable operation
of combustion devices. Condensate traps to remove water are placed at low points in the
piping system at intervals of approximately 15 m (50 ft) and just upstream of the air inlet.
A typical liquid condensate trap design is illustrated in Figure 4-4 (2). The packed bed
scnibbcrs also serve the function of moisture removal to prevent upsets of burner operation,
4.12
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Liquid
FIGURE 4-4
LIQUID CONDENSATE TRAP FOR NONCONDENSABLE
GAS HANDLING SYSTEM
or flameouts, by liquid droplet entry to the combustion zone, and to eliminate false signals
to the flameout control.
Flame arresters of the leaf or grid type are commonly added to the noncondensable gas line
to prevent the passage of any flame fronts to the process units. The flame arrester normally
is added to the noncondensable gas line immediately upstream at the point of introduction
to the primary air ducts, as shown in Figure 4-5 (2). One or more additional flame arresters
can be installed in long piping systems to gain further protection from possible explosions.
Two features are added to provide emergency venting of excess gas pressures during possible
explosions. Rupture discs are added to noncondensable gas lines at approximately 30 m
(100 ft) intervals. These devices have diaphragm discs set to explode at certain bursting
pressures. To cope with power failures, an emergency vent release is normally placed in the
noncondensable line connected to the flameout control for the combustion device. A
continuous recording device with input from an orifice meter provides a useful record of
system flow rate.
Design for safe confluence of the noncondensable gas stream with the primary air to the
combustion device requires consideration of inlet draft, gas velocity, and physical features.
A damper is normally placed in the primary air inlet upstream of the point of gas
introduction to provide an inlet vacuum of 2.5 to 7.5 cm H 2 0 (1 to 3 inches of water)
necessary to maintain gas flow. Otherwise, an auxiliary fan is needed in the noncondensable
gas handling system. The gas is added through a horizontal pipe placed across the primary
4-13
-------
FIGURE 4-5
SAFETY DEVICES FOR NONCONDENSABLE GAS HANDLING SYSTEMS
air duct with holes evenly spaced along the pipe on the downstream side to achieve even
distribution and rapid dilution of the gas to well below explosive limits. Gas velocities in the
primary air inlet duct after mixing with air are normally above 9mIs (30 ft/sec).
4.3 Gas Treatment Systems
The major techniques for treatment of malodorous sulfur gases to prevent their emission to
the atmosphere are thermal oxidation and liquid absorption. The major types of devices for
thermal incineration of noncondensable gases are the lime kiln and catalytic furnaces, but
limited use is also made of other combustion systems. The liquid solutions empLoyed to date
are acidic chlorination bleaching effluent and caustic solutions. Thermal oxidation in lime
kilns provides a positive means for destruction of malodorous sulfur gases. Liquid scrubbing
can also prove effective as an alternate treatment system or as a pretreatment technique for
safety considerations or for sulfur recovery.
4.3.1 Lime Kiln Incineration
Incineration of digester and evaporator noncondensable gases in lime kilns provides a positive
method for odor control, without the necessity of constructing an additional combustion
unit, and also allows heat recovery. The noncondensable gases normally are added to the
primary air inlet of the lime kiln at a dilution of at least 50 to 1 and an air inlet velocity of
Flow Record
And Control
Vent
Flame
Cut L :i
Control ‘
I I
I I
S I
Lime
Kiln
Damper
Flame
Primary
Arrester
Air Fan
4-14
-------
at least 9 rn/sec (30 ft/sec). A flameout control on the lime kiln is connected to a three-way
emergcncy bypass vent in case of power failure.
Gases are incinerated in the kiln at maximum temperatures of about 1200 to 14000 C (2200
to 2550° F) to achieve complete oxidation of the sulfur compounds present. The SO 2
formed is largely collected by the lime, as calcium sulfite (Ca 2 SO 3 ) instead of being emitted
to the atmosphere. To date, mills employing noncondensable gas incineration have not
observed any adverse effects on either .reburned lime quality or causticizing system
operation from the burning of these gases.
The basic layout of the burning system for incinerating noncondensable gases consists of the
collection piping and burning sections. Systems where batch digester gases are collected
require a flow equalization device or suitably large condenser capacity. The layout for a
typical unsteady state system for batch digester blow and relief gases is illustrated in Figure
4-6; a steady state system for a continuous digester is illustrated in Figure 4-7 (2). A
combination system for gases from both batch and continuous digesters, plus evaporator
gases, is illustrated in Figure 4-8. The noncondensable gas flow rates, their sources, and
burning devices employed for eight existing installations are presented in Table 4-9.
FIGURE 4-6
UNSTEADY STATE FLOW SYSTEM FOR BATCH DIGESTER NONCONDENSABLE
GAS INCINERATION (2)
Flow
Flow Record
& Control
Flame Out
Auxiliary
Fan
Tank
TraPS
‘-4— Fan
Condenser
Scrubber
Flame
ATreste rs
Effluent
4-15
-------
Cyclone or
Condensate Traps
STEADY STATE FLOW SYSTEM FOR CONTINUOUS DIGESTER
NONCONDENSABLE GAS INCINERATION (2)
TABLE 4 -9
GAS FLOW RATES TO BURNING DEVICES FROM NONCONDENSABLE GAS
HANDLING SYSTEMS (2)
Gas Flow Rate
m 3 /h (cfm)
Sources Included**
Burning Device
A
363
(400)
42
(25)
BD, TD
Lime Kiln
B
999
(1,100)
48
(28)
BD, EH, ID
Lime Kiln
C
499
(550)
357
(210)
CD
Lime Kiln
D
499
(550)
1,490
(875)
CD, CS
Lime Kiln
E
545
(600)
51
(30)
BD,CD,EH
LimeKilti
F
1,135
(1,250)
.-
--
BD, CD, EH
Lime Kiln
G
499
(550)
850
(500)
BD
Cat. Furn.
H
145
(160)
17
(10)
BD
Cat. Furn.
*Reported at stack conditions
* Sourcc Codc
BD . Batch digester (both blow and relief).
CD . Continuous digester (both blow and relief).
CS . Condensate stripping.
EH . Evaporator hotwell.
TD . Turpentine decanter.
Steamin
Flow Recorder
Flame Out
Control
Aeration
Fan
Rupture
Discs
Arresters
FIGURE 4-7
Mill Daily Pulp Production
t (ton)/day
416
-------
Gases From Multiple
Effect
Floating Cover
Gas Holder
- ‘ I
Blow & Relief
Gases From 4
Botch Digesters
8 A ; Heat
Accumulator
Defibrotor Off
Gases
Entrainment
Separator
Lime Ki lns
FIGURE 4-8
Fuel Oil
Tank
NONCONDENSABLE GAS INCINERATION SYSTEM
-------
4.3.2 Other Incineration Systems
Both catalytic and auxiliary furnaces can be used for incineration of noncondensable gases,
but they require a separate combustion unit and the heat cannot normally be recovered.
Coleman (5) and DcHaas and Hansen (6) report on the use of an auxiliary furnace for
combListion of noncondensable gases. The piping system is the same as for lime kilns. The
system burns the gases at 850 to 10000 C (1560 to 1830° F), hut does not recover any of
the heat produced. The use of the auxiliary furnace has been discontinued because of
control maintenance problems and unstable burner operation caused by water and
turpentine mist entrainment.
A second installation employs an auxiliary furnace for incineration of batch digester and
evaporator noncondensable gases without the use of a flow equalization device. The system
employs both blow heat condensers with large heat transfer surface areas and surge tanks of
large capacity. Large air dilution is required along with careful gas flow rate control; one
explosion with the system did occur (9).
Catalytic furnaces are employed for noncondensable gas incineration at two mills.
Noncondensable gases are diluted with air, as in lime kiln systems, and are incinerated at
400° C (750° F) in the presence of porcetam rods coated with an alumina-platinum catalyst.
Shortcomings of the systems include incomplete oxidation of organic sulfur compounds,
requirements for considerable maintenance of automatic controls, and frequent replacement
of the catalyst cells if these are allowed to contact water droplets (2).
Other methods of incineration of incineration of noncondensable gases have been reported.
Chisoni (7) reports on the burning of digester and evaporator noncondensable gases in a
natural gas-fired power boiler; Lindberg (10) describes the addition of noncondensable gases
to the primary air inlet of a kraft recovery furnace. Adams (11) describes the use of
recovery furnaces, auxiliary furnaces, and waste wood burners for incineration of
noncondensable gases.
4.3.3 Liquid Scrubbing Systems
The two major types of scrubber liquids employed, to date, are caustic soda and acidic
chlorination bleaching effluent. Chase (12) describes the use of alkaline scrubbing to remove
1-12 S from evaporator noncondensable gases for return to the chemicaL makeup system.
Alkaline solutions do not have any great affinity for nonpolar organic sulfur gases and do
not achieve any significant removal of them.
Morrison (13) describes a system where batch digester blow and relief gases are incinerated
in a lime kiln. A backup system employs addition of the digester gases to the dropleg of the
acidic chlorination bleaching stage when the burning system is not in use. The excess
4-18
-------
chlorine convcrts the sulfur compounds Lo elemental sulfur, sulfonvi chlorides, sulfoxy
compounds, and oxidized Lerpenes. ftc system appears to oxidize the sulfur compounds
sufficiently to prevent their release to the atmosphere.
4.3.4 Standby Systems
The maximum degree of control of non-condensable gases is achieved by providing
alternative combustion units for Lhe incineration of these gases. If lime kiln incineration is
practiced, the alternative could he routing the gases to a second lime kiln, a boiler or a
furnace. Such a system offers the possibility for continued noncondensable gas control
during periods when the primary incineration device is out of service.
4.4 System Economics
4.4. I Capital Costs
The capital costs for a noncondensablc gas handling system depend on these parameters:
1. Diameter, length, and materials of piping;
2. Number, type, and size of safety appurtenances; an(l
3. The possible usc of flow equalization gas holders and auxiliary fans.
The flow equalization gas holder is basically a fixed cost item of $25,000 to $50,000, based
on type and size of unit, and is independent of the amount of pulp production. Piping and
safety devices add an additional $25,000 to $75,000 or more to the cost of the system. The
use of auxiliary furnaces instead of lime kilns increases the cost of the system. A summary
of installed capital costs for noncondensable gas handling systems is presented in Table 4.10.
4.4.2 Operating Costs
The major operating cost variables for noncondensable gas handling systems are those for
the additional electric power for increased kiln air draft or auxiliary fans, and for system
maintenance. Combustion of noncondensable gases in the lime kiln achieves additional fuel
savings by reducing oil or gas requirements. In addition, turpentine can be burned in the
lime kiln to reduce fuel requirements. i ’lorrison (13) reports system maintenance costs of
$2,100 per year; while fuel savings of $1 ,200 result from burning noncondensable gases in
the lime kiln. Operating costs for noncondensable gas incineration are estimated to be S0.01
to $0.05 per metric ton of pulp produced, with an average of SO.03 per metric ton (S0.01 to
0.05 per short ton, avg. $0.03/short ton).
4 19
-------
TABLE 4-10
CAPITAL COSTS FOR INSTALLED NON-
CONDENSABLE GAS HANDLING SYSTEMS
Burning Device
or System Installed Capital Cost
S/daily t (S/daily ton)
Lime Kiln 110-165 (100-150)
Auxiliary Furnace 165-220 (150-200)
Catalytic Furnace 192-275 (1 75-250)
4.5 References
I. Personal communication with Mr. Andrew F. Reese, Fibreboard Corporatrnn, Antioch,
California, February 1 970.
2. Blosser, R. 0., and Cooper, H. B. H., Current Practices in Thermal Oxidation of
Noncondensable Gases in the Kraft Industry. Atmospheric Pollution Technical Bulletin
No. 34. National Council of the Paper Industry for Air and Stream Improvement, Inc.,
New York, New York, November 1967.
3. Perry, J. H. (ed.). Chemical Engineers Handbook, 3rd. edition. New York.
McGraw-I-till Book Company, 1950. p. 1585-1586.
4. Ginodman, G. M., Parification of Waste Streams from Sulfate Cellulose illanafacture.
Bu mazh naya Promysh lennost, (Moscow) 22 (7): 16.22, November-December 1947.
5. Coleman, A. A., The Combustion of Noncondensable Blow and Relief Cases in the
Lime Kiln. Tappi, 41: 166A .168A, October 1958.
6. DeI-laas, G. C., and Hansen, C. A., The Abatement of Kraft Mill Odors by Burning.
Ta pp’ 38:732.738, December 1955.
7. Chisoni, P., Eluninatton of Odors in a Sulfate Pulp Mill. l’appi, 37:201-205, May 1955.
8. Hansen, C. A., Odor and Pnllout Control in a Kraft Pulp Mill. Journal of Air Pollution
Control Association, 12:409.412, September 1962.
4-20
-------
9. Personal communication with Mr. Dwayne j. Clark, Simpson Lee Paper Company,
Everett, Washington, 1972.
] 0. Linclbcrg, S., I-low One Swedish Mill Destroys Air and J1’ater Pollutants. Pulp and
Paper, 41 (3):35.39, January ]6, 1967.
Li. Adams, I). F., A Survey of European Kraft Mill Odor Iteduction Systems. Tappi,
48:83A-85A, May 1965.
12. Chase, S. J., Control of Air Pollution at the Champion Paper and Fibre Company.
(Paper Read to the Scmi.Annual Technical Meeting of the Air Pollution Control
Association. 1-louston, l)cccmbcr 3, 1956.)
13. Morrison, J. L, Collection and Combustion of Noncondensable l)tgester and
Evaporator Gases. Tappi, 52:2300-2301, Dcccmher 1969.
4.21
-------
CHAPTER 5
CONDENSATE TREATMENT
Besides noncondensable air polluting gases the kraft process produces condensates that are
contaminated with different compounds. These compounds can generate both air and water
pollution.
A general course of action is to decrease the amount of condensates, collect them, and reuse
them; if they arc not reusable, the procedure then is to strip them of their contaminating
compounds and oxidize these compounds to less harmful forms.
5. I Condensate Components
The kraft process produces two main condensates, namely digester and evuporaLor
condensates. Both contain compounds that fall into two l)road classes:
I. BOD producing compounds that mainly generate water pollution, and
2. Odorous compounds that mainly generate air pollution.
Characteristically, the first class of compounds are volatile (boiling points between 56 and
1500 C (133 and 302° F)), chemical oxygen demanding, and partly toxic (turpentine).
Typical of the second class of compounds is that they are volatile (boiling points —59 to
117° C (-75 to 243° F)), reduced sulfur containing, odorous, chemical oxygen demanding,
and toxic.
Although this section of the manual focuses chiefly on the second class, the odorous
compounds, it is advantageous and sometimes necessary to briefly touch upon the first class,
the BOD compounds. They may be treated with the same methods, keeping in mind that
they are less volatile than the odorous substances.
The main components of typical kraft mill contaminated condensates are enumerated in
Table 5-1.
Components I to 3, especially CH 3 OH, arc mainly responsible for the BOD load of the
condensates; components 4 to 8 arc mainly responsible for the toxicity of the condensates;
and components 5 to 8 arc mainly responsible for the odor of the condensates. Because of
the entrained black liquor, the condensates are, for the most part, on the alkaline side. All
these components will exist in kraft mill condensates in varying amounts depending on place
of origin (digester or evaporator), pulp raw material (wood species), operating practices
5-1
-------
TABLE 5-1
MAIN COMPONENTS OF TYPICAL KRAFT MILL CONDENSATES (1. 2. 3)
Odor
No. Component Boiling Point BOD Sulfur Threshold
°C (°F’) kg/kg PP 1
Cl-1 3 0H 64.7 (148.5) 1.00 0 100,000
2 CH 3 CH 2 O I I 78.5 (173.3) 1.23 0 10,000
3 CH 2 COCH 3 56.5 (133.7) 0.67 0 100,000
4 Turpentine I 54 (309) 0
(Phi en c)
5 H 2 5 -59.6 (-75) 0.60 94 0.4-5
6 Cl-I 3 SH 7.6 (45.7) 0.07 67 0.4-3
7 CH 3 SCH 3 37.5 (99.5) 0.31 52 1-10
8 CH 3 SSCH 3 117 (243) 0.61 68 2-20
(such as sulfidity and cooking time), type of equipment (continuous or discontinuous), and
condition of equipment (such as capacity and age).
Because of the difficulty in separating air and water pollution aspects of condensate
treatment, both arc covered together.
5.2 I)igcster Condensates
Digester condensates will var ’ in flow and composition. Thc quantities of condensates arc
especially different for batch digesters and for continuous digcstcrs.
5.2.1 Batch Digester Condensates
The amount of turpentine decanter condensates from batch digesters is fairly similar in
different mills. The condensates may vary depending on digester pressure relief method. A
typical composition is given in Table 5.2 and flow arid load range in Table 5 3.
The amount of batch digester blow condensates, or the overflow or hlcc(l from the blow
heat accumulator (flow 3 in Figure 2 -I) is heavily dependent on the amount ol additional
fresh water (flow 2 in Figure 2-1) that is injected into the condenser (blow steam, direct
contact) to boost vapor condensation (section 2.1.1 .1). Additional water means more
condensate to treat. The importance of a properly clinicnsioncd and efficiently working
blow hcat. recovery system is again emphasized.
5-2
-------
TABLE 5-2
TYPICAL KRAFT MILL CONDENSATE COMPOSiTIONS, MEAN VALUES FOR
10 MILLS (2)
Condcnsa tc Turpen tine Means. Digester Evapora [ or Evaporator
Corn pound Decati ter I3low Effects I-b [ well
mg/I mg/I mg/I mg/I
H 2 S 90 60 40 100
CH 3 SH 250 80 10 40
CH 3 SCFI 3 400 70 5 7
CH 3 SSCI-1 3 130 50 5 IS
TotalS 550 180 51 135
CH 3 OH 6,500 4,300 10,000 1,000
CH 3 CH 2 OH 1,600 500 60 40
CH 3 COCH 3 160 40 6 10
Total BOD 860 490 I ,070 I ,060
The volume of blow condensates will diminish if the 1) 10w heat recovery system works
without heat exchanging (i.e., if the direct contact condenser is fed with cool or warm water
from the outside and the hot contaminated water in the blow heat accumulator is used
directly for pulp washing. Such a system, however, increases the odors from the pulp in the
washing area. This I)iilP will require bleaching. Therefore, such a system ‘ill probably not
meet future air pollution requirements.
5.2.2 Continuous Digester Condensates
The continuous digester relief and flash condensates are rather siniilar in composition to the
corresponding conclensalcs from batch digestcrs.
The continuous digester flash condensate will be found in different parts of the kraft mill,
depending on where the flash steam is used. Usually the black liquor from the continuous
digester is expanded or flashed in two stages, and the steam from the primary flash tank is
used in the prcsteaming vessel (Figure 2-9). Most of the noncondensable and low boiling
compounds end tip in the turpentine recovery system and are vented from that system. The
secondary flash steam may be put into the prestcaming vessel, into a condenser for hot
water generation, or into an evaporation plant as primary steam (section 3.2.1.4).
Consequently, the flash condensates will end up in those places, perhaps mixed with other
condensates. Large variations in the amounts of flash and turpentine condensates actually
emanating from the digester area will consequently occur.
5-3
-------
TABLE 5-3
TYPICAL KRAFT MILL CONDENSATE CHARACTERISTICS FOR 17 MILLS (4)
Condensates
Terpen Line Digester Evapora [ or Evaporator
Characteristic Decanter Blow Effects I lotwdll
Flow, m 3 /t (gal/ton)
Max. 0.3(72) 4(960) 6(1,440) 16(3,840)
Mean 0.15(36) 2(480)* 6(1,440) 7(l,680)
M m. 0.08(19) 0.9(216) 6(1,440) 1.5(360)
Sulfur, kg/t (lb/ton)
Max. 0.08 (0.16) 0.72 (1.44) 2.25 (4.5)* - -
Mean 0.05 (0.]0) 0.36 (0.72) 0.30 (0.6) 0.60 (1.2)
M m. 0.01 (0.20) 0.10 (0.20) 0.10 (O.2)** .. -.
BOD 7 , kg/t (lb/ton)
Max. 5 (l0)° -- -- t5 (30 - . - .
Mean ] (2) 2 (4) 6.5 (13) 3.5 (7)
M m. I (l 5 (10) - - - -
*Values exceeding 1 m 3 /t (240 gal/ton) pulp indicate fresh water addition to blow heat recovery system
or condensing system.
**lncludes evaporator hotwells.
#* 4 ’lncludes digester blow.
5.3 Evaporator Condensates
The evaporation plant condensates are usually divided into primary condensates emanating
from the first stage, secondary condensates emanating from the other stages, and hotwe ll
condensates. The hotwell condensates include condensates from primary and secondary
condensers and from the vacuum pulp (refer to Figure 3-2). If the primary steam is pure
back pressure steam, the primary condensates are clean (with no leaks) and may be returned
to the power station feedwater system.
If the primary steam is wholly or partly flash steam the primary condensate will he
contaminated and need treatment. The hotwell condensates will he usually more
contaminated than the secondary condensates. Fairly typical values for composition and
amounts of the different condensates are given in Tables 5-2 and 5-3. The composition of
the condensates varies greatly depending on wood species cooked, sulfidity, evaporation
tenipcratu rcs, and condensers.
5.4
-------
5.3.1 Evaporator CondcnsaLc Quantity Reduction
The quantity of combincd evaporator condensates should be about 7.5 m 3 per metric ton of
pulp (1800 gal/ton) without (lirect contact evaporation and with a dry solids yield of 1.5 kg
per kg of pulp (1.5 tons/ton). Thc total evaporator condensates flow, howcver, is usually
larger (see Table 5-3), because fresh water is added (see Figure 5-1). Surface condensers will
reduce the amount of condensate ( 5cc Figure 3-4). To eliminate the addition completely, a
heat exchanger can he installed and the hotwcll condensate circulated through it for cooling
an(l reuse in condenser and vacuum pull) (Figure 5-2). By feeding white liquor into the loop
there will he a simultaneous gas scrubbing and condensate return to the white liquor system
(section 3.2.1 .2).
SECONDARY
CONDENSATE
TAIL
CONDENSATE
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY (3
FIGURE 5-1
EVAPORATION PLANT SURFACE CONDENSER WITH WATER JET CONDENSER
AND WATER RING PUMP
WATER JET
CONDENSER
BAROMETI C
SURFACE
CONDENSER
LAST
STAGE
VAPOR 1
COND. 2
9 WARM WATER
8 FRESH WATER
VENT
WATER RING
VACUUM PUMP
( 5+j
5-5
-------
SECONDARY
CONDENSATE
WATER JET
CONDENSER
9 WARM WATER
8 FRESH WATER
VENT
WATER RING
VACUUM PUMP
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY (3
FIGURE 5-2
EVAPORATION PLANT SURFACE CONDENSER WITH RECIRCULATED
WATER JET CONDENSER AND WATER RING PUMP
5.3.2 Evaporator Condensate Segregation
A natural and very effective way of facilitating contaminated condensate treatment is to
separate the condensates at their point of origin and group thcni according to their sulfur
and BOD content. This procedure applies especially well to the evaporation plant, where
most of the sulfur tends to concentrate in the tail end, and most of the BOD after the weak
liquor feed.
The use of segregation is illustrated on an evaporation plant with 5 stages such as that
presented in Figure 3-2. The liquor flow is 3.4-5-2-I and stages 4 and 5, following the liquor
feed stage, are equipped with two-stage venting through liquor prchcatcrs (section 3.2. I .1).
About 15 1 rccnt of the vapor to the stage condenses in the preheater. The evaporation
plant condenser is divided into two surface condenser stages. About 13 percent of the vapor
SURFACE
CONDENSER
LAST
STAGE
VAPOR 1
COND. 2
CIRCULATION
PUMP
TAIL
CONDENSATE
5-6
-------
is condensed in the second stage and 2 percent in a water-ring vacuum pump that has
indirect cooling. Undcr these circumstances the distribution of flows, sulfur, CH 3 0 1- I, and
BOD will be approximately as presented in Table 5-4.
TABLE 5-4
CALCULATED EVAPORATOR CONDENSATE FLOW, SULFUR, METHANOL,
AND BOD DISTRIBUTION FOR LIQUOR SEQUENCE 3-4-5-2-1 (4,5,6, 7)
Point of Release
(No. refers to flows in
Fig. 3-2) Temp. Flow(a) Sulfur(b) cu 3 Oi i(c) BO1)( )
°C (°1?) % %
2—Effect,condensate l]5 (239) 21.0 1 1
2—Effect, vent
3—Effect, condensate 100 (212) 18.0 1 2 2
3—Effect, vent
4—Effect, condensate 85 (185) 15.3 3 24 22
4—Effect, preheater 80 (176) 2.7 3 33 32
4—Effect, vent 8(e) 8(e)
5—Effect, condensate 70 (158) 17.0 2 6 6
5—Effect, preheater 65 (149) 3.0 2 12 11
5—Effect, vent 4(e) 4(e)
6—Primary condenser 55 (131) 19.6 12 3 4
7—Secondary condenser 40 (104) 2.9 49 ]2 14
8—Vacuum pump 30 (86) 0.5 27 7 8
Total 100 100 100
(a) 100% flow = 6.7 5 m 3 /t of pulp (1440.1800 gal/ton of pulp).
(b) 100% sulfur = 1-2 kg/t of pulp (2.4 lb/ton of pulp).
(c)]oo% CH 3 OH = 6-9 kg/t of pulp (12-18 lb/ton of pulp)
(d) 100 % BOD = 7.10 kg/t of PUlP (14-20 lb/ton of pulp).
(C)Alrcady included in secondary condenser & vacuum pump.
5.7
-------
By combining various effect condensates, for example, from effects 2 to 5. with those from
the primary condenser, 90.9 percent of the condensates with 19 percent of the total
reduced sulfur (TRS), and 35 percent of the total BOl ) at a mixing temperature of 86° C
(187° F) are obtained without flashing. This flow may be reused within the process. The
remaining 9.1 percent of the condensates will contain 81 percent of the sulfur and 65
percent of the BOD. This small flow can be rather easily treated, for instance by steam
stripping. Then, by combining flows from 2, 3, and 5 with those of the primary condenser,
75.6 percent of the condensates are obtained that contain only 16 percent of the total
sulfur and 13 percent of the total BOD. Other liquor sequences, vaporization distributions,
and preheater locations will change the distribution of contaminants.
Condensate segregation requires special piping arrangements and usually causes a small
increase in the primary steam consumption in the evaporation plant because of incomplete
secondary condensate flashing.
Segregation of the hotwell condensates is needed solely for odor abatement. Hotwell
condensates include those from the primary condenser, the secondary condenser, and the
vacuum pump. These condensates alone contain 88 percent of the TRS in 23 percent of the
condensate volume.
5.3.3 Weak Black Liquor Oxidation
An effective way to eliminate odors is to oxidize the weak black liquor.
By oxidizing the weak black liquor, the sulfide is converted to thiosulfate and the CH 3 SI-I to
CH 3 SSCH 3 . Therefore, 1125 and CH 3 SH are not liberated in the evaporation process.
Consequently the condensates will require little, if any, treatment for odor abatement. The
oxidation can be quite effective, and different systems using air have been developed. One
example is the British Columbia Research Council (BCRC) system that works as a weak
black liquor and odorous gas oxidation system. This particular system uses bleach plant
effluent containing rest chlorine as one gas scrubbing and oxidation agent, and so obviously
it a bleach plant must be present.
There are, however, some drawbacks to weak liquor oxidation. During evaporation and
storage the elemental sulfur generated tends to reconvert to H 2 S, and the CFI 3 SSCH 3 to
Cl-I 3 511. Such reconversions will nullify to some extent the previous oxidation effort (8).
Another difficulty, especially with resinous softwoods, is the foaming of the weak liquor in
the oxidation process. Its foaming tendency is a function of its concentration. At high
concentrations the foaming decreases. Black liquors that can be oxidized as thick liquors
may be impossible to oxidize as weak liquor, even up to 23 percent dry solids (8).
5.3
-------
Furthermore, black liquor oxidation, weak or thick, will decrease the heat value of the dry
solids by an average of 523 Mi per metric ton of pulp (0.45 X I 0 BTU short ton of pulp).
Also, the oxidizing air strips off odorous compounds from the liquor. Current trends are
toward oxidation of thick black liquor just before its evaporation by direct contact with Lhe
recovcry boiler flue gases. For more information on weak black liquor oxidation, see section
9.1.
5.4 Condensate Chlorination
It is possible to deodorize condensates from digesters and evaporators by mixing elemental
chlorine (Cl 2 ) in the condensates. Since the chlorine demand of reduced sulfur compounds
is high (Table 5.5) this is an expensive method unless there is an inexpensive source of
TABLE 5.5
CHLORINE DEMAND OF REDUCED SULFUR
COMPOUNDS FOR OXIDATION TO SULFUR (9)
Compound pH Cl 2 Demand Redox Potential
kg Cl 2 1kg Compound (+) Volts
H 2 S 4 9.2 0.25
H 2 S 8 5.6 0.26
CH 3 SH 4 6.7 0.45
CH 3 SH 8 5.1 0.61
CH 3 SCH 3 4 2.2 0.44
CH 3 SCH 3 8 2.4 0.62
CH 3 SSCH 3 4 1.9 0.58
CH 3 SSCH 3 8 2.2 0.70
chlorine available. For instance, in some mills bleach plant effluent containing rest chlorine
is mixed with odorous condensates, and the combined effluent has neither H 2 S nor chlorine
odor (10).
5.5 Condensate Stripping
Stripping the contaminated condensates has proved an efficient and economical way of
removing odor and BOD. Condensate stripping is becoming the dominant treatment
5.9
-------
method. The two principal ways are air stripping and steam stripping. To facilitate stripping,
certain preconditioning techniques should be utilized.
5.5.1 Condensate preconditioning
After the volume of condensates to be stripped has been reduced to the minimum feasible
through using reduced fresh water input and condensate segregation, there are other
conditions to he met.
A higher temperature of the released contaminated condensate will aid treatment by
stripping. Air stripping will he more effective since the volatilcs will have a higher vapor
pressure. Similarly, steam stripping will require less steam for heating the condensates to the
boiling point.
The condensates should have as low a pH as possible. A low ph means a high volatility for
the ionized sulfur compounds, H 2 S and CH 3 SI-l (Figures 2-4 and 2-5); whereas, high pH
(greater than 9) results in low volatility as well as foaming. High pH is the result of black
liquor entrainment and, thus, may indicate insufficient drop and mist separation in the
vapor flows or equipment overloading. The points to watch for high pH values are the blow
tank and the evaporator drop separators.
5.5.2 Air Stripping
A very simple condensate treatment is to strip the condensates in a multistage column (that
is, a column with multiple trays) with a large countercurrent flow of air or flue air (Figure
5-3)(11).
Use of an atmospheric vent (sec flow 5 in Figure 5-3) for air or flue gas stripping simply
translates a water pollution problem into an air pollution problem. Although part of the
TRS will be oxidized while feed and air or flue gas mix in the column, most of the TRS will
pass out through the vent. With small feeds and air flows the vent can be connected to a
boiler furnace or to an incinerator, but for air flows equal to thousands of cubic meters per
hour this is hardly feasible.
Experience with working units (7) indicates that column features and performance are
approximately as follows:
1 - ‘l’ray type, bubble cap:
2. Liquid feed rate, Fre(.(l iii 3 /h (Ficed gpm)
3. Tray number, 10-20;
5-10
-------
+ ®VENT
FOUL CONDENSATES
STRIPPING
COLUMN
2 REST ACID
4 AIR
1 CLEAN CONDENSATES
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY (3
FIGURE 5-3
CONTAMINATED CONDENSATES AIR STRIPPING
PLANT FLOW SHEET
4. Tray distance, 500 mm (20 in)
5. Column diamcter, 130 V’Fi . mm (2.44 V’Ffecd in)
6. Air flow, 18 X Ff d;
7. Feed temperature, 65.70° C (149-] 58° F);
8. Feed pH, beloW 9;
9. TRS of feed, 130-320 g/m 3
10. BOD (CH 3 OH) of feed, 390-1030 gJm 3
11. TRS removal, 80 percent; and
12. BOD removal, 0-10 percent.
Tripling the air flow will increase TRS removal to 84-94 percent and BOD removal to 10-15
percent.
5-11
-------
The condensate fced must he as hot as possible and the pH must stay below 9. If pH rises
above 9, reduced stripping cfficicncv and serious foaming problems will follow. The
temperature of the ambient air will affect the feed temperature, but this effect is small,
abouL 6°C (110 F)
if the p l- l of the condensates cannot l)c brought below 9, rest acid must be added to the
condensate feed (see flow 2 in Figure 5-3). Rest acid is obtained, for example, from the
manufacture of bleach plant C 10 2 , and added at the feed pump suction. A control loop
from after the pump can be set to keep Lhe feed pH at a suitable level of about 7.8. Air or
flue gas sLripping will not remove more than 95 percent of TRS nor more than 15 percent of
BOD at reasonable gas flow rates.
5.5.3 Steam Stripping
Steam stripping was first used about 25 years ago at Skoghall, Sweden (9). This steam
stripping plant is still working and has been modernized. Steam stripping research was
carried out in piloL plants (7, 12) and, subsequently, put into full-scale operation in a num-
ber of kraft in ills.
5.5.3.1 Separate Steam Stripping Plants
Steam stripping seems to lie the most feasible method of purifying contaminated kraft mill
condensates. Steam stripping can be divided into two categories: TRS removal and BOD
removal. Characteristic features of stripping TRS are strong pH dependency and low steam
consumption, which is about 2 percent of the condensate feed for a 90 percent TRS
removal. Characteristic features of stripping BOD are much lower pH dependency and high
steam consumption, which is about 20 percent of the feed for a 90 percent BOD removal.
A typical steam stripping plant is shown in Figure 54. Condensates (flow 3) will enter
a storage tank with level alarms. Resi acid addition (flow 2) is controlled to keep pH at 7.8
after leaving the tank to enter the preheater/primary condenser. Condensates then receive
heat from outgoing stripped condensates in a heat exchanger. Then the)’ pass a direct
contact steam injector with temperature control to insure that column feed temperature is
sufficient and constant. Condensates next pass down the column countercurrently to fresh
steam injected at the bottom of the column. Condensate feed is kept constant by flow
control subject to alarm from the storage tank.level monitor. Stripping steam flow also is
kept constant, but the flow control set point will follow feed flow in a fixed ratio that can
be adjusted. This ratio will essentially be equal to the steam/condensate ratio, taking into
account steam consumed for heating the condensates to the boiling point.
The vapors rise through a fortifier section of the column countercurrently to condensate
feedback and enter the primary condenser/preheater, where most of the vapor coiidenses.
5-12
-------
8 VENT
7 FRESH WATER
6 WARM WATER
5 TURPENTINE PHASE
4 STEAM
3
2 REST ACID
FIGURE 5-4
CONTAMINATED CONDENSATES STEAM STRIPPING
PLANT FLOW SHEET
1 CLEAN CONDENSATES
5-13
-------
The noncondensables and the rest of the waLer vapor pass to the secondary condenser,
which serves as a final condenser and gas cooler. The cooling is controlled In’ [ he outgoing
gas temperature, which will sLav constant. Adjusting the set point of the controller means
adjusting the temperature and the dew point of the noncondcnsable gases from the vent
(flow 8 in Figure 5-4). By adjusting the dew point, these gases can be kept at a humidity
level (40%) that will greatly diminish the explosion risk.
The column condensates flow to a separator where eventually turpentine and other organic
compounds may form an oily layer to be drawn off with a level control and piped to
incineration (flow 5 in Figure 54). The underfiow is pumped back to the column fortifier
section top under level control. To give a very general idea of how to dimension and what to
expect from a steam stripping column, the following data are given:
Tray type, l)ul)l)le cap;
2. Steam flow rate, Fstcani in metric tons per hour (Fsteain short ton/hour)
3. Tray number, 10.20;
4. Tray distance, 500 mm (20 in);
5. Column diameter, 780 V’Fstcarn mm (29.2 V’Fstcam in)
6. Feed pH, below 9;
7. TRS of feed, 210 g/m 3
8. BOD of feed, 300 g/m 3
The removal percentage as a function of the steam/condensate ratio is approximately as
shown in Figure 5-5. A lower p ’- I of the feed would improve these removal rates for ‘ ‘2 S,
cspcciaUy at low steam/condensate ratios. After the contaminants have been stripped out of
the condensate, and the main part of the water vapor condensed, the remaining gases are
usually incinerated. Such gases may be incinerated in either the lime kiln, a separate
incinerator, or possibly the recovery boiler.
5.5.3.2 Evaporation Steam Stripping Plants
Stripping of the condensates is usually performed in a separate stripping column with fresh
steam. One way to reduce the high treatment costs involved is to combine this stripping
with the evaporation plant, using “secondary steam” from the black liquor evaporation step.
Such an arrangement is shown in Figure 5-6.
5-14
-------
100
80
a 2
- 60
>-
C-)
2
LU
0
40
LU
-j
0
20
L i i
0
l0
STEAM/CONDENSATE, AS %
FIGURE 5-5
STRIPPING EFFICIENCY FOR DIFFERENT STEAM-CONDENSATE
RATIOS WITH 10 THEORETICAL PLATES
The stripping column is placed on top of the sccond effect, similarly to the alcohol stripping
in a sulfite spent liquor evaporation plant. All dirty condensates arc piped to the stripping
column, where the stripping efficiency is 95 percent. The CH 3 OH stripped off is withdrawn
through [ he liquor preheater of the third effect together with a small amount of steam and
is condensed to form an 8.10 percent CH 3 OH ‘ater solution. The CH 3 OH is recovered in a
small stripping column and destroyed by burning. Alternatively, it can be recovered and
sold. The total BOD-removal of the condensates by this stripping arrangement is estimated
at 90 percent. If the stripping column is placed on top of the second evaporator effect, the
heating surface of the evaporation plant has to be increased by approximately 7 to 8
percent. This increase is necessary to compensate for the pressure loss in the column and the
lower condensation temperature in the third effect. The steam consumption of the
evaporation plant increases by 5 to 7 percent.
0 2 4 6 8
515
-------
STRIPPING
COLUMN
STEAM
C’
FOUL CONDENSATE
FROM DIGESTER
CONDENSATE TO REUSE
FIGURE 5-6
CONDENSATE STRIPPING IN AN EVAPORATION PLANT (13)
-------
By using the stripping arrangement shown in Figure 5-6, the evaporation plant will produce
thick liquor and clean, practically distilled water for reuse, plus concentrated volatile
organics as a byproduct.
5.5.4 Condensate Finishing
Condensate finishing may be feasible when very high removal efficiencies of greater than 99
percent are required. Instead of stripping with an excessively large steam/condensate ratio,
enough stripping steam is first used to decrease TItS by 90 or 95 percent, and then the
stripped condensate is treated with ozone, chlorine, or activated carbon (14). Condensate
stripped of TRS may he finished in a biological treatment plant where most of the
remaining traces of TR.S and 1301) are removed. Biological condensate finishing can be
accomplished if the mill already has a biological Lrcattnent plant for its other wastewaters.
5.5.5 Condensate Reuse
Condensate reuse includes reuse of both treated and untreated condensates. Vhen using
treated condensates that have been stripped of their TRS and BOD, there are no reuse
problems since the treated condensate is clean, hot, distilled water.
When reusing condensates that have been stripped of their TRS components only, the BOD
components will come out from the process at some other point. If the components are
recirculated, they will build up within the process unit until such a level is reached that
discharge takes place. The best way to treat such condensates is to direct the discharge of
BOD components into the recovery boiler through the black liquor.
When using untreated contaminated condensates, the TRS must be removed since it will
either go down the drain to the receiving waters or it will start circulating and budding up
until it is discharged to the atmosphere and/or to the water. A possible solution is to put the
TFtS back into the white liquor, where it will provide part of the sulfide and decrease the
sulfur make-up demand.
5.5.5.1 Pulp Washing
Untreated contaminated condensates have been used for pulp washing, especially in mills
with direct blow heat recovery (see section 5.2.1). But such an arrangement causes air
pollution from washer hoods, and water pollution from the pulp screening, if an
open-screening system is used. A closed-screening system will reintroduce the water
pollutants, raise their level, and increase air poHution.
Before contaminated condensates are used for pulp washing, they should be stripped
completely for open-screening systems, and at least of their IRS components for
5-17
-------
closed-screening systems. A stripped condensate is a very good washing liquid, namely hot
distilled water. The total condensate amount, without adding fresh water, will be about
7.5-8.5 m 3 /per metric ton of pulp (1800-2000 gal/ton). The amount will depend upon
whether there is continuous or batch digestion. Of this amount, about 8 m 3 per metric ton
of pulp (1900 gal/ton) can be used for washing pulp, thus meeting thc whole wash water
demand.
5.5.5.2 Lime Kiln Flue Gas Scrubbing
Using contaminated condensates as scrubbing liquids in lime kiln fluc gas scnibbers amounts
to stripping them with hot gases while they reciprocally wash particulates from the gascs.
The result is increased air pollution caused by TRS. To usc stripped condensates in lime kiln
or flue gas scnibbers offers no advantagc, since ordinary fresh water will do the same job
while additionally picking up hcat from the hot gases. Stripped condensates are hot distilled
water and should be used in proper applications, such as washing pulp in the washing
department or in the bleach plant. Fresh water demand for a lime kiln scrubber is about
4 m 3 per metric ton of pulp (960 gal/ton).
5.5.5.3 White Liquor Liquid Makeup
In the digester, the liquid/dry wood ratio can be trimmed by black liquor reeireulation. The
necessary liquid makeup to the white liquor can be supplied at three points, namely in the
smelt dissolving tank, in the mud washers, and in the white liquor itself.
The smelt dissolving tank requires liquid at about a rate of 1.5 m 3 per metric ton of pulp
(360 gal/ton). The liquid should preferably be cool water, since the heat input to the smelt
dissolver is already so great that addition of hot water would cause vaporization of both
water and volatile and odorous substances. Condensates, therefore, are not suitable for smelt
dissolver liquor makeup unless they are first stripped of TRS and cooled.
The mud washers need hot water and usually 50 percent of the white liquor liquid demand,
approximately 1.5 m 3 per metric ton of pulp (360 gal/ton), enters the process this way.
Contaminated condensates will cause odor problems at the mud washer, but TRS stripped
condensates may be used although BOD components then will be fed back into the process.
If the mud washer system is highly effective, that is, if it will work with less wash water than
the 1.5 m 3 per metric ton of pulp (360 gal/ton), the remaining makeup liquid can be
supplied directly to the white liquor. Treated or untreated condensates can be used since
odorous compounds will, for the most part, be absorbed by the alkali of the white liquor.
The BOD level will rise, of course.
5-18
-------
A plan for condensate segregation, stripping, and reuse is suggested in Figure 5-7. This i 1 n
will require about 1.5 m 3 per metric ton of pulp (360 gal/ton) of fresh water makeup and
about 1 m 3 per metric ton of pulp (240 gal/ton) of hot water makeup for the pulping cycle.
Seven percent of the combined digester and evaporator condensates, containing 9 I percent
of the combined TEtS and 73 percent of thc combined 130D, arc stripped and used for
washing pulp and lime mud. Seventeen percent of the BOD is returned to the white liquor.
Ten percent of the BOD and 9 percent of the TRS arc returned to pulp washing.
Approximately 50 percent of these contaminants will remain with the pulp and get washed
out. during screening; the rest will return to the black liquor.
(2iDS) WOOD
4m’
1 3 rn 3
1 0m 5
0 2m 1
12
PULP )09i OS)
55m’
I I I I I I I I I
15 15 13 02 13 02 13 02 —
115 tOO 85 80 10 65 55 40 30 ‘C
01 02 22 32 06 11 04 14 08 kgBOD
— 01 05 O3k 9 TRS
_ PP-
0 6m’
1 6m’
0 2m
SIMPLIFIED LIQUID
FIGURE 5-7
FLOW SHEET FOR KRAFT PROCESS WITH CONDENSATE
SEGREGATION, STRIPPING, & REUSE
I
1 m’
8m’
rule L q
Black L’q
15t DS
8 5m’
een Liq
T tOrn’
Hot Waiei
1 4in’
56rn’
Flue Gas
Fresh Water
lrn
3 4a 4b 5a 5b 6 1 81
5-19
-------
5.6 References
1. Nylander, C., Report on Forest Industry Waste Waters. Svcnsk Papperstidning
67(1 5):565-572, August 1964 (Stockholm).
2. Lcornados, C.. Kendall, D., and Barnard N., Odor Threshold Delernunations of 53
Odorant Chenucalc. Journal of Air Pollution Control Association, 19:91 -95, February
1969.
3. Withy, F. V., Variation in Recognition Odor Threshold of a Panel. Journal of Air
Pollution Control Association, 19:96-] 00, February 1969.
4. EKOl’ O Oy, Helsinki, Finland, files.
5. Arnc, H. C., and Bcrgkvist, S., Methanol Distribution in an Evaporation Plant. Svcnsk
Pappcrstidning, 77(10) :380-382, 1973 (Stockholm).
6. Jonsson, S. E., Black Liquor Evaporation. Svensk Pappcrstidning, 74(7):191.196, April
15, 1971 (Stockholm).
7. Backstrom, B., Helistrom, H., and Kommonen, F., Purification of Malodorous Sulfur
Containing Condensates from Turpentine Separation, Digester I3low and Spent Liquor
Evaporation at the Oy Kaukas Ab, Kraft Mill. Paperi ja Puu, 52(3):113-120, 1970
(Helsinki).
8. Murray, F. E., The Oxidation of Kraft Black Liquor. Pulp & Paper Magazine of Canada,
64:82-86, January 5, 1965.
9. Ruus, L., Report on Forest Industry Waste Waters. Svensk Papperstidning, 67(19):751-
755, October 15, 1964 (Stockholm).
10. Lindberg, S, How Uddeihoim Destroys Air and Water Pollutants at the Skoghall Works.
Pulp and Paper Magazine of Canada, 69(7):T178.T183, April 5, 1968.
11. Morgan, 1. P., and Murray, F. E.,A Comparison of Air and Steam Stripping as Methods
to Reduce Kraft Pulp Mill Odor and Toxicity from Contaminated Condensate. Pulp &
Paper Magazine of Canada, 73(5):62.66, May 1972.
12. Matteson, M. I., Johanson, L. N., and McCarthy, J. L., SEKOR II; Steam Stripping of
Volatile Organic Substances from Kraft Pulp Mill Effluent Streams. Tappi, 50:86-91,
February 1967.
5-20
-------
13. Study of Pulp and Paper 1ndustry s Effluent Treatment. EKONO Oy. Prepared for
FAO Advisory Committee on Pulp and Paper. Session 13, Rome, May 15-16, 1972.
14. hansen, S. P., and Burgess, F. t., Carbon Treatment of Kraft Condensate Wastes. Tappi,
51:241-245, June 1968.
5-21
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CHAPTER 6
BROWN STOCK WASHER GASES
After mixing with the original liquor and added black liquor, the pulp passes from the blow
tank to the washing plant. The washing process is a minor source of air pollution compared
to combustion, evaporation, and digestion. Generally, the washing produces large quantities
of ventilation air slightly contaminated with organic sulfur compounds through contact with
the black liquor. The amount of air and of sulfur compounds will mainly depend on the
type of washing process and equipment and to a minor degree on wood, sulfidity, pH, tem-
perature, and other factors that are generally determined by the production process. The
two main washing processes are displacement washing and diffusion washing.
6.1 Displacement Washing
Displacement washing usually takes place on rotary drum filters using air for maintaining
the pressure difference over the washed pulp sheet. Thus, the hot black liquor is exposed to
large quantities of air. This exposure will have two effects. One effect, for example, is that a
portion of the reduced sulfur in the black liquor will be oxidized and will decrease the sub-
sequent odor generating capacity of the black liquor during evaporation. The other effect is
that part of the reduced sulfur will become volatile and will contaminate the air.
6.1.1 Vacuum Washers
The most common type of kraft pulp washer is the vacuum washer (Figure 6-1). The two
main points of odor release to the atmosphere are the hood vent (flow 5 in Figure 6-1), with
large flows of slightly contaminated air and the foam tank vent (flow 6 in Figure 6-1), with
smaller flows of more polluted air. Tables 1-2, 1-3, and 14 summarize typical emissions
from these sources.
Because of their large volume and low concentration of odorous components, the only prac-
tical way of treating washer gases, especially those from the hood vent, is to incinerate them
in an existing boiler. For example, the gases from the washer vents are used as part of the
combustion air in an auxiliary furnace or a black liquor recovery boiler at several U.S. and
Swedish mills. Proper safety precautions must be designed into the system. These precautions
include condensate traps, rupture disks, flame arresters and flame control, and emergency
vents. The smaller flow rate of the foam tank vent gases allows incineration in the lime kiln.
The sulfur content of foam tank gases varies with their source, that is, whether softwood or
hardwood, see Table 6-1.
Some mills use contaminated condensates from blow heat recovery accumulators or evapo-
rators for washing. This practice will increase the odor release from the washers significantly.
The hood vent TRS emission may increase 5 to 15 times, and the foam tank vent TRS emis-
sion may increase from 20 percent to 4.5 times when changing from fresh hot water wash to
condensate wash (1). The abatement method is to strip the condensates before use.
6-1
-------
( ) VENT
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY Q
4 WASH
3 WASI-1
2 WASHED PULP
PULP
VENT
WEAK
LIQUOR
6.1.2 Pressure Washers
FIGURE 6-1
VACUUM WASHERS FLOW SHEET
Pressure washers have closed hoods with blowers circulating air for maintaining the pressure
difference across the pulp sheet (Figure 6-2). The hood vent (flow 5 in Figure 6-2) and the
foam tank vent (flow 6 in Figure 6.2) will both havc a small flow rate and a high sulfur
concentration compared to the vacuum washer hood vent. Black liquor oxidation and
release of reduced sulfur to the atmosphere will be less, too.
Odor abatement is much easier than for a vacuum washing plant and can best be
accomplished by incineration of the vent gases in the lime kiln.
6.2 Diffusion Washers
Diffusion washing usually takes place in a closed reactor, and ideally there is no air involved.
Therefore, black liquor oxidation and odor release are very small, when compared with
displacement washers.
HOOD
1
6-2
-------
TABLE 6-1
FLOWS, COMPOSITION AND SULFUR RELEASE OF
VACUUM WASHER FOAM TANK VENT GASES (2)
Wood Species
Flow
m 3 /t
(ft 3 /ton)
Component
Sulfur
kg/t (lb/ton)
Pine
64
H 2 S
0
0
(1930)
CH 3 SH
0.03
0.06
CH 3 SCH 3
0.06
0.12
CH 3 SSCH 3
0.04
0.08
Total
0.13
0.26
Birch
65
(2080)
H 2 S
CH 3 SH
CH 3 SCH 3
CH 3 SSCH 3
Total
0
0.02
0.13
0.13
0.28
0
0.04
0.26
0.26
0.56
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY (3
FIGURE 6-2
PRESSURES WASHERS FLOW SHEET
VENT
4 WASH
3 WASH
2 WASHED PULP
VENT
WEAK LIQUOR
1
6-3
-------
6.2.1 Batch Diffusers
Batch diffusers can still be found in old mills, but in many cases have been replaced by
vacuum washers (Figure 6-3). Because of the batch operation there is liquor-air contact
during blowing, washing, and emptying of the diffuser. Consequently, there is some black
liquor oxidation and odor release, symbolized by flow 5 in Figure 6-3. The sulfur release
will be difficult to determine and abate. Continuous venting to the lime kiln is probably tl e
best method of gas treatment.
VENT
3 WASH
2 WASHED PULP
WEAK LIQUOR
FIGURE 6-3
BATCH DIFFUSION WASHERS FLOW SHEET
6.2.2 Continuous Diffusers
Continuous diffusers have been integrated with continuous digesters for the past decade.
They are also now made as separate washers (Figure 6-4). The washing process is closed-off
to minimize air infiltration. Thus, the oxidation of the black liquor and the subsequent
release of reduced sulfur, symbolized by flow 5 in Figure 6-4, are kept at a minimum. The
reduced sulfur is rather easily contained and is incinerated in the lime kiln.
Scrubbing of washer gases with an alkaline solution, such as white liquor, is not generally
practiced since the washer gas TRS is predominantly nonionizable sulfur compounds.
6-4
-------
CONTINUOUS DIFFUSION WASHERS FLOW $HEET
I I
I I
VENT
WEAK LIQUOR
3 WASH
2 WASHED PULP
I PULP+ LIQUOR
6.3 References
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY CJ
FIGURE 6-4
CONTINUOUS DIFFUSION WASHERS FLOW SHEET
1. Atmospheric Emissions from the Pulp and Paper Manufacturtag Industry.
EPA-450/1 -73-002. September 1973. (Also published as NCASI Technical Bulletin
No. 69, February 1974.)
2. Kekki, R., Kraft Mill Odor Abatement
Incineration. M.S. Thesis, Wood Industry
Finland. September 18, ] 969 (Finnish).
by Condensate Stripping and Waste Gas
Department, Helsinki Technical University,
6-5
-------
CHAPTER 7
STORAGE TANK VENT CASES
All storage tanks, especially black liquor tanks [ hat hold sulfide-containing liquor, are
potential air polluters when vented to the atmosphere. All of the storage tanks of a mill
together, however, are a minor source of air pollution, even smaller than the vacuum washer
hoods.
7.1 Storage Tank Vent Gas Composition
The liquor in the tank is usually hot, and the tank vent will emit water vapor contaminated
mainly with organic reduced sulfur compounds. The composition of the noncondensable
gases is rather similar to that of the washer gases. The main factors influencing release of
odorous gases are liquor sulfide concentration, pH, and temperature. These factors are fixed
by external circumstances (i.e., the general production process), and cannot be changed
except within a very narrow range.
7.2 Storage Tank Vent Gas Treatment
The fact that storage tanks are dispersed over the mill, thus constituting several air pollution
sources, makes their treatment quite difficult. An effective program is to connect all storage
tank vents to a central duct leading to an incinerator or an oxidation tower. This system is
applied in one Scandinavian mill (1) and at least one U.S. mill (2).
One way of reducing the release of odor from black liquor storage tank vents is to use weak
black liquor oxidation. (See Chapter 9.)
7.3 References
1. Air Pollution Abatement Problems of the Forest Industry. Statens Naturvardsverk
(Sweden). Publication 1969: 3, July 1969.
2. Michigan Department of Natural Resources, Division of Air Pollution Control. Staff
Activity Report dated April 24, 1973.
7-1
-------
CHAPTER 8
TALL OIL RECOVERY GASES
When pulping softwood with an alkaline process, recovery of tall oil can be quite profitable.
Soap is skimmed from weak, intcrmediatc, and strong black liquor storage tanks and
evaporators, and from black liquor oxidaLion plants. To liberate the fatty acids from their
sodium salts, the soap is acidified with sulfuric acid (H 2 SO 4 ). The H 2 SO 4 will also displace
other weak acids present, such as l-1 2 S and CH 3 SH, creating a potential air pollution
problem. Tb two factors with the greatest influence on the odor releasc arc the soap
washing efficiency and the rccovcry mode (hatch or continuous). Tall oil recovery gases are
minor odor sources in the kraft mill.
8.1 Batch Tall Oil Recovery
A typical batch tall oil plant flow scheme is depicted in Figure 8-1. The major odor emission
point is the boiler vent (flow 5 in Figure 8.1). When H 2 SO 4 is mixed with soap, the sulfide
of the residual black liquor, which has not been washed away, reacts to form 1-12 S. Thus,
there is first a sudden surge followed by a gradual decrease in evolution of H 2 S. Another
surge of 112 S may follow when the brine is neutralized with white liquor. The major part of
the noncondensable flow is air. The 1-12 S concentration may vary between zero and 2
percent. Other TRS components and some typical concentrations arc given in Table 8-I.
Flow rates and sulfur emission rates of TRS arc given in Table 8-2.
7 ACID
6 WATER
VENT
4 WHITE LIQUOR
3 STEAM
TALLOIL
SALT SOLUTION
TO EVAPORATOR
FIGURE 8-1
BATCH TALL OIL PLANT FLOW SHEET
8
9
TO EVAP
8-1
-------
TABLE 8-1
BATCH TALL OIL RECOVERY PLANT
TRS COMPONENTS AND TYPICAL
CONCENTRATIONS (1,2)
IRS Component Typical Concentration
g/m 3 (gr/cu It)
H 2 S 8.7 (3.8)
CH 3 SH 0.3 (0.13)
CH 3 SCH 3 0.2 (0.09)
CH 3 SSCH 3 0.03 (0.01)
TABLE 8.2
TALL OIL RECOVERY NONCONDENSABLE GAS
FLOWS AND REDUCED SULFUR
EMISSIONS (1)(2)
Condition Gas Flow TRS Emission
m 3 it (Cu ft/ton) kg of Sit (lbs/ton)
Maximum 20 (641) 0.66 (1.32)
Mean 11 (352) 0.15 (0.30)
Minimum 1 (32) 0.01 (0.02)
Since most of the rl RS is H 2 S, one treatment method often used is to duct the boiler vent
(flow 5 in Figure 8-1) directly to a white liquor scrubber.
8.2 Continuous Tall Oil Recovery
A typical continuous tall oil plant flow scheme is depicted in Figure 8.2. The major odor
emission point is the reactor vent (flow 5 in Figure 8.2). The Ilow is much smaller but more
conccntratcd in i’RS than for the batch process. There arc no surges of l I2 S, but a continu-
ous flow. The best treatment method is to apply a white liquor scrubber to the ‘cnL.
8-2
-------
6 WATER
3 STEAM
®VENT
TALL OIL
SALT
SOLUTION
FIGURE 82
CONTINUOUS TALL OIL PLANT FLOW SHEET
8.3 References
I. Air Pollutton Abatement Problems of the Forest Industry. Statcns Naturvardsverk
(Sweden). Publication 1969:3, July 1969.
2. EKONO Oy, Helsinki, Finland, files.
7
4
8
9
POINTS OF POSSIBLE ODOR RELEASE ARE ENCIRCLED BY Q
8-3
-------
CHAPTER 9
BLACK LIQUOR OXIDATION
Black liquor oxidation is extensively applied to facilitate odor control and chemical
recovery in kraft pulp mills. Its immediate purpose is oxidation of Na 2 S to innocuous salts
to prevent the release of H 2 S. Black liquor oxidation can be performed on either weak or
strong black liquor by air or molecular oxygen. Overall reviews of black liquor oxidation
practices have been prepared by Collins (1), Landry (2), Hendrickson (3), and Blosser and
Cooper (4).
9.1 Weak Black Liquor Oxidation — Air
Weak black liquor can be oxidized with air to decrease reduced sulfur emissions from both
multiple-effect and direct-contact evaporation systems. Systems in extensive usc for weak
black liquor oxidation include sieve tray towers (5), porous carbon black diffusers (6),
packed absorption towers (7), vertical-slat falling-film packed towers (8), and agitated air
spargers (9). Rotating fluid contactors (10) and pressurized vessels (11) are uscd to a lesser
extent for weak liquor oxidation.
A previous survey by Blosser and Cooper (4) indicates that it is possible to obtain
consistently high efficiency of Na 2 S oxidation with the porous diffuser (Collins) and the
sieve tray tower (Trobeck), provided that sufficient gas-liquid interfacial contact areas and
air flow rates are used and that liquor and Na 2 S loadings are kept sufficiently low. Less
effective performance is observed for packed towers because of inadequate liquor retention
times and for agitated air spargers because of frequent mechanical breakdowns.
Most weak black liquor oxidation systems, employing air in the United States, are located in
the Pacific Northwest, upper Midwest, and Northeast, .where highly resinous pine wood
species are not pulped. Blosser and Cooper (4) report the following problems with weak
black liquor oxidation systems using air:
1. Excessive foaming when pulping highly resinous pine wood species,
2. Incomplete Na 2 S oxidation efficiency caused by improper or under-design of
systems,
3. System overload caused by increased pulping capacity without expansion of
existing facilities, resulting in inadequate liquor retention time, and
4. Inability to achieve effective oxygen mass transfer from air into the black liquor.
9-1
-------
9.1.1 Porous Plate Diffusers
Collins (12) reports on the development of a two-stage system for weak black liquor
oxidation with air in consecutive aeration and deaeration steps (see Figure 9-1). The
aeration stage employs passage of black liquor across a series of horizontal porous plates
arranged vertically in a parallel flow arrangement. Air is blown consecutively in series
through the porous plates in a erossflow configuration to create gas-liquid interfacial contact
by the generation of foam. The liquor depth on the plates is normally 10 to 20 cm (4 to 8
in).
The air-foam-liquor mixture flows off the plates for deaeration in a retention tank to
provide for gas-liquid phase separation. Foam is dissipated by mechanical foam breakers
atop the deaeration tank, and liquid droplets entrained in the exit gas stream are returned to
the deaeration tank via a cyclone separator, as shown in Figure 9-1. Recent studies by Van
BLOWER
TO
ATMOSPHERE
MOTOR
FOAM BRAKER
FIGURE 9-1
COLLINS POROUS PLATE DIFFUSER WEAK BLACK LIQUOR
OXIDATION SYSTEM (14)
FOUR STAGE
OXIDATION
TOWER
STACK
BLACK
LIQUOR
PUMP
9-2
-------
Donkelaar (13) and Shah and Stcphenson (14) indicate that the Na 2 S concentration in weak
black liquor can be reduced to below 100 mg/I by a porous plate diffuser oxidation system.
Design and operating parameters calculated from published data for the two mills are
presented in Table 9-1. The systems are not normally suitable for highly resinous pine black
liquors because of excessive foaming.
TABLE 9-1
DESiGN AND OPERATING PARAMETERS FOR POROUS PLATE DIFFUSER
BLACK LIQUOR OXIDATION SYSTEMS (13, ]4)
Parameter Mill A (13) Mill B (14 )
Production, t (ton)/day 450 (500) 180 (200)
Liquor flow, m 3 /h (gpm) 205-227 (900-975) 68-108 (300-475)
Air flow, rn 3 /h (efm) 25,600 (16,500) 13,700 (8,850)
Power, kW (hp) 186 (250) 75 (100)
Plate area, m 2 (It 2 ) 650 (7,000) 31.2 (3,350)
Na 2 5 loading, kg/rn 2 /b (lb/It 2 /hr) 2.4-2.9 (05-0.6) 1.2-24 (0.2.0.5)
Liquor loading, m 3 /m 2 lh (gal/ft 2 /hr) 0.30-0.35 (7.4-8.6) 0.22-0.35 (5.3-8.5)
Air loading, rn 3 lm 2 /h (eu ft/ft 2 /hr) 39 (128) 44 (144)
Power loading, kW/t/clay (hp/ton/day) 0.34 (0.41) 0.34 (0.41)
Na 2 5 inlet,g/l 8.0-10.0 2.3-6.6
Na 2 S outlet, gil 0.01-0.05 0.01-0.10
Oxid. effie., % Na 2 S 994 98-99k
Oxygen ratio, actJtheor 7.7 14.2
9.1.2 Sieve Tray Towers
The Trobeck-Ahien, Bergstrom-Trobeek, or Lundberg weak black liquor oxidation systems
consist of countercurrent flows of air and liquor through a series of perforated sieve trays
(15). The system is arranged in consecutive aeration and deaeration stages for consecutive
air-liquor contact and air-liquor separation in series. The aeration stage consists of a
seven-stage vertical sieve tray absorption tower where black liquor can be added at the third,
fifth or seventh tray (numbered upward). The liquid is allowed to cascade downward from
one tray to the next through a series of overflows and downcomers.
Air is introduced upward through the bottom of the tower and passes upward through the
sieves, generating small gas bubbles and foam to facilitate interfacial gas-liquid contact. The
liquor is drained from the oxidation tower into a deaeration tank. Sufficient retention time
allows entrained gas bubbles to separate from the liquid. Mechanical foam breakers are used
9-3
-------
for foam control. The exhaust air stream from the oxidation tower passes through a cyclone
separator for removal of entrained foam and liquor droplets. The system is illustrated in
Figure 9.2.
Extensive experience has been obtained with sieve tray systems for wcak black liquor
oxidation in Europe, Canada, and the United States since 1942. Yemchuk (16) and
Kacafirek (]7) report Na 2 S oxidation efficiencies from 85 to 95 percent with parallel
operation of Trobeck-Ahlen weak black liquor oxidation towers. Sylwan (18) reports Na 2 S
oxidation efficiencies of up to 98 percent for a single tower Trobeck weak black liquor
oxidation unit.
The higher oxidation efficiencies reported by Sylwan are at least parLially the result of the
greater air-to-liquor flow ratio of 36 m 3 air/m 3 black liquor (4.8 cu ft air/gal black liquor)
as opposed to the lower values reported for the other systems of 15.0 m 3 air/m 3 black
liquor (2.0 cu ft air/gal black liquor).
Blosser and Cooper (4) report on extensive experience with use of multiple tray-type black
liquor oxidation towers at several kraft pulp mills in the United States. Relatively high Na 2 5
oxidation efficiencies of 96 to 99 percent are noted for all units without excessive foaming
when the pine furnish is 20 percent or less. A summary of results is presented in Table 9-2.
Two recent installations, employing Trobeck-Ah len weak black liquor oxidation systems,
have been modified to prevent excessive foaming at mills pulping substantial quantities of
pine wood species. Rippee (19) reports on a system for foam control at a western U.S. kraft
mill pulping approximately 40 percent pine wood species (primarily ponderosa and western
white). The liquid collected from the cyclone separator is drained to the tall oil recovery
system instead of the deaeration tank. A portion of the strong black liquor is recycled to the
inlet of the weak liquor oxidation tower to increase the solids concentration from 11 to 18
percent by weight. Approximately 20 percent of the liquor entering the tower is strong
black liquor, resulting in a substantial decrease in foaming problems. Na 2 S oxidation
efficiency for the system has averaged 96 percent over an extended period.
Robinson (20) reports on modification of the design of a Trobeek oxidation unit to control
foam at a southern U.S. kraft mill pulping about 70 percent pine wood species. A portion of
the air is introduced at the tangential inlet of the cyclone separator to act as a piston for
controlling the foam layer. The liquid stream passes through a series of deaeration tanks for
further foam suppression and soap recovery. The Na 2 S oxidation efficiencies have averaged
approximately 98 percent over an extended period.
9-4
-------
1
CYCLONE
SEPARATOR
—c,G-1 OXIDATION
I TOWER
— X3--I ______
___% 1 \f ___
— STORAGE AIR ‘ — — DEAERATION
TANK INLET’ E -TANK
BLACK
BLACK
LIQUOR
INLET - LIQUOR
______ _____ EXIT
-, p ____
FIGURE 9-2
TROBECK..AHLEN MULTIPLE SIEVE TRAY WEAK BLACK LIQUOR OXIDATION SYSTEM (15)
-------
TABLE 9-2
DESIGN AND OPERATING PARAMETERS FOR TROBECK-AHLEN MULTIPLE SIEVE TRAY WEAK BLACK LIQUOR
OXIDATiON UNiTS (4)
Parameter Mill A Mill B Mill C Mill I )
Liquor flow, m 3 /h (gpm) 73 (320) 79 (350) 98 (430) 613 (300)
Air flow, m 3 /h (cfm) 8,500 (5,000) 8,500 (5,000) 8,500 (5,000) 8,500 (5,000)
Plate area, m 2 (ft 2 ) 72 (780) 62 (665) 62 (665) 63 (672)
Na 2 S loading, Kg/m 2 /h (lb/ft 2 /hr) 6.3 (1.3) 8.3 (1.7) 11.2 (2.3) 12.2 (2.50)
Liquor loading, m 3 /m 2 /h (galIft 2 lltr) 1.0 (24.6) 1.3 (31.5) 1.6 (38.6) 1.1 (26.7)
Air loading, m 3 /m 2 /h (cu ft/ft 2 /hr) 118 (388) 137 (451) 137 (451) 135 (444.)
Na 2 S inlet, gIl 5.6 6.5 7.0 II .3
Na 2 S outlet, gIl 0.03 0.2 0.3 0.1
Oxid. effic., % Na S 99 97 96 99
Oxygen ratio, act./thcor. 14.3 12.7 16.8 19.0
-------
9.1.3 Packed Towers
Packed towers are used for weak black liquor oxidation with air at kraft pulp mills in the
Pacific Northwest with varying and somewhat limited effectiveness. The basic principle of
operation is to provide gas-liquid contact by providing a large interfacial surface area
through the use of packing instead of foam. Two different types of packed tower systems
for weak black liquor oxidation with air have been developed, one by the Weyerhaeuser
Company (21) (22) and the other by the British Columbia Research Council (23) (24).
The Weyerhaeuser packed tower weak black liquor oxidation system uses concurrent
downward vertical contact of air and black liquor in the tower to control foaming (25). The
tower is packed with conventional packing materials to provide for gas-liquid interfacial
contact area as shown in Figure 9-3. Design and operating parameters for several actual
systems are presented in Table 9-3 (4).
The weak black liquor oxidation system of British Columbia Research Council introduces
air and liquor concurrentl r at the top of parallel towers using successive layers of vertically
layered packing sheets. The sheets allow wetting of both sides, and the spaces provide for
lower pressure drops and resultant lower power requirements than do conventional packed
tower configurations. Murray (25) reports that vertically layered asbestos packing sheet
allows effective gas-liquid contact. West (26) observes that Na 2 S oxidation efficiency ap-
proaches 100 percent at relatively low loading rates below 140 kg Na 2 S per hour per m 2 of
packing area (29 lb/hr per ft 2 ) when two oxidation towers are located in a series arrangement.
Packed towers are relatively simple devices to construct and operate, and they have minimal
foaming problems and low horsepower requirements. They can be successfully operated
when followed by a deaeration tank to prevent air entrainment in pumps. Most installations
do not provide sufficient retention time for nearly complete oxidation of Na 2 S to occur,
have insufficient capacity for the pulp production rates involved, and are subject to plugging
with pulp, particularly from continuous digesters.
9.1.4 Agitated Air Sparging
Agitated air spargers are used for weak black liquor oxidation at a few kraft pulp mills. The
air sparger system employs a completely mixed tank containing weak black liquor with air
dispersed through a turbine aerator at the bottom of the tank. A rotary agitator, located
immediately above the aerator, shears the air into small bubbles to maximize the gas-liquid
interfacial contact area. Additional power is provided to break the air into small bubbles
instead of providing additional fan capacity and excess air. A typical unit is diagramed in
Figure 9-4.
9.7
-------
FORCED
DRAFT
FAN
FROM
BLOW— ,.
TANK
HOT
WATER
ACCUM ULATOR
a
BLACK
LIQUOR
I NLET
PACKED
TOWER
AAA
EXHAUST
GAS
DEAERATION
TAN K
BLACK
LIQUOR
EXIT
WEYERHAEUSER SYSTEM (21)
AIR HEADER
OXIDIZED BLACK LIQUOR
TO EVAPORATION
BRITISH COLUMBIA RESEARCH COUNCIL SYSTEM
FIGURE 9-3
(23)
PACKED TOWER SYSTEMS FOR WEAK BLACK LIQUOR OXIDATION
VENT
L .
L
FORCED
DRAFT
FAN
BLACK
LIQUOR
INLET
PACKED
TOWERS
Al R
VENT
- ‘ S
AIR
VENT
- ‘ S
-V
9-8
-------
TABLE 9-3
DESIGN AND OPERATING PARAMETERS FOR WEYERHAEUSER CONCURRENT FLOW PACKED TOWER WEAK
BLACK LIQUOR OXIDATION SYSTEM (4)
Parameter Mill A Mill B Mill C Mill D
Liquor flow, m 3 /h (gpm) 102 (450) 68 (300) 238 (1,050) 127 (560)
Air flow, rn 3 /h (cfm) 6,100 (3,600) 17,000 (10,000) 11,400 (6,700) 10,800 (6,400)
Packing area, m 2 (It 2 ) 12,900 (139,000) 13,950 (150,000) 22,000 (237,000) 35,000 (376,000)
Packing volume, m 3 (cu It) 1,300 (46,000) 1,050 (37,200) 2,410 (85,000) 3,370 (119,000)
Tower height, m (ft) 7.0 (23) 11.0 (36) 13.1 (43) 11.0 (36)
Na 2 S loading:
Surface area, kg/10 3 ni 2 /h (lb/10 3 ft 3 /hr) 9.8 (2.0) 31.2 (6.4) 86.4 (17.7) 32.2 (6.6)
Volumetric, kg/10 3 m 3 /h (lb/10 3 ft 3 /hr) 97.2 (6.0) 415.0 (25.9) 790.0 (49.2) 334.0 (20.8)
Liquor loading:
Surface area, m 3 /10 3 m 2 /h (gal/10 3 ft 2 /hr) 7.9 (0.2) 4.9 (0.1) 10.8 (0.3) 3.6 (0.1)
Volumetric, m 3 /10 3 m 3 /h (gal/10 3 1t 3 /hr) 79 (0.6) 55 (0.4) 99 (0.7) 38 (0.3)
Air loading:
Surface area, m 3 /m 2 /h (cu ft/ft 2 /hr) 0.47 (1.54) 1.22 (4.00) 0.52 (1.70) 0.31 (1.02)
Volumetric, m 3 /rn 3 /h (cu ft/cu ft/hr) 4.7 (4.7) 16.2 (16.2) 4.7 (4.7) 3.2 (3.2)
Na 2 S inlet, g/l 1.20 6.40 8.00 8.90
Na 2 S outlet, g/l 0.03 1.30 1.60 2.50
Oxidation efficiency, % 98 80 80 72
Oxygen ratio, act./theor. 33.0 25.0 3.9 6.2
-------
FIGURE 9-4
EXHAUST
AIR
BLACK
LIQUOR
EXIT
AGITATED AIR SPARGING SYSTEM FOR BLACK LIQUOR OXIDATION
Two major problems noted with agitated air sparging units for weak black liquor oxidation
are foaming caused by either low weak liquor solids content or by a wood furnish of greater
than 10 percent pine (27) and mechanical breakdown of the liquid agitator system.
Methods employed for foam control are:
1. Placement of mechanical foam breakers at the top of the tank to break up stable
foam, requiring 0.08 to 0.16 kW per metric ton per day (0.1 to 0.2 hp per short
ton per day).
2. Use of chemical or kerosene defoamer in the black liquor,
3. Recycling of strong black liquor to the weak black liquor to increase the net
liquid viscosity, and
4. Reducing retention time with resultant lowered oxidation efficiency. A summary
of performance data for two units is presented in Table 9-4 (4).
FOAM
BREAKER
/
BLACK
LIQUOR
INLET
AIR
AGITATOR
/
TURBINE
AERATOR
9-10
-------
TABLE 9-4
OPERATING AND PERFORMANCE DATA FOR AGITATED AiR
SPARGED WEAK BLACK LIQUID OXiDATION SYSTEMS (4)
Parameter Mill A Mill B
Liquor flow, rn 3 /h (gpm) 243 (1,070) 262 (1,150)
Air flow, m 3 Ih (cErn) 8,500 (5,000) 4,250 (2,500)
Tank volume, m 3 (ft 3 ) 200 (7,060) 200 (7,060)
Retention [ tine. ruin. 49 46
Na 2 S loading, kg/m 3 /h (Ib/1t 3 /hr) 3.64 (0.23) 8.75 (0.55)
Liquor loading, rn 3 /in 3 /h (gal/1t 3 /hr) 1.2 (9) 1.3 (10)
Air loading, m 3 /rn 3 /h (ft 3 /ft 3 /hr) 42.5 21.2
Na 2 5 inlet, gIl 3.0 6.7
Na 2 S outlet, g Il 0.2 0.2
Oxidation efficiency, % 93 97
Oxygen ratio, act./theor. 8.0 2.6
Power required:
Agitator, kW/t/day (hp/ton/day) 0.14 (0.17) 0.10 (0.12)
Foam breakers, kW/t/day (hp/ton/day) 0.16 (0.19) 0.10 (0.12)
9.1.5 Rotating Fluid Contactors
A limited amount of experience has been obtained with the Asheroft dual vortex eon [ actor,
primarily in polishing weak black tiquor to upgrade existing units. The system introduces
weak black liquor tangentially into an axially downward flowing pipe of small diameter. Air
is introduced to the liquid tangentially at the point where the small pipe connects with a
wider pipe, resulting in large shearing forces and rapid gas-liquid mixing (28).
Preliminary results indicate that the system is suitable for weak black liquor oxidation with
nonresinous wood species, but excessive foaming occurs for pine black liquor. One system
employed to upgrade the performance of an existing weak liquor oxidation system feeds air
into an Ashcroft unit at 5,000 m 3 /h (3,000 efm) with the black liquor flowing at 227 m 3 /h
(1,000 gpm). Installation of the unit results in increasing the overatl Na 2 S oxidation
efficiency during weak black liquor oxidation from the former 50 to 60 percent to between
95 and 99 percent.
9.2 Strong Black Liquor Oxidation—Air
Strong black liquor oxidation with air, following multiple-effect evaporation, is employed at
mills that pulp substantial quantities of resinous pine wood species to alleviate potential
foaming problems, particularly in the southeastern United States. Strong black liquor
9-11
-------
oxidation reduces malodorous sulfur gas emissions from the direct contact evaporator and
counteracts the tendency for Na 2 S to reform. The major types of systems employed for
strong l)laCk .liquor oxidation are single. and two.stage unagitated air sparging. Agitated air
spargers, plug flow reactors, and dual vortex contactors also are used for strong black liquor
oxidation to a limited extent.
9.2.1 Single Stage Unagitated Air Sparging
The major technique employed for strong black liquor oxidation, to date, is a completely
mixed unagitated air sparging tank. Hawkins (29) (30) describes the development of the
original system at the Pasadena, Texas, mill of the Champion Paper Company. The system
employs aeration and dcaeration tanks arranged in series. The black liquor is oxidized by
sparging s1th air in a single stage aeration tank, followed by separation of entrained air
bubbles from the liquid in the deaeration tank. The aeration tank consists of a cylindrical
section mounted between two conical sections, as shown in Figure 9.5.
Air is introduced in th bottom of the cylindrical section of the aeration tank through a
series of eight radially branching sprayer arms in a Christmas tree arrangement. The sprayer
arms arc constructed of 20 cm (8 in) pipe with 19 nozzlcs for air outlet per branch and are
connected to a central header. Each nozzle is 3.8 cm (1.5 in) in diameter. The air is caused
to deflect downward against a deflector plate to achieve a fanning air curtain effect.
Air is introduced to the central header at a rate of 10,000 m 3 /h (6,000 cfm). The exhaust
air is drawn through a cyclone separator to remove entrained foam and liquor droplets,
which are then returned to the aeration tank. The black liquor is introduced to the top of
the tank through a 15 cm (6 in) diameter pipe located above the liquid level. The liquor is
sprayed in from the bottom of the pipe through a series of 2.5 cm (1 in) holes evenly spaced
along the pipe to obtain even liquid distribution and to provide a means for controlling
foam. The black liquor is withdrawn from the bottom of the tank and passed to two
deacration tanks for gas.liquid separation to prevent pump malfunctions.
A Na 2 S oxidation efficiency of 97 to 98 percent occurs at an inlet concentration of 30 gIl
and at a liquid retention time of 2.5 hr. The process results in reducing H 2 S emissions by 90
percent from the recovery furnace following direct evaporation as compared to 1-12 S
emissions from unoxidized black liquor.
Blosser and Cooper (4) have prepared an extensive review of strong black liquor oxidatIon
practices at kraft pulp mills, it is necessary to provide 3- to 5- times the stoichiometric
amount of air for oxidation of Na 2 S to sodium thiosulfate (Na 2 S 2 03). Air-to.liquid flow
ratios must normally be greater than I 10 to 190 m 3 air/rn 3 strong black liquor (15 to 25
ft 3 /gal of black liquor.) It is normally necessary to maintain a minimum liquor depth of 2.7
to 3.6 m (9 to 12 It) with a minimum liquor retention time of 120 to 150 minutes to
9L2
-------
EXIT GAS
LEGEND
____ 1 HEAVY LIQUOR STORAGE TANK
2 LIQUOR INLET NOZZLES
_____ _____ 3 OXIDIZED LIQUOR OUTLET
4 BLOWER
5 __
5 AIR SPARGER
_____ ____ 6 CYCLONE SEPERATOR FOR
BLACK
____ I BLACK AIR DISCHARGE
LIQUOR
‘ LIQUOR 7 OVERFLOW
INLET
EXIT 8 HEAD TANKS
AIR
COMPRESSOR
FIGURE 9-5
CHAMPION UNAGITATED AIR SPARGE STRONG BLACK LIQUOR OXIDATION SYSTEM (29)
-------
provide for effective oxidation of the Na 2 S. The product of the oxygen ratio (actual oxygen
addition rate to theoretical oxygen addition rate) and liquor retention Lime in minutes must
be 600 or greater to reduce exit Na 2 S levels to less than 0.5 g/l in [ lie oxidation tower exit
liquid. The product of oxygen ratio and retention time inusi be greater than 1 ,000 to reduce
Na 2 S levels in [ lie exit liquor lo less than 0.2 gIl, as shown in Figure 9-6. These findings
point to the possible development of two-stage systems for strong black liquor oxidation
with air.
Morgan (31) (32) reports on the results of a study on a new strong black liquor oxidation
system. He finds that the efficiency of strong black liquor oxidation is a function of liquor
height, air flow rate, inlet Na 2 S concentration, and retention time. The rate of Na 2 S
oxidation increases with increasiiig Na 2 S concentration and air flow rate, decreasing
retention time, and decreasing liquor height. Na 2 S oxidation efficiencies of up to 99 percent
arc observed, but oxygen transfer efficiency is relatively low so that large quantities of
excess air (3.5 Limes theoretical or more) arc required to achieve high degrees of oxidation.
Foam control is particularly a problem during strong black liquor oxidation with air at kraft
pulp mills in the southeastern United States. Several control methods are available. Effective
soap removal during multiple-effect evaporation upstream of the oxidation unit removes
6
C.)
z
w
80Z3
U-
U-
.TOW
z
6O
N
0
z
5
4
3
2
0
X RETENTION TIME (minutes)
FIGURE 9-6
OPERATING & PERFORMANCE DATA FOR SINGLE STAGE STRONG
BLACK LIQUOR OXIDATION SYSTEM (4)
OXYGEN SUPPLY
REQ’T
-J
(I)
CiJ
0
2
-J
U)
w
100
0 200 400 600 800 1000
40
9-14
-------
foam-producing materials (33). Cyclone separators in the exhaust gas line remove entrained
foam and liquor droplets. Provision of adequate height above the liquor level allows for
foam dissipation. Two to 5 in (6 to 1 5 ft) is the normal minimum required.
Mechanical and chemical methods of foam dissipation also arc employed. Mechanical foam
breaking requirements for strong black liquor oxidation systems arc aeration tank only—0 to
0.022 kW per daily metric ton of pulp (0 to 0.027 hp/ton pulp per day) and aeration plus
deacration tank—0 to 0.007 kW per daily nictnc ton of pulp (0.0 to 0.1 hp/ton pulp per
day). Chemical defoaming agents used include diesel oil and kerosene in dosages from 50 to
250 I/rn 3 (50 to 250 gal/I ,000 gal) strong black liquor, primarily on an intermittent basis.
Mechanical foam breaking has not proved stiitahle in many cases where there are
considerable foaming problems. The cost of chemical defoaming runs as high as 0.55 to
$1 .65/t ($0.50 to $1.50/ton) of pulp.
9.2.2 Two Stage Unagitatcd Air Sparging
Limitations in achieving Na 2 S oxidation efficiencies above 99 percent to reduce exit
concentrations of Na 2 S to less than 0.1 g/l with single stage units have led to the
development of multiplc.stage strong black liquor oxidation systems. Padlield (34) reports
on efforts to upgrade the single stage strong black liquor oxidation sysLcm from the former
97-98 percent Na 2 S oxidation efficiency to a desired 99-100 percent. A second aeration
lank of design similar to the first placed after the initial aeration tank provides for
additional Na 2 S oxidation. About a 60 minute retention time is provided in each oxidation
tank with an additional 30 minutes of deacration in the bottom of each tank to facilitate
liquor pumping. The exit liquor from the second stage oxidation tank is sent directly to the
direct-contact evaporators to lrcvcnt reversion to sodium sulfide.
The system has two air blowers for adding air to the black liquor, one for each tank. The
blower for the first oxidation tank provides 10,400 m 3 /hr (6,000 cfm) of air at 143 kP i
total pressure (]41 atmospheres or 6 psig) with a 187 kW (250 hp) motor. The blower for
the second oxidation tank provides air at 8,500 m 3 /hr (5,000 elm) at the same pressure and
using the same power. Two 19 kW (25 hp) mechanical foam breakers are required, one on
each tank. Total power requirement for the two stage system is thus 410 kW (550 hp) for a
pulp production of 770 Mg (850 tons) per day, or about 0.53 kW per daily metric ton of
pulp (0.65 hp/ton/day). The system is illustrated in Figure 9-7 (34).
The average overall Na 2 S oxidation efficiency for the system averages 99.95 percent,
resulting in an average Na 2 S concentration in the exit black liquor of 0.02 gIl. Emissions of
H 2 S from the recovery furnace arc about 2 ppm by volume as compared to 50 ppm with the
original single stage system.
9-15
-------
BLACK BLACK
LIQUOR LIQUOR
EXIT INLET
FIGURE 9-7
CHAMPION TWO STAGE UNAGITATED STRONG BLACK LIQUOR
OXIDATION SYSTEM (34)
9.3 Agitated Air Sparging
The completely mixed agitated air sparging units employed for strong black liquor oxidation
arc of the same design as those described in section 9.1.4 for weak liquor systems. One unit
introduces 10,400 m 3 air/h (6,000 cfm) into a black liquor flow of 80 m 3 /h (350 gpm) at 48
percent solids using 1 50 kW (200 hp) agitated turbine aerator (4). The aeration tank has a
diameter of 9.5 m (31 ft) and a height of 7.9 m (26 ft), and is operated at a liquid depth of
3.7 m (12 ft). Five 19 kW (25 hp) mechanical foam breakers are installed on the top of the
unit, but have severe maintenance problems. No definitive results regarding efficiency of the
unit are available because of the lack of sufficient operating experience. The system requires
frequent maintenance because of agitator and foam breaker malfunctions.
9.4 Combination Systems
Tobias and Robertson (35) describe the development of an agitated concurrent plug-flow
reacLor system for oxidizing of strong black liquor with air. The primary purpose of the
system is to polish the strong black liquor following multiple-effect evaporation to
counteract the reversion to Na 2 S of oxidized weak black liquor in an existing system. The
system employs one oxidation unit on the exit pipe of the strong black liquor storage tank,
and the other on the strong black liquor recirculation line, as shown in Figure 9-8. Each
reactor is located so that a pipe with a central axially located baffle forms two mixing
chambers, each with an agitator. Two parallel pipes are located immediately upstream of the
agitators in both chambers, with holes drilled so as to obtain a fanning air curtain effect.
EXIT
GAS
BLOWER PUMP BLOWER PUMP
9-16
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BLACK
INLET
60% HEADER
40% HEADER-S
-J
VENTURI
RECIRCULATION LINE
—... IN-LINE
OXIDIZER
RETENTION TANK
(RECIRCULATION)
CYCLONE
LC V
CYCLONE
SEPARATOR
STRONG BLACK
LIQUOR TANK
FIGURE 9-8
WESTERN KRAFT PIPELINE REACTOR STRONG BLACK LIQUOR OXIDATION SYSTEM (35)
-------
The air bubbles arc sheared into smaller bubbles by the rotating action of the agitators to
increase gas-liquid interfaeial contact area, turbulence, and mixing.
Preliminary results indicate that the Na 2 S oxidation efficiency for the system is virtually
100 percent with both units operating at inlet concentrations of 3 gIl or less. The oxidation
efficiency is practically I 00 percent with the system on the storage-tank exit line alone at
Na 2 S concentrations of 1.3 gIl or less. The system also appears to have higher oxygen
utilization efficiencies than conventional air sparged units. No major problems with foaming
are observed, but the wood furnish of the mill is primarily oak hardwoods. The two phase
flow system reduces liquor line plugging but also causes substantial line vibration, requiring
secure fastening of the pipes. When the system was installed at a second mill, the operation
of the recovery furnace bed became unstable, possibly because of sulfate formation with
resultant pH reduction and subsequent lignin precipitation from the black liquor.
Martin (36) describes reversion phenomena during multi-effect evaporation and also
describes a secondary strong black liquor oxidation system installed downstream of an
existing high efficiency weak black liquor oxidation unit. The system injects 680 m 3 /h
(400 cfm) of air into the bottom of an existing storage tank through a series of eight vertical
pipes connected to a central header, where the black liquor flow is about 68 to 80 m 3 /h
(300 to 350 gpm). Results indicate that the Na 2 S concentration of the black liquor entering
the direct contact evaporator could be maintained at below 0.005 g Il with thc inlet Na 2 S
concentration below 1.5 WI.
9.5 Molecular Oxygen Systems
The severe foaming problems accompanying oxidation of weak black liquor with
atmospheric oxygen (air) at kraft pulp mills pulping pine woods in the southeastern United
States led to considering use of molecular oxygen as an alternative. The foaming problems
of air oxidation can be alleviated by using molecular oxygen because the inert
nitrogen-argon diluting medium is no longer present. The major drawback to using oxygen
for black liquor oxidation has been its cost, but recent trends arc toward lower oxygen
prices, thus making it a more competitive alternative. Additional possible uses of oxygen
include pulp bleaching, chemical pulping, wastewater treatment, and addition to recovery
furnace firing zones for odor control. Such applications may lead to further increases in
oxygen consumption at a given mill.
9.5.1 Preliminary Studies
Early laboratory investigations of Na 2 S oxidation in black liquor with molecular oxygen
have been made by Bergstrom and Trobeck (37), Venemark (38), Ricca (39), Sakhuja.and
Bosu (40), and Miller (41). Major findings of these laboratory studies are that the sulfide
oxidation occurs in more than one rate-limiting regime, that Na 2 2 03 is the major reaction
9-18
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product, but that varying amounts of Na 2 SO 4 are formed. Thc reaction rate is influenced
by temperature and catalyzed by the presence of organic constituents in the black liquor.
Cooper (42) reports on weak and strong black liquor oxidation with molecular oxygen in a
series of pilot scale experiments. Design criteria are listed in Table 9-5 for plug flow reactor
systems.
TABLE 9-5
DESIGN CRITERIA OF PLUG FLOW REACTOR SYSTEMS FOR
BLACK LIQUOR OXIDATiON WITH MOLECULAR OXYGEN (42)
Weak Strong
Black Liquor Black Liquor
Criteria Oxidation Oxidation
Performance:
Na 2 S efficiency, % 99+ 99+
Na 2 S outlet, gIl 0.01-0.02 0.01-0.02
Liquor:
Reynolds number 100,000 10,000-20,000
Velocity, rn/s (ft/see) 1.5-4.5 (4.9-14.8) 1.0-3.0 (3.3-9.8)
Temperature, °C (°F) 77-88 (170-190) 110-115 (230-239)
Liquid pH 12.0-12.6 12.0-13.0
Solids,%bywt. 15-17 48-51
Oxygen:
Oxygen ratio, Act/theoret. 1.1-1.3 1.2-2.5
Total pressure, atm 3.0-4.4 3.0-5.1
Partial pressure, % purity 90-99.5 90-99.5
Retention time (Liquor basis):
Piping section, seconds 60-120 60-180
Storage tank, minutes 15-45 30-60
Both weak and strong black liquor oxidation systems employ plug flow reactors followed in
series by tall storage tanks, as shown in Figure 9-9.
9.5.2 Weak Black Liquor Oxidation
Weak black liquor oxidation of lightly resinous pine black liquors with oxygen allows
stabilization of Na 2 S without causing the excessive foaming as with air. The process also has
a lower capital cost than comparable air units because it can be carried out in a plug flow
reactor within the piping of an existing mill. The use of oxygen would probably not
9-19
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OXYGEN
EXHAUST
GAS
I
—.-I
STRONG BLACK
LIQUOR TO
DIRECT CONTACT
EVAPORATORS
FIGURE 9-9
WEAK BLACK
LIQUOR FROM
WASHERS
EXHAUST
GAS
1 ri
! i i
! I i
! j i
! I
L.J L.
.—.—I
MOLECULAR OXYGEN
WEAK BLACK LIQUOR
—• — STRONG BLACK LIQUOR
TWO STAGE COMBINATION WEAK & STRONG BLACK LIQUOR OXIDATION WITH OXYGEN
-------
overcome the problems of Na 2 S reversion in oxidized black liquor unless the process were
carried out at a temperature above 1300 C (266° F).
Kosaya (43) describes the first reported use in 1956 of molecular oxygen for weak black
liquor oxidation at the Kehra pulp mill in the Soviet Union. The system was originally
installed to reduce the H 2 S emission from mulLiple-effect evaporators, with the aim of
improving the water quality and reducing odor. The pulp mill was a kraft mill that did not
use recovery furnace flue gas for direct contact evaporation. The system injects oxygen into
the black liquor upstream of the multiple effect evaporators through a “dosing apparatus.”
Results indicate that it is possil)lc to achieve “essentially complete” oxidation of the Na 2 S
in the black liquor at an inlet concentration of 7.5 gIl within “several minutes of retention
time” in the pipe at a temperature of 70° C (158° F). No problems with foaming were
observed, and absorption of oxygen in the pipeline reactor was virtually complete.
The study indicates that the oxygen consumption is about 12 percent greater than the
stoichiometric amount for conversion of Na 2 S to Na 2 S 2 0 3 , as shown in Table 9-6. An
overall finding is that oxidation of weak black liquor with molecular oxygen is potentially
attractive from an economic standpoint.
TABLE 9-6
OXYGEN REQUIREMENTS FOR WEAK BLACK LIQUOR
OXIDATION* (43)
Dimension Theoretical Actual
m 3 ofO 2 /tofNa 2 S 315 370
(cu ft of 0 2 /ton of Na 2 S) (10,080) (11,850)
kgof O 2 Itof pulp 24 58
(lb of 0 2 /ton of pulp) (48) (116)
m 3 of 0 2 /m 3 of black liquor* 2.1 2.5
(cu ft of 0 2 /gal of black liquor) (0.28) (0.33)
*BlaCk liquor conditions. 15% solids, 7.54 g Na 2 SI I.
**Based on black liquor flow of 8 m 3 /t pulp (1900 gal/ton pulp).
Freedman (28) reports on the oxidation of weak black liquor with molecular oxygen in
1970, when an Ashbrook rotating cyclonic fluid contactor was used to provide contact
between the oxygen and black liquor. The system introduces black liquor tangentially into a
9-21
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small diameter pipe and redirects its flow axially to form a centrifugal cyclonic liquid flow
pattern. Oxygen is then introduced tangentially into the liquid at the point of concentric
expansion into a larger diameter pipe. The violent mixing of the fluids causes a large number
of small oxygen bubbles to form. Collectively, these bubbles will provide a large interfacial
contact area between the gas and the liquid.
Pilot studies, using the contactor for oxidation of weak black liquors with oxygen, indicate
that effective oxidation of Na 2 S can be achieved with liquors from both northern
hardwoods and southern pine woods. Excessive foaming does not occur in either case
because of the absence of the argon-nitrogen diluent from the incoming gas stream.
Essentially complete oxidation of sodium sulfide can be obtained within 15 to 60 seconds
for hardwood black liquors, and within about 60 seconds for southern pine black liquors.
Oxygen usage efficiencies of 85 to 90 percent are obtained in both cases, making the
technique economically favorable. The exothermic oxidation also warms the black liquor by
up to 10° C (18° F), thus decreasing evaporator stream heating requirements.
Galeano and Amsden (44) report on an extensive study of using molecular oxygen tor weak
black liquor oxidation at the Owens-Illinois, Inc. kraft pulp mill in Orange, Texas, where
southern pine wood species are primarily pulped. Basically, the system introduces molecular
oxygen into a flowing stream of black liquor in a two-step pipeline reactor with a retention
time of 15 to 40 seconds, followed by tank storage with a retention time of 8 to 12 hours.
Oxygen gas enters the 61 cm (24 in) diameter horizontal inlet pipe at a total pressure of
0.45 MPa (50 psig) at a flow rate of 850 to 1,360 m 3 /h (500 to 800 elm) through an
injector oriented perpendicular to the liquid, it is broken up into small bubbles under high
turbulence by a venturi effect. The black liquor enters the inlet pipe at a flow rate of 227 to
386 m 3 /h (1,000 to 1,700 gpm) with a Na 2 S concentration of 9 to 12 g/l, a solids content
of 12 to 15 percent, and a temperature of 93 to 1000 C (200 to 230° F). After contacting
the oxygen, the black liquor flows through a 30 m (100 ft) length of 25 cm (10 in) diameter
pipe, where the oxygen is absorbed and the Na 2 S oxidized. The initial contact section
provides for a liquid Reynolds number of 100,000 to 200,000 in highly turbulent flow. The
system is diagramed in Figure 9-10.
Galeano and Amsden observed Na 2 S oxidation efficiencies of 85 to 98 percent (94 average)
without excessive foaming with an oxygen usage efficiency of 75 to 95 percent (91 average).
About 90 percent (of the total 94 percent) of the Na 2 S is oxidized to Na 2 S 2 03 within the
pipeline reactor section. There is a subsequent conversion of about 15 percent of the
Na 2 S 2 03 to Na 2 SO 4 , as listed in Table 9-7.
Not all the sulfur is accounted for by the above chemical analyses alone, indicating ither
the formation of other products, such as polysulfide, sulfite, or polythionate ions, or loss by
gasification. The initial Na 2 S oxidation reaction is extremely rapid. Additional changes in
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OXYGEN
(.
STRONG
BLACK
LIQUOR
STORAGE
TA N K
OX EN
j I
! !
! !
! !
_.) ____ STRONG BLACK
LIQUOR T0
DIRECT CONTACT
EVAPORATOR
FIGURE 9-10
OWENS-ILLINOIS SYSTEM FOR TWO STAGE WEAK & STRONG BLACK LIQUOR
OXIDATION WITH OXYGEN (45)
WEAK
BLACK
LIQUOR
EXHAUST
GAS
I
-------
TABLE 9-7
EFFECT OF WEAK BLACK LIQUOR OXIDATION ON LIQUID
CHEMICAL COMPOSITiON (44)
Chemical Component
Location Na 2 S Na 2 S 2 O 3 Na 2 SO 4 Total Account
gil gil gil gil %
Reactor inlet 4.56 0.80 0.53 5.89 100.0
Reactor outlet 0.45 4.35 0.70 5.50 93.7
Storage outlet 0.]9 4.61 0.63 5.43 92.8
black liquor during the oxidation process are a rise of 0.1 to 0.3 pH units because of
thiosulfate formation, and a temperature rise of 5 to 8° C (10 to 15° F) caused by the
exothermic oxidation.
Galeano and Amsden observed several benefits of weak black liquor oxidation with
molecular oxygen as compared to lack of oxidation, namely:
I. H 2 S emissions during direct contact evaporation declined by 95 to 99 percent
during multiple-effect evaporation,
2. Na 2 SO 4 makeup requirements declined by 30 kgit pulp (60 lb/ton pulp),
3. Evaporator condensate water quality improved, and
4. Tall oil yield increased by about 15 percent.
The effects of weak black liquor oxidation on sulfur gas emissions from various process
sources arc listed in Table 9-8.
The mill is located adjacent to an oxygen pipeline where the low oxygen cost of $9.35/t of
02 ($8.50/ton) results in a maximum net mill operating cost of $0.06 to $0.09 per metric
ton of pulp ($0.05 to $0.08/ton). Projected capital costs for similar oxygen generating
installations arc $55 to $83 per daily metric ton of pulping capacity ($50 to $75 per daily
short ton).
9-24
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TABLE 9-8
EFFECT OF WEAK BLACK LiQUOR OXIDATION ON SuLFUR GAS EMiSSIONS
FROM KRAFT MILL PROCESS SOURCES (44)
Concentration
Source Constituent Unoxidized Oxidized Rcdudion
ppm, by vol. ppm, by vol. %
Cycloneevaporatorexit H 2 S 191 10 95
RSH 161 5 97
RSR+RSSR 20 0 100
So 2 243 241 1
Weak black liquor storage H 2 S 553 3 99
tank vcnL RSH 760 375 50
RSR I ,274 280 78
Evaporator noncondcnsablc H 2 S 1,208 4 99
gas vent RSH >>500 300 >50
EvaporaLor condensate Sulfide 72 28 60
BOD 863 530 38
9.5.3 Strong Liquor Oxidation
Very little work has been devoted to the use of molecular oxygen for strong black liquor
oxidation. Galcano and Amsden (45) describe the use of strong black liquor oxidation with
oxygen for polishing to counteract reversion to Na 2 S from the previously oxidized weak
black liquor. The system introduces the oxygen into the strong black liquor in an expanded
pipe section at two places. The black liquor then flows downward to take advantage of the
gas bubble “holdup” phenomenon with its attendant increased effective retention time, and
then through a horizontal pipe section to a strong black liquor storage tank. The liquid
Reynolds number in the pipeline reactor ranges from 10,000 to 20,000 in the liquid at 53
percent solids, at a temperature of 120° C (250° F), and with a liquid retention period of 30
to 50 seconds.
Initial studies indicate potential problems with black liquor cooling, ligniri oxidation, and
incomplete oxygen mass transfer into the black liquor, resulting in incomplete Na 2 S
oxidation. Little Na 2 S reversion in the oxidized weak black liquor is observed, possibly
because of the high inlet temperatures, 93 to 110° C (200 to 230° F), of liquor from the
continuous digestcrs. The liquid retention lime in each pipeline reactor stage is about one
9-25
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minute. The result is a reduction in the Na 2 S concentration in the strong black liquor from
1 .5 gIl to between 0.05 and 0.10 gIl.
9.5.4 Digester Injection
It is possible to introduce oxygen into the kraft recovery system at the digester at the end
of the cook. Oxidation of the Na 2 S and sodium mercaptide (NaHS) at the end of the cook
prevents gaseous conversion and release of these compounds into the atmosphere. The
technique has the advantages of performing the oxidation in an enclosed reactor, thus
assuring complete use of the oxygen added, and performing the operation at high
temperatures to assure maximum reaction rates and minimal possibilities for reversion to
Na 2 S. The technique can also result in oxidation of organic sulfur compounds to prevent
their discharge and to limit emissions from the digester, washers, multiple-effect
evaporators, tall oil and storage tank vents, and the direct contact evaporators.
Possible digester corrosion and excessive oxygen consumption can possibly occur because of
competing side reactions, such as sulfate formation and lignin oxidation and potential
degradation in pulp quality. Two studies investigated addition of molecular oxygen to kraft
mill digesters at the ends of cooks to oxidize the Na 2 S and NaHS present in the black
liquor.
Fones and Supp (46) report on the addition of oxygen to a kraft digester to oxidize black
liquor in 1960. Oxygen is added at the ends of successive cooks to a pressure vessel
containing pulp alone, black liquor alone, and a mixture of pulp and black liquor. For the
test cook with pulp alone, the lignin content of the pulp is reduced by oxidation, and its
brightness is increased.
When oxygen is injected into a digester containing both pulp and black liquor at the end of
the cook, the Na 2 S is rapidly and completely oxidized. The high reaction temperatures and
pressures result in the formation of substantial quantities of Na 2 SO 4 with a resultant drop
in liquid pH and a sharp increase in the amount of oxygen consumed. In addition, the
oxidation process also reduces the bursting strength and brightness of the pulp, and its lignin
content. Oxidation of the black liquor with oxygen at the end of the cook is uneconomical
because of the excessive oxygen requirements and the detrimental effects on pulp quality.
Tests were made to determine the effect of oxygen addition on the oxidation of Na 2 S in a
pressurized vessel containing only black liquor. Oxygen was added to black liquor alone in a
digester at a total pressure of 0.8 MPa (100 psig) and 150° C (302° F) in a series of stages to
recirculated black liquor. The oxygen then reacted with the Na 2 S and other constituents
until the pressure returned to its initial value in a time span of four hours. Results indicate
that the initial reaction product was Na 2 S 2 O 3 , but substantial quantities of Na 2 SO 4 were
also present causing a resultant drop in liquid pH. Because oxygen consumption was about
9.26
-------
double that for stoichiometric convcrsion of Na 2 S to Na 2 S 2 O 3 , the process proved
prohibitively expensive with oxygen costing $22/t ($20/ton).
Kringsted and McKcan (47) describe oxidation of Na 2 S and NaHS in black liquor at the end
of a draft cook in the presence of pulp. Their results indicate that by oxygen addition Lo a
kraft digester at the end of a cook, the Na 2 S concentration in black liquor can be reduced
by 90 percent and NaHS by 99 percent but twice the stoichiometric amount of oxygen is
required.
An additional finding is that no sodium polysulfide or Na 2 SO 3 could be detected in the
black liquor during thc oxidation period, indicating that it might be possible to reduce or
eliminate potential problems caused by reversion to Na 2 S. No measurements for Na 2 SO 4
were made during the tests.
The process does not affect the pulp strength or yield, but does make the pulp easier to beat
and also reduces its brightness. Oxygen addition to a digester at the end of a kraft cook can
effectively oxidize Na 2 S and NaHS, and can minimize potential problems caused by
reversion. Of particular importance is that the relatively long retention time of about 20
minutes at a high temperature of approximately 180° C (356° F) can result in the possible
oxidation of substantial amounts of Na 2 S 2 O 3 to Na 2 SO 4 , and also of lignin. The
nonselective oxidation process can require excessive quantities of oxygen in assuring
oxidation of the Na 2 S and NaHS, thus making the process economically unattractive.
9.6 Process Effects
Black liquor oxidation influences the operation of the kraft chemical recovery system in
other ways in addition to reducing malodorous sulfur gas emissions. Major influences
include:
1. Improvement in multiple.effect evaporation through reduced scaling on heat
transfer surfaces,
2. Reduced corrosion rates of metal evaporating surfaces,
3. Possible increases in tall oil yield,
4. Reduced chemical makeup requirements for Na 2 SO 4 and calcium oxide (CaO),
5. Possible increases in green and white liquor sulfidities, with resultant effects on
pulp yield and quality, and
6. Possible effects on black liquor heating values.
9-27
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9.6.1 Evaporator Scaling
Berry (48) finds that the H 2 5 evolved from unoxidized weak black liquor causes the
formation of an iron sulfide (FeS) scale, which inhibits [ he rate of heat transfer in the
multiple-effect evaporators. The precipitation of lignin on the heat transfer surfaces can be
minimized by maintaining a sufficiently high liquid pH, thus alleviating another potential
source of evaporator scaling. Any Na 2 SO 4 formed during weak liquor oxidation becomes
less soluble as liquid temperature increases and can CULI5C a scaling problem during the latter
stages of evaporation at higher solids concentrations.
9.6.2 Corrosion During Evaporation
Weak black liquor oxidation can substantially reduce the corrosion of multiple-effect
evaporator surfaces, particularly in the vapor shell sections. Von Esscn (49) reports that the
corrosion of evaporator surfaces is caused primarily hy the formation of FeS from H 2 S in a
moist atmosphere on metal surfaces. The rate of corrosion is reduced by about 85 percent
by reducing the inlet Na 2 S concentration in black liquor to less than 3 g/l. Cyr and 1-larper
(50) report that the average life of multiple-effect evaporator tubes is substantially increased
by weak black liquor oxidation.
9.6.3 Tall Oil Yield
Weak and strong black liquor oxidation can increase tall oil yields for byproduct recovery
by either physical or chemical mechanisms (33). Foaming, however, could be a serious
problem with air oxidation of pine black liquors where large quantities of tall oil arc
obtained. Galeano and Amsden (44) report that tall oil yield is increased by approximately
15 percent as the result of weak black liquor oxidation with molecular oxygen. Also, there
is no apparent decrease in tall oil quality and there arc increases in yield of as much as 7.5 to
12.5 kg per metric ton of pulp (15 to 25 lb/ton). Rippee (19) reports an approximate 10
percent increase in tall oil yield at a western U.S. kraft pulp mill achieved by oxidizing weak
black liquor with air and recycling strong black liquor for foam control. At a 1972 National
Council symposium (51), black liquor oxidation was reported to result in increased tall oil
yields. But the quality of the tall oil decreased, possibly because of its oxidation.
9.6.4 Chemical Recovery
Black liquor oxidation reduces the sulfur losses from kraft process sources, resuItin in
decreased Na 2 SO 4 and lime chemical makeup requirements. Calcano and Amsdcn (44, 45)
report that to maintain a given sulfidity level the Na 2 SO 4 makeup rate can be reduced by
15 to 30 kg per metric ton of pulp (30 to 60 lb/ton) of pulp. To provide a portion of the
sodium makeup requirements, NaOFI can be used. The lime makeup requirement can he
reduced by 1.0 to 1.5 kg per metric ton of pulp (2-3 lb/ton), and the total lime mud
9-28
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processing rate can be reduced by 5.0 to 7.5 kg per metric ton of pulp produced (10 to
15 lb/ton). Specific chemical savings for maintaining particular sulfiditv levels will var ’
between mills.
9.6.5 Liquor Sulfidity
A major effect of black liquor oxidation on the kraft pulping process is the increase in green
and white liquor sulfidities. The reason for the increased sulficlity levels is the retention of
additional sulfur in the recovery system; less H 2 S is lost to the recovery furnace flue gas
stream during the direct contact evaporation (52). The increase in green liquor sulfidity
results in a decrease in 1)0th lime makeup and lime burning requirements, with a resultant
decrease in fuel requirements at the lime kiln (44).
Black liquor oxidation can result in increased white liquor sulfidity levels of two to five
percent or more for a given Na 2 SO 4 chemical makeup rate, or a reduction of 15 to 30 kg
per metric ton of pulp (30 to 60 lb/ton) to maintain a given sulfidity level (14). The
increased white liquor sulficlity levels can increase the rate of clclignification during digestion
of wood chips. The greater sulfidities also can increase the emission of malodorous organic
sulfur compounds from digester blow and relief gases, brown stock washer vents, and
multiplc.effect evaporator gases (53).
Shah and Stephenson (L4) observed that the installation of weak black liquor oxidation
results in a 2. to 5.percent increase in white liquor sulfidity for given chemical makeup rates.
Increases in pulp quality and yield were observed during the digestion process, but they
probably are offset, at least in part, by increased corrosion eaused by higher Na 2 S levels.
The increased sulfidity may cause increased malodorous sulfur gas emissions from the kraft
recovery system because of the higher sulfur circulation rates.
One drawback to the usage of black liquor oxidation is that it can cause the inlet Na 2 S
concentrations in the black liquor to risc sharply, thus overloading the oxidation capacity of
the existing system. Ritchcy (54) reports that installation of a strong liquor oxidation
system using air at an existing mill resulted in about a 50 percent increase in the inlet Na 2 S
concentration in the unit, that is, from about 40 to 60 g/l. The equivalent sulfidity increase
was approximately 10 percent, resulting in a reduction in Na 2 S oxidation efficiency from
the expected 98 percent to about 80 percent. To maintain proper sodium-sulfur ratios in the
recovery system it is necessary to add NaOH as part of the sodium makeup.
9.6.6 Energy Balances
Black liquor oxidation affects energy conservation in the kraft recovery system. Lime kiln
fuel requirements are reduced with increasing liquor sulfidity levels (see section 9.6.5).
Oxidation of black liquor results in a reduction in net heating value of black liquor because
9-29
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of the oxidation of both the reduced sulfur compounds and the organic lignin fuel materials.
Losses in heating value can range from 2 to 6 percent (55). The oxidation of black liquor
can act to maintain cleaner heat transfer surfaces, thus reducing steam requirements during
multiple-effect evaporation. With air oxidation there is a slight evaporation of water, which
acts to concentrate the solids. But the effect can be offset for air by liquor cooling,
particularly for strong black liquor.
The oxidation of kraft black liquor increases the efficiency of heat transfer by reducing
mulLiplc.effcct evaporator scaling. Weak black liquor oxidation with air increases the liquor
solids concentration by I to 2 percent by weight because of the water evaporated, thus
reducing steam and fuel requirements. Freedman (28) observes that weak black liquor
oxidation with pure oxygen results in a warming of the black liquor by 100 C (18° F),
thereby reducing the sensible heat requirements during multiple.effect evaporation. There
are several potentially detrimental effects. Miller (41) observed during a series of pilot scale
studies that oxidation with molecular oxygen caused a cooling of strong black liquor.
Tomlinson and Douglas (22) note that the endothermic heat requirements for reduction of
Na 2 S 2 O 3 and Na 2 SO 4 in the kraft recovery furnace can result in slightly lower heat release
and lower subsequent steam generation for use in digestion and evaporation sections. The
decreased heat release could cause possible increases in supplementary fuel costs.
Robcrson (56, 57) observed that weak black liquor oxidation with air reduced its heating
value by about 2 percent, with a resulting decreased heat release. Lindholm and Stockman
(55) find that weak black liquor oxidation with molecular oxygen reduced the liquor
heating value by 2.0 and 3.6 percent for Na 2 S oxidation efficiencies of 90 and 100 percent,
respectively. The loss in heating value is caused primarily by oxidation of organic matter,
resulting in increased oxygen consumption and decreased heat availability. Cooper (58)
observes that black liquor oxidation with oxygen can decrease the heating value of weak
black liquor by 2 to 4 percent, and strong black liquor by 6 to 8 percent.
9.7 Air Pollution Effects
Weak black liquor oxidation has several beneficial effects on reducing malodorous sulfur gas
emissions from several sources, including the multiple-effect evaporator noncondensabic
gases, tall oil vents, storage tank vents, and the direct contact evaporator of the recovery
furnace. Weak black liquor oxidation also can result in improvement of evaporator
condensate water quality to facilitate process water reuse. Strong black liquor oxidation
results in a reduction in malodorous sulfur gas emissions from direct contact evaporation,
but does not provide any of the benefits for multiple-effect evaporation as does weak black
liquor oxidation. Black liquor oxidation with air creates an additional source of odorous gas
emissions. These emissions arc practically negligible when molecular oxygen systems arc
employed.
9-30
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9.7. I Direct Contact Evaporation
Direct contact evaporation is used at most kraft pulp mills to conccntrate black liquor from
50 percent solids to 60-65 percent solids. Direct contact evaporation may liberate
malodorous sulfur gases on contact of liquor with recovery furnace flue gas because the CO 2
of the flue gas may acidify the lignin sufficiently to generate 1-12 S from Na 2 S. Black liquor
oxidation can minimize reduced sulfur gas emissions from the direct contact evaporator by
oxidizing Na 2 S to stable products in the liquid.
Major variables affecting the emission rate of malodorous sulfur gases are inlet Na 2 S and
NaHS concentrations, liquor pH, and degree of gas-liquid contact. Murray (59) observes that
increasing the black liquor pH above 12.0 causes a substantial decrease in H 2 S emission
from recovery furnaces following direct contact evaporation. Increasing the degree of
gas-liquid contact by increasing the pressure drop at high liquor firing rates can increase
total reduced sulfur (TRS) emissions from the direct-contact evaporator.
The primary variable affecting H 2 S emission from the direct-contact evaporator is the Na 2 S
concentration of the incoming black liquor. Oxidation of the Na 2 S and NaHS to innocuous
products in either weak black liquor upstream or strong black liquor downstream can
substantially reduce malodorous sulfur gas emissions. This reduction is particularly
significant for the multiple-effect evaporator noncondensable gases, tall oil vent gases, and
recovery furnace flue gases, and also evaporator condensates. The substantial reduction in
sulfur gas emissions from direct-contact evaporation is shown in Table 9-9.
Due to their growing relative significance, organic sulfur emissions from direct contact
evaporation will also be important to consider as more stringent emission standards are
adopted. Methyl mercaptan (CH 3 SH) is oxidized in the presence of oxygen to CH 3 SSCH 3 ,
which has a lower odor threshold level. CH 3 SCH 3 and CH 3 SSCH 3 can be removed by the
TABLE 9-9
EFFECT OF BLACK LIQUOR OXIDATION ON SULFUR GAS
EMISSIONS DURING DIRECT CONTACT EVAPORATION (60)
Liquor
Sulfur Compound Unoxidized Oxidized
kg S/t (lb S/ton) kg S/t (lb S/ton)
H 2 S 2.50-15.00 (5.0-30) 0.05-1.00 (0.1-2.0)
C l- I 3 I-IS 0.15-1.00 (0.30-2.0) 0.02-0.10 (0.04-0.2)
CH 3 SCH 3 0.02-0.08 (0.04-0.16) 0.01-0.03 (0.02-0.06)
CU 3 SSCH 3 0.05-0.15 (.1-.3) 0.01-0.08 (0.02-0.16)
9.31
-------
stripping action of the flue gases. The degree of removal depends on the inlet concentrations
in the liquid, the liquid and gas tcmperaLures, and the pressure drop across the direct con-
tact evaporator. Polishing with oxygen at temperatures of 120 to 1400 C (250 to 285° F)
may be necessary to oxidize the organic sulfur constituents in strong black liquor and,
therefore, prevent their release to the atmosphere.
9.7.2 Evaporator Noncondensable Gases
Weak black liquor oxidation has substantially reduced H 2 S and CH 3 SH emissions from
evaporator noncondensable gases, according to Douglass (61), Rcid (62, 63), and Galcano
(44). The sulfur gas emissions arc reduced by increasing black liquor pH and decreasing inlet
Na 2 S and NaHS concentrations. The effect of ‘cak black liquor o,ciclation on malodorous
sulfur gas emissions from multiple-effect evaporator noncondensable gases is presented in
Table 9-10 (60).
TABLE 9-10
EFFECT OF WEAK BLACK LIQUOR OXIDATION ON
MALODOROUS SULFUR GAS EMISSIONS FROM
EVAPORATOR NONCONDENSABLE GASES (60)
Liquor
Sulfur Compound Unoxidized Oxidized
kg S/t (lb S/ton) kg SIt (lb S/ton)
H 2 S 0.05-1.40 (0.10-2.8) 0.00.0.01 (0..02)
CH 3 SH 0.05-0.50 (0.10-1.0) 0.05-0.10 (0.1.0.2)
CH 3 SCH 3 0.01-0.02 (0.02-0.04) 0.01-0.04 (0.02-0.08)
CH 3 SSCI- 1 3 0.00-0.01 (0.0-0.02) 0.02-0.05 (0.04-0.10)
9.7.3 Evaporator Condensate Waters
The oxidation of weak black liquor results in reducing the amounts of malodorous sulfur
gases liberated for subsequent absorption in the evaporator condensate waters. As reported
l)y Kosaya (43) and Chisoni (64), two European mills employ weak black liquor oxidation
primarily to control odor levels in evaporator condensate waters and to gain, as a side effect,
control of water pollution.
Threc additional studies have been made regarding the effect of weak black liquor oxidation
on evaporator condensate water quality. Shah and Stephenson (14) find that the installation
of a weak black liquor oxidation system reduced the BOD of the condensate water by 27
percent and substantially reduced the odor level. The liquid pH is raised from 6.5 to 9.0,
making the evaporator condensate water suitable for process reuse as brown stock washer
9-32
-------
makeup water. Galcano and Amsden (44) note a 38 percent rcduction in BOD and 60
percent reduction in sulfide ion concentration with weak black liquor oxidation using
molecular oxygen at a southern U.S. kraft mill.
Turner and Van Horn (65) find that Cl-I 3 OH contributes 64 percent of the total ROD of
evaporator condensate water from evaporating unoxidized black liquor at a southern kraft
mill. Galeano and Amsden (44) speculate that one reason for the decrease in l3OD of the
condensate water is the partial oxidation of CII 3 OH to formaldehyde (I-Id-tO) during weak
black liquor oxidation, along with the reduction in sulfide ion concentration. Weak black
liquor oxidation may promote reuse of evaporator condensate water by reducing the
treatment required for recycling.
9.7.4 Tall Oil Vcnt Gases
The oxidation of weak black liquor reduces the inlet Na 2 S concentration of the liquid
supplied to the tall oil recovery system; the t 12 S liberated during aeidulation of the tall oil
proportionately declines. Little information is available on sulfur gas emissions from tall oil
processing.
9.7.5 Storage Tank Vents
Black liquor oxidation can reduce the emission of H 2 S and CH 3 SH from storage tank vents.
Galcano and Amsden (44) report reductions of 99 percent for H 2 S, 50 percent for CH 3 SH,
and 78 percent for organic sulfides from the storage tank vents following weak black liquor
oxidation with oxygen.
9.7.6 Ambient Air Quality
Hendrickson and Harding (3) observe that the installation of parallel strong black liquor
oxidation systems substantially reduces the malodorous sulfur gas concentrations in the
ambient air near the mill. The ambient level of odorous gas concentrations, as measured by
“reducible sulfur” concentration (a general indicator of H 2 S levels), was 50 to 75 percent
lower after installation of black liquor oxidation facilities.
9.8 Oxidation Tower Emissions
Black liquor oxidation facilities can reduce sulfur gas emissions from the direct contact and
multipIe effect evaporators, tall oil vents, and storage tanks. The use of weak or strong black
liquor oxidation with air does provide an additional source of reduced sulfur compounds to
the atmosphere. Reduced sulfur emissions from tank vents where black liquor is oxidized
with molecular oxygen arc negligible in comparison to those from air oxidation systems
because the nitrogen.argon dituent is no longer present to cause any stripping.
9-33
-------
Primary factors affecting sulfur gas emissions from black liquor oxidation tower vents when
air is used for oxidation are the inlet concentrations in the black liquor, the temperature of
the black liquor, the height of the black liquor in the tank, and the air flow rate per unit
volume. Sulfur gas emissions tend to increase for higher liquid temperatures and for greater
air flow rates because of the greater volatility of warmer gases and the stripping action of
the air. Sulfur gas emissions also tend to increase with increasing inlet concentrations in the
incoming black liquor. Use of hardwood species or contaminated condensates for brown
stock washer makeup can be the cause of increased inlet concentration.
The primary sulfur gas constituents present in black liquor oxidation tower exhaust arc
organic sulfur compounds, such as CH 3 SCH 3 and CH 3 SSCH 3 . In addition, other volatile
organic nonsulfur constituents can be stripped from the black Liquor, such as terpenes,
alcohols, and hydrocarbons. The emissions from weak and strong black liquor oxidation
systems are summarized in Table 941 (66).
TABLE 9.11
REDUCED SULFUR EMISSIONS FROM BLACK LIQUOR
OXIDATION TOWER VENTS USING AIR (66)
Condition Weak Black Liquor Strong Black Liquor
kg S/daily t (lb S/daily ton) kg S/daily t (lb S/daily ton)
Average 0.06 (0.11) 0.05 (0.10)
Minimum 0.01 (0.02) 0.005 (0.01)
Maximum 0.11 (0.22) 0.09 (0.18)
Large volumes of exhaust gases with high moisture content are generated during black liquor
oxidation. These gases can be incinerated in large boilers that have sufficient combustion
capacity, provided that safety precautions are taken to accommodate the wet gases. Hisey
(67) describes a system at a kraft pulp mill in South Africa where exhaust gases from the
black liquor oxidation tower are added to the primary air inlet of the recovery furnace. This
is the only installation known to date to employ this technique.
9.9 Process Economics
Two recent surveys of capital and operating costs were made for weak and strong black
liquor oxidation systems that use air (4) (56). The weak systems normally have higher
capital costs than the strong ones for equivalent production because of greater liquid volume
9.34
-------
at lower solids concentration. The strong liquor oxidation systems normally have higher
operating costs than the weak ones of equivalent production because a greater amount of
energy is required for oxygen mass transfer into the viscous strong liquor.
9.9.1 Air Oxidation Systems
CaicLilations by Roberson (56, 57) verify the higher capital and operating costs for strong
black liquor oxidation systems than for equivalent weak ones. Roberson’s design features
for a hypothetical processing mill were an inlet Na 2 S concentration of 6.0 gIl in the weak
black liquor or the equivalent in strong liquor concentration, and a Na 2 S oxidation
efficiency of 99 percent across the system. Results of the calculations are shown in Figure
9.11.
Blosser and Cooper (4) present a compilation of capital and operating costs for existing
weak and strong black liquor oxidation systems from data supplied by individual mills. Capi-
tal cost figures are adjusted to a base of December 1968 from reported values to correct for
the effects of inflation and are listed in Table 9-12.
TABLE 9-12
ESTIMATED CAPITAL COSTS FOR BLACK LIQUOR
OXIDATION SYSTEMS
Unit Description
Location Type Installation Cost* Reference
$/daily t ($/daily ton)
Weak Packed tower 440-660 (400.600) 4
Multiple tray 440-880 (400-800) 4
Agitated sparger 330-385 (300.350) 4
Rotating fluid 55-165 (50-150) 28
Strong Unagitated sparger 300-770 (275.700) 4
Agitated sparger** 440-660 (400-600) 4
Plug flow reactor** 6-55 (5-50) 35
*Adjusted basis of December 1968.
**Estimated values.
Annual operating costs for black liquor oxidation systems depend on a number of variables.
The major component expense is for electric power, but equipment maintenance, interest
9-35
-------
A. CAPITAL COST
500
400
0
0
2 300
U )
I-
C i )
8 200
-J
100
0
0
B. OPERATING COST
50,000
40,000
30,000
C,)
0
C,
c 20,000
2
10,000
a-
0
0
0
PRODUCTION - TONS PULP/ DAY
FIGURE 9-11
EFFECT OF PRODUCTION RATE ON CAPITAL & OPERATING COSTS FOR
WEAK & STRONG BLACK LIQUOR OXIDATION WITH OXYGEN (56}
0 500 1000
PRODUCTION - TONS PULP/ DAY
1250
500 1000 1500
2000
9-36
-------
on invested capital, and depreciation must also he included. A summary of direct annual
operating costs for black liquor oxidation systems is presented in•Table 9-13.
TABLE 9-13
APPROXIMATE ANNUAL OPERATING COSTS FOR BLACK
LIQUOR, OXIDATiON SYSTEMS USING AiR
Black Liquor Power Requirement Operating Cost*
kW/daily t hp/daily ton $/daily t $/daily ton
Weak 0.16-0.49 0.2-0.6 11-55 10-50
Strong 0.41-0.82 0.5-1 0 22-220 20-200
Docs not include operating cost credits.
Operating variables for systems include inlet Na 2 S concentration; liquor depth in the tank;
auxiliary facilities, such as foam breakers and agitators; and the possible need for chemical
addition for foam control or pl - l adjustment. An additional factor is whether a system is
single- or multiple-staged. Sheppard (68) reports a 30 percent decrease in annual operating
costs in converting from a single- to a double-stage strong black liquor oxidation system, as
listed in Table 9-14.
TABLE 9-14
EFFECT OF NUMBER OF STAGES ON ANNUAL OPERATING COSTS
FOR STRONG BLACK LIQUOR OXIDATION WITH AIR (68)
Number
of
Stages Flow Power Annual Operating Cost
m 3 Ih (cfm) kW (hp) S/daily t (S/daily ton)
1 76,500 (45,000) 2,720 (3,650) 100.50 (91.15)
2 41,000 (24,000) 1,870 (2,500) 69.25 (62.81)
3 35,800 (21,000) 1,810 (2,430) 67.00 (60.77)
9-37
-------
9.9.2 Molecular Oxygen Systems
The capital costs for black liquor oxidation systems using molecular oxygen are
considerably lower than for air oxidation systems. This is because molecular oxygen
proccssing can be carried out within the piping of an cxisting mill without constructing large
separate tanks. Galeano and Amsden (44) estimate the cost for two stage weak black liquor
oxidation system at $55 to $83 per metric ton of installed daily capacity ($50 to $75 per
ton/day) and the cost for strong black liquor polishing at $83 to $165 per metric ton of
installed daily capacity ($75 to $150 per ton/day). Cooper (58) cstimates capital costs for
black liquor oxidation with oxygen, depending on the materials of construction used, as
listed in Table 9 15.
The primary factors affecting operating costs for strong black liquor oxidation with oxygen
are the inlet Na 2 5 concentration and the oxygen unit purchase cost. Normally, 15 to 25
percent excess oxygen must be added to provide for maximum Na 2 5 oxidation efficiency.
An additional 5 to 10 percent more oxygen than that required for weak black liquor
oxidation alone must also be added for strong black liquor polishing. The effect of oxygen
unit costs and inlet Na 2 S concentration on unit operating costs for weak black liquor
oxidation with oxygen is presented in Figure 9-12 (66). A summary of calculated ranges in
operating cost credits and debits for black liquor oxidation with oxygen is presented in
Table 9-16 (69).
A summary of estimated operating parameters, reduced sulfur emissions, and capital and
operating cost values for weak and strong black liquor oxidation with air and oxygen is
presented in Table 9-17.
9-38
-------
0 5
10 15 20 25 30 35
OXYGEN PRICE — DOLLARS/TON
FIGURE 9-12
OPERATING COSTS FOR WEAK BLACK LIQUOR OXIDATION
WITH OXYGEN (66).
1.50
1.00
0.50
0.00
0
-J
0
z
0
I —
C ,)
0
0
Q
z
uJ
Q.
0
9-39
-------
TABLE 9-15
ESTIMATED CAI’ITAL costs FOR BLACK LIQUOR OXIDATION SYSTEMS
USING MOLECULAR OXYGEN
Reactor Section
Weak Liquor System —
Carbon Steel Stainless Steel Carbon Steel
(All costs in $/claily I. ($Idaily Lou))
Strong Liquor System
Slaiuiless Sled
*Cost data froni Popper, Fl. (ed ). Modern Cost Euiguiccring& l’echnuIucs. New York, McGraw-l-IilI Book Co., 1970. p. 80-178.
4 *Liquid retention times of 30 to 60 seconds.
Liquid retention tunes of 15 to 30 seconds.
t Liquid retention times of 30 to 60 seconds.
Liquid pumping
2-3
(2.3)
2-3
(2-3)
1-2
(1-2)
1-2
(1-2)
Oxygen injection
1-2
(1-2)
1-2
(1.2)
1-3
(1-3)
1-3
(1-3)
Piping section**
Storage tank
Total
2-6
11-22
(2.5)
(10-20)
13-17
28-55
(12-15)
( 25 . 50 )t
3-6
6.11
11-22
(3-5)
(5-10)
(10-20)
6-I l
14-28
22-44
(5-10)
( 13 . 25 )tf
(20-4.0)
16-33
(15-30)
44-77
(40-80)
-------
TABLE 9-16
OPERATING COSTS AND OPERATING CREDITS FOR BLACK LIQUOR
OXIDATION WITH OXYGEN (69)
Cost or Cost/Credit
Credit Item Amount Range
$it ($/ton)
Cost Oxygen:
Weak Black Liquor 10-6OkgO 2 /t 0.11-1.65
(20-120 lb 0 2 /ton) (0.10-1.50)
Strong Black Liquor 0.5-7.5 kgO 2 /t 0.01-0.28
(1-15 lb 0 2 /ton) (0.01-0.25)
Cost Electric Power 0.04-0.12 kW per t/day 0.01-0.03
(0.05.0.15 hp per ton/day) (0.01-0.03)
Cost Operating and Maintenance 0.5-1.0 h/wk 0.01-0.02
( 0.01-0.02 )
Total Costs 0.14-1 .98
(0.13-1.80)
Credit Tall Oil Yield 0-13 kg/t 0.0-1.10
(0-26 lb/ton) (0.0-1.05)
Credit Sodium Sulfate 0-15 kg/t 0.0-0.33
(0.30 [ b/ton) (0.0-0.30)
Credit CaO 0.5-2.5 kg/t 0.01-0.06
(1-5 lb/ton) (0.01-0.05)
Credit Kiln fuel savings 3-10% 0.03-0.11
( 0.03-0.10 )
Total Credits 0.04-1.66
(0.04-1.50)
9-41
-------
TABLE 9-17
TYPICAL RANGES IN OPERATING VARIABLES, REDUCED SULFUR EMiSSIONS
LIQUOR OXIDATION SYSTEMS (69)
Air Oxidation Systcms
%VBLO* SBLO** WBLO &
Item Only Only SBLO ____
Reduced Sulfur Emissions:
ANI) cost FACTOItS VOlt BLACK
Operating Variables:
Oxygen rcqm’L, act/theor.
Power rcqm’t., kW/daily
(hp/daily ton)
Na 2 S to direct contacL evaporator,
gil
Na 2 S oxidation efficiency, %
Molecular Oxygen S stems
WBLO SBLO
Only Only
4-6
0.2-0.8
(0.3-1.0)
0.3-1.5
5-8
04-1.2
(0.5-1.5)
0.02-0.1
Digester
On I y
2-3
0.1-0.2
(0.1 -0.3)
0.05-0.5
7-10
0.6-1.5
(0.7-I .8)
0.01-0.5
I .i-i:i
0.02-0.08
(0.03-0. 10)
0.3-1.5
,lII•O &
SIILO
1.2-1.5
0.04-0.12
(0.05-0.15)
0.01—0.10
97.99 95.99 9999k 90-99
1.4-1.7
0.02-0.04
(0.02-0.05)
0.1-1.0
98-99 96-98 99-100
Evaporator gases, kg/t
Tall oil vent, kg/L
BLO tower vent, kglt
Recovery furnacc,directcontact
evaporator, kg/t
0.05-0.5
0.05-0.1
0.05-0.25
0.1-1.5
0.5-5.0
0.5-0.75
0.05-0.15
0.05-1.0
0.05-0.5
0.05-0.1
0.05-0.3
0.05-0.5
0.025-0.25
0.025-0.1
0
0.025-0.5
0.05-0-25
0.025-0.1
0-0.01
0.05-1.5
0.05-5.0
0.5-0.75
0.-0.005
0.1-1.5
0.05-0. 1
0.025-0.25
0.05-025
Economic Factors:
Capital cost, S/daily t
(S/daily ton)
Annual operating cost, S/daily t
(S/daily ton)
330-880
(300-800)
11-110
(10-100)
550-880
(500-800)
27-220
(25-200)
660-1100
(600-1000)
55-275
(50-250)
11-55
(10-50)
33-550
(30-500)
27-137
(25-125)
11-330
(10-300)
55-220
(50-200)
22-440
(20-4.00)
55-165
(50150)
17-385
(15-350)
* VI3LO = Weak Black Liquor Oxidation.
**SBLO = Strong Black Liquor Oxidation.
-------
9. 10 References
1. Collins, T. ‘1’., The Oxidation of Sulfate Black Liquor and Related Problems. Tappi,
38:172A- 175A, August 1955.
2. Landry, J., Black Liquor Oxidation Practice and Development—A Critical Review.
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3. l-lendrickson, E. R., and Harding, C. J., Black Liquor Oxidation as a Method of
Reducing Air Pollution from Sulfate Pulping. Air Pollution Control Association,
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4. Blosser, R. 0., and Cooper, I-I. B. H., Survey of Black Liquor Oxidation Practices in the
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943
-------
14. Shah, I. S., and Stephenson, W. I)., Jl 1 eak Black Liquor Oxidation: Its Operation and
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9.44
-------
27. Scott, C. XV., IPeak Black Liquor Oxidation to Reduce Air Pollution with Foam
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Paper, 43:125-126, June 1969.
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University of Florida, Gainesville, February 1962.
9.45
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40. Sakbuga, L.. and Bosu, S., Studies on the Fixation of Sulfide Sulfur in Sulfate Black
Liquor. Indian journal of Technology, 6:149-1 52 May 1968.
41. Personal communication with Mr. Roy L. Miller, Si. Regis Paper Company, Pensacola,
Florida, February 1972.
42. Cooper, H. 13. H., and ftossano, A. T., Jr., Black Liquor Oxidation with Molecular
Oxygen in a Plug Flow Reactor. rIapI)t 56:100-103, June 1973.
43. Kosaya, G. S., Black Liquor Oxidation wit/i Oxygen. Buniayhnoya Promyshlennost,
31:15, June 1956.
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Oxygen. Tappi, 53:21 42-2146, November 1970.
45. Owens-Illinois Odor Reduction Oxidation System Made Available. Southern Pulp and
Paper Manufacturer, 36(3):38, March 10, 1973.
46. Fones, ft. E., and Sapp, J. E., Oxidation of Kraft Black Liquor with Pure Oxygen.
Tappi, 43:369-373, April 1960.
47. Kringstad, K. P., McKean, W. J., Libert, J., Kieppe, P. J., and Laishong, C., Odor
Reduction by In-Digester Oxidation of Kraft Black Liquor with Oxygen. Tappi,
55:1528-1533, October 1972.
48. Berry, L. R., Black Liquor Scaling in Multiple Effect Evaporators. Tappi, 49:68A-71A,
April 1966.
49. Von Essen, C. C., Corrosion Problems in Sulfate Pulp Mills. Tappi, 33:]4A-32A, July
1950.
50. Cyr, M. F., and 1-larper, A. M, ! ‘!ultiple Effect Evaporator Project. Pulp and Paper
Magazine of Canada, 61:T247-T249, April 1960.
51. NCASI and Members Host Symposium on Black Liquor Oxidation. NCASI Monthly
Bulletin, 10:] -3, February-March 1972.
52. Staicll, J. A., and Schmitt, M. C., Some Practical Aspects of the Chemistry of Sulfur in
the Kraft Recovery Process. Paper Mill and Wood Pulp News, 61(44):12, October 29,
1938.
53. Sarkancn, K. V, 1-Irutfiord, B. I., Johanson, L. N., and Gardner, H. S., Feature Review
Krafl Odor. Tappi, 53:766-783, May 1970.
9-46
-------
54. Personal communication with Mr. Rick Ritchcy, Southland Paper Mills, Inc., Luikin,
Texas. Octubcr 1972.
55. Lindholm, I., and Stockman, L., 1-leat Evolution daring Black Liquor Oxidation.
Svcnsk Papperstidning, 65(19): 755-759, October 15, 1962.
56. Robcrson, J. E, How Does Recovery Odor Control Affect a Kraft Mill Energy Balance.
Pulp and Paper, 43:151-154, November 1969.
57. Roberson, J. E., The Effect of Odor Control on a Kraft Mill Energy Balance. Air
Pollution Contro [ Association, 20:373.382, June 1970.
58. Cooper, I-I. B. H., Black Liquor Oxidation with Molecular Oxygen in a Plug Flow
Reactor. Ph.D. Dissertation, University of Washington, Department of Civil Engineer-
ing, Seattle, Washington, August 1972.
59. Murray, F. E., and Rayncr, I-I. B., Emissions of hydrogen Sulfide from Kraft Black
Liquor daring Direct Contact Evaporation. Tappi, 48:588-593, October 1965.
60. Hendriclcson, L. R., Roberson, J. E., and Koogler, J. B., Control of Atmospheric
Emissions in the Pood Pulping Industry. Volume I. Final Report, Contract No. CPA
22-69.18, U.S. Department of Health, Edtication, and Welfare, National Air Pollution
Control Administration, Raleigh, North Carolina, March 15, 1970.
61. Douglas, 1. B., Sources of Odor in the Kraft I’rocess, III. Odor Formation in Black
Liquor Multiple Effect Evaporators. Tuppi, 52:1 738-1742, Septeni ber 1969.
62. Reid, H. A., The Odour Problem at Maryvale. APPITA, 3(2):479-500, December 1949.
63. Reid, H. A., Soda Recovery and Losses in Kraft Pulping. APPITA, 4(3):338.360,
December 1950.
64. Chisoni, P., Elimination of Odors in a Sulfate Pulp Mill. Tappi, 37:201-205, May 1954.
65. Turner, B. C., and Van Horn, J. E., Identification of Volatile Compounds in Krart Mill
Evaporator Condensates. (Tappi Southeastern Section Meeting, Atlanta, March 1969.)
66. Blosser, R. 0., Miscellaneous Sources and Trë nds in Kruft Emission Control: Overview.
Tappi, 55:1189-1191, August 1972.
67. Hiscy, W. 0., Abatement of Sulfate Pulp Mill Odor and Effluent Nuisances. Tappi,
34:]-6, January 1957.
9.47
-------
68. Sheppard, M., Design of a High Efficiency Heavy Black Liquor Oxidation System.
(Presented at the Tappi Environmental Conference, San Francisco, May ] 5, 1973.)
69. Cooper, H. B. H., Recent Developments and Future Trends tn Black Liquor Oxidation.
(Presented at the Tappi Environmental Conference, San Francisco, May 15, 1973.)
9-48
-------
CHAPTER 10
RECOVERY BOILER DESIGN AND OPERATION
10.1 General Conditions
10.].] Process Parameters Outside Recovery Boiler
Recovery boilers arc used for the combustion of spent liquor from the following
sodium-base pulping processes:
1. Sodium sulfate (kraft)
2. Sodium bisulfite, and
3. Neutral sulfite semi-chemical (NSSC) in combination with kraft.
The kraft process can be used for making paper grade pulp or, when combined with a
prehydrolysis stage, this process can be used to make dissolving pulp. The pH of the bisulfite
process can range from acidic to alkaline at the finishing stage of the cooking. The NSSC
process also has a certain pH range.
The amount of black liquor and its characteristics depend to a great extent on the type of
pulping used (owing to differences in the cooking liquor and the pulp yield), the species of
wood used for pulping and the site of the growth of the wood.
The part of the wood no longer present in the pulp after cooking is converted to:
1. Organic and inorganic components in the dry solids in the black liquor;
2. Volatile compounds, such as terpenes, CH 3 OH, CO 2 ;
3. Soap and waxes; and
4. Water.
The organic part of the black liquor dry solids can be classified as lignin derivatives and
carbohydrates. Lignin contains more carbon and less oxygen than the carbohydrates. The
hydrogen content is almost the same for both. The compositions of softwood lignin and
hardwood lignin are slightly different. The heat values and oxygen demand for combustion
differ for the different compounds (1) (2).
10-1
-------
Part of the volatile compounds, the soaps and the waxes may be present in black liquor
when it rcachcs the recovery boiler for combustion. Part of these components may also have
been stripped or skimmed from the black liquor.
The chemicals in the cooking liquor are present in the black liquor as the inorganic fraction.
A portion of the inorganic fraction is involved in reactions with the organic material
dissolved in the black liquor during the cooking and another portion of the inorgarlics is
involved in reactions after the cooking. Only the sullur present as sulfide ion (S 2 ), that can
be converted to thiosulfate ion (S 2 0 3 2 ) in the brown stock washing or in an oxidation
plant, is discussed here.
The pulp yield calculated after the cooking on a bone dry pulp basis can vary considerably
for different processes, grades of pulp, and species of wood.
The majority of recovery boilers are used for burning kraft liquor. Some are used for
bisulfite liquor or for cross-recovery for kraft and neutral sulfite pulping liquor. The
characteristics of bisulfite and NSSC liquors can deviate considerably from the normal kraft
liquors.
The dry solids content of the black liquor from the kraft pulping process can vary over a
wide range. Some extreme values are shown in Table 10-1.
TABLE 10-1
BLACK LIQUOR DRY SOLIDS CONTENT
Process Dry Solids Content*
kg/t (lb/ton)
Linerboard pulp 1,100 (2,200)
Paper-grade pulp 1,000 (2,000)
Dissolving pulp 2,250 (4,500)
#Based on air-dried ton of production.
Changes in Lhc reduction ratio (ratio of sulfide sulfur to total sulfur) and the caIisticn ing
efficiency (molar ratio of NaQU to NaOH plus Na 2 CO 3 ) in the white liquor must be
monitored since these changes influence the dry solids composition.
10.2
-------
The heat value of the I)lack liquor dry solids varies with the composition and is higher for
black liquors of high lignin content. rfte amount of air needed for complete combustion is
also higher for liquors of high lignin content. The heat value and the air required for
complete combustion arc relate(l linearly. The deviations, however, from this linear
relationship are, in many cases, of the same magnitude as the amount of excess air that can
be uSed for combustion in a recovery boiler. The deviations from the linear relationship arc
increased by variations in the reduction, causticizing efficiency, and residual alkali in the
black liquor. Figure L0- [ shows the bomb heat values (BUY) of softwood and hardwood
lignins and an average pulp carbohydrate composition as functions of the oxygen required
for complete combustion. Figure 10-2 shows the BHV for some North American black
liquors and the maximum observed deviations from the indicated linear relationship.
The amount of air needed for the combustion and the amount of flue gas can easily be
determined if the elemental analysis of the black liquor dry solids and the water content in
the black liquor are known. The elemental analysis should represent the same stage in the
process as the figure for the dry solids fired. One method of calculation of air and flue gas
flows is shown in Appendix 10.1.
The BHV should be determined for the black liquor without drying it in advance.
Investigations show that considerable changes take place in the heat value of the black
liquor dry solids if it is dried to a powder before it is put in the bomb (3) (4). A black liquor
sample of 60-70 percent solids is placed in the bomb and a small amount of paraffin oil is
added on top to increase the hcut generation, to get complete combustion without residual
carbon in the ash, aiid to prevent the dry solids ash from splattering out of the sample
holder. The air should not be purged from the bomb as the very small amounts of nitric
oxides formed during combustion are useful as a catalyst for oxidizing all sulfur to sulfate.
The heat value determined in this way has a much smaller variance than the heat value
determined according to the method of drying the black liquor to a powder. By relating the
BHV to the black liquor, the error introduced by the method of determining the l5lack
liquor to dry solids is bypassed. Comparisons of heat values before and after black liquor
oxidation should he made using sodium as a reference element because of the change in dry
solids concentration. Comparisons of heat values before and after a direct contact
evaporator can be made by adding another reference element to the black liquor as sod ium
is packed up with ash in the flue gas.
The combustion products in the bomb are in the fully oxidized stage if made with the
modified proceduire. Two cases can be considered depending upon the molar ratio S/Na 2 .
This ratio is normally less than 0.4 for kraft IiqLlor, but could be more than I for
sodium bisulfite liquor. The combustion products are shown in Table 10.2.
The influence of the formation of carbonic acid (H 2 CU 3 ) and the uneven distribution of
1 I2 SO 4 in different condensate drops may be neglected.
10-3
-------
12.
a)
U i
-J
Ui
I
a)
OXYGEN FOR COMPLETE COMBUSTION
Id Ibmoles / lb
FIGURE 10-1
HEAT VALUES VS. OXYGEN DEMAND FOR COMPLETE COMBUSTION OF
LIGNIN AND CARBOHYDRATES
SOFTWOOD LIGNIN.
STRAIGHT LINE THROUGH ORIGIN
9,000.
— 2E%
LI G N IN
5,000
10-4
-------
8000
I —
m
7000
-J
w
I
0
6000
5000
OXYGEN DEMAND FOR COMPLETE COM8USTION
IO Ibmoles/Ib
FIGURE 10-2
HEAT VALUES VS. OXYGEN DEMAND FOR COMPLETE COMBUSTION FOR
SOME NORTH AMERICAN DRY SOLIDS
Thc combustion products from the components of the dry solids in the boiler differ from
those of the bomb. For the different S/Na 2 ratios, the combustion products arc shown in
Table 10-3.
The compounds given in parentheses in Table 10-3 should be kept low in quantity for
acceptable operating conditions. The flue gas might, during incomplete combustion, also
contain H 2 S,CO, H 2 CFI 4 , and CH 3 SH.
Straight Line Through Origin
e
25 30 35 40
10-5
-------
TABLE 10-2
BLACK LIQUOR COMBUSTION PRODUCTS IN CALORiC
BOMB USING MODIFIED TECHNIQUE
S/Na 2
Phase
Less than 1.0
Greater than 1.0
As ash
As gas
As condensate
Na 2 SO 4 . Na 2 CO 3
CO 2
1 -120, H 2 CO 3
Na 2 SO 4
CO 2
H 2 O, H 2 SO 4
TABLE 10-3
BLACK LIQUOR COMBUSTION PRODUCTS
IN RECOVERY FURNACE
SIN a2
Phase
Less than 1.0
Greater than 1.0
As smelt
Asgas
As dust entrained in the gas
As condensate
Na 2 S, Na 2 CO 3
(Na 2 SO 4 , Na 2 2 03, Na 2 S )
CO 2 ,(S0 2 ), H 2 O
Na 2 SO 4 , Na 2 CO 3
nil
Na 2 5, Na 2 CO 3
(Na 2 SO 4 , Na 2 5 2 0 3 , Na 2 S )
C0 2 , SO 2 , H 2 0
Na 2 SO 4 , Na 2 CO 3
nil
Chlorides in thc dry solids will form sodium chloride in the smelt and the dust, as well as
hydrogen chloride (HCI) and free chlorine (Cl 2 ) in the gas.
Corrections have to be made for the actual combustion products in the recovery boiler and
applied to the calculations for the furnace design and steam generation.
The amount of flue gas and the release of heat can vary considerably, as calculated per ton
of pulp. Furthermore, the ratio between flue gas and heat release does not have a constant
value, but has to be considered during design of the recovery boiler to produce the correct
temperature in the furnace and to produce the correct temperature of the superheated
steam generated in the boiler. These facts complicate the design and operation of recovery
boilers. The American method of stating the recovery boiler capacity in pounds of dry solids
per day is less inexact than the old European method of rating in tons of pulp per day.
Turpentine and soap are formed during the cooking, especially rn the pulping of softwood.
The turpentine is released in the vapor form but the soap is dissolved in the liquor. Some of
the soap is skimmed from the black liquor and used for the production of tall oil, which is a
10-6
-------
valuable byproduct. The remaining soap in the black liquor adversely affects the operation
of the evaporation plant. The soap has a high heat value and will cause complications in the
furnace if it is not efficiently rein oved.
10. 1.2 Comparison of American and Scandinavian Liquor Concentration Methods.
the concentration of the black liquor from the brown stock washing depends on the
equipment in the washing department. The black liquor concentration is increased with the
number of theoretical exchange units in the brown stock washing and with higher losses of
pulp, if such can be allowed. The black liquor concentration normaJly ranges between 15
and 18 percent dry solids. It is not possible to burn black liquor at this low concentration. It
must be concentrated to at least 55 percent, and normally 60-65 percent dry solids, before
injection into the recovery boiler. The dry solids concentration refers to the solids prescnt
when the black liquor leaves the evaporation unit and does not include the chemicals from
the ash hoppers or dust collectors. ‘l’he hopper and precipitator ash mixed with the black
liquor increases the black liquor concentration by 3-4 percent.
The Flow of the inorganic and organic compounds within a North American black liquor
recovery boiler is shown in Figure 10-3. This figure assumes the usc of a dry ash conveyor
system for the boiler tube bank, the economizer, and the electrostatic precipitator. The
normal practice in North America has been, however, to use the black liquor in the ash
hoppers and the electrostatic precipitator to convey the ash back to the furnace. Figure 10-3
illustrates the change in the ratio between inorganic and organic contents of the black liquor
within the recovery boiler department. It does not, however, show the changes in the
composition of the black liquor that take place in the direct contact evaporator.
It is possible to use a fluidized bed incinerator for the combustion of black liquor, in which
ease a lower concentration, down to 35 percent, can be injected into the reactor. The
chemicals are, however, not recovered in a suitable form for further processing, and the heat
recovery as steam is much decreased. These facts make the use of fluidized bed apparatus
for the recovery of kraft black liquor uneconomical except for very small plants. Figure
10-4 shows, in principle, the flow of the inorganic and organic contents through such a
system.
The concentration of the black liquor before injection into the recovery boiler furnace can
be effected in several ways. Most older recovery boilers in the United States and Canada use
a multiple-effect evaporation plant to concentrate thc brown stock wash liquor to about 50
percent. Direct contact evaporators are then used for concentrating up to about 65 percent
solids before injection into the boilers. Three different types of direct contact evaporators
have been used. These arc cascade evaporators, cyclone evaporators, and venturi scrubbers.
The flue gases are normally cooled from 400° C (750° F) to an exit temperature of tlO to
150° C (230 to 300° F).
10-7
-------
/
/
::::::: /
jJDIRECT HEATING
ECONO-
MIZER
FIGURE 10-3
TO MILL
1/ L4 1 ? 4/c4 H20
______ t FLUE
DUST 1 GAS
BT uJ
_____ _________ STEAM
FURNACE
BOILER
TUBE
BANK
DIRECT CONTACT
EVAPORATOR
4 4 :; 7 t
/
ELECTROSTATIC
PRECIPITATOR
1
V
SMELT
1ll,
I ,
MAKE UP
MIX TANK
:
:: :::;:;:::::::
.•.::•:::::::•;
—w.---—-
.. . .:: ••.
——
•:.: .: .. .: .: .:.: .: .:.:.: .: .: .: . . .:.: .: >
4—< .
-
BTU
Iu / / 77/ H 2 0
__________ INORGANICS LIQUOR
4< ORGAN ICS
FLOW DIAGRAM FOR BLACK LIQUOR THROUGH RECOVERY BOILER-NORTH AMERICAN SYSTEM
-------
STEAM TO MILL
PRIMARY DUST
COLLECTOR
I1 WASTE HEAT
[ lIT BOILER
4
4-
‘////J2 YJYE/ ////L
SECONDARY
DUST
-COLLECTOR
_______ FLUE
(GAS
BT ,)
________ — b -- - ‘ ‘
C
1/A j’PELLETS - !DUST
______________ ‘ //// __ H20
FIGURE 10-4
BTU BLACK
4 ( INORGANIC 0R
ORGA N IC )
J-/j H20
4 4* co 2
FLOW DIAGRAM FOR BLACK LIQUOR THROUGH FLUIDIZED BED
REACTOR WITH WASTE HEAT RECOVERY BOILER
REACTOR
—p:::::
-------
The use of multiple-effect evaporation alone to achieve 60-65 percent concentration has
been practiced at all ptilp mills in Sweden anti about half of the pulp mills in Finland. The
eoncentralioii of the black liquor at furnace injection has been increased from about 50
percent in 1940 to 60-65 percent today. Figure 10-5 illustrates the flow of the inorganic and
organic contents through such a system.
The two different approaches to the evaporation of the black liquor both have their
advantages. rrhe Scand iiiavian method of concentrat lug the black I iq Lior exel u sivelv I))’
multiple-effect evaporators for injection into the boiler will give greater steam generation.
The American method of employing reeover ’ boilers with direct contact evaporators are less
expensive to install. Thes’ are also self-compensating for capacity at overloading of the
recovery boilers. The exit temperature of flue gas at the direct contact evaporator will
increase at increasing load. This increase means that more heat is available for the final
concentration of the black liquor. The additional heat is normally sufficient to compensate
for the attendant lower dry solids concentration from the multiple-effect evaporator, which
also is a result of overloading.
Overloading a Scandinavian-type recovery boiler normally means that the black liquor from
the evaporation plant will have a lower solids concentration: however, the overall heat
transfer coefficients will increase at lower concentrations, and the influence of the boiling
point rise on heat requirements wills therefore, lie reduced. This reduction will limit the
decrease in concentration to some extent. The steam-generation rates of the Scandinavian
system and of the American system would be approximately the same if the dry solids
concentration after evaporation are the same. Furthermore, the flue gas flow in the furnace
would be higher with the Scandinavian system since the final concentration of the black
liquor is made in the furnace. This fact tends to increase the vertical velocity’ of the flue gas
and to increase carryover of dry particles from the black liquor and their subsequent
combustion in suspension. The emission of H 2 S in the flue gases is lower than with direct
contact evaporators operating with no oxidation of the black liquor, even tinder the adverse
conditions of overloaded boilers. There may’ also be a certain decomposition of the dry
solids in the direct contact evaporators, which will decrease the steam generation rate.
investigations, however, of the decomposition reactions taking place in the direct contact
evaporators are not conclusive. Decreases in heat value of the dry solids by as much as 6
percent have been reported (5).
The lower capital cost for the American system and the possibility’ of increasing the load oii
the recovery’ boilers while maintaining combustion conditions as allowed by the self-
compensating characteristics of the direct contact evaporators were very’ attractive to the
Scandinavian pulp engineers. Investigations were made oii the economics of the American
and Scandinavian recovery boiler designs and operation. rl he feasibility’ of the Scandinavian
design was investigated, and the additional investment was, iii most cases, found attractive
because of higher energy’ efficiencies. The fuel and po ver prices in Scandinavia, however,
10-10
-------
STEAM TO MILL
I _______
H20
2CO2 1
‘ -DUST
— . BTt. )
FIGURE 10-5
FLOW DIAGRAM FOR BLACK LIQUOR THROUG 1 RECOVERY BOILER
SCANDINAVIAN SYSTEM—LOW ODOR SYSTEM
FURNACE
COMPENSATION FOR INCREASED
EVAPORATION
BOILER TUBE ECONO- MIZER ELECTROSTATIC
BANK PRECIPITATOR
.-\-\ \ \ !\- \-\ \-\ *
J;j
\ \\\ flu.
SMELT
FLUE GAS
MAKE-UP
MIX TANK _iir
4— =
__________________________________ H2O -
4 BTU (Bi.. CK
4 — INORGANICS ( LIQUOR
____ ORGAN ICS)
-------
were t.onsiderabl higher than in most parts of the United States and Canada. An updated
comparison of the conditions on the marginal investments is given in Tables I 0-4, 1 0-5. and
10-6.
The flue gas from an American recovers’ boiler with a direct contact evaporator has a higher
water vapoi- content than Lhc flue gas from a Scandinavian recovery boiler with a large
economizer. The higher moisture condition is ‘cry advantageous for the operation of the
electrostatic precipitators traditionally used for collection of dust from flue gas. A higher
voltage can be used with higher water vapor content, and the handling characteristics of the
dust are better at a lower temperature and a higher water content. (The favorable tendency
of cooling has a practical limit at about 115° C (240° F), and the dust absorbs so much
TABLE 10-4
OPERATING CONDITIONS FOR AMERICAN AND SCANDINAVIAN BLACK
LIQUOR CONCENTRATION*
American Scandinavian
Assumed Operating Conditions Practice Practice
Dry solids, kg/hr (lb/hr) 45,000 (100,000) 45,000 (lOO,000)
Fuel heat value, cents/million Btu 85 85
Steam heat value, cents/million Btu 95 95
Cost of electric power, cents 120/kW month 1 20/kW month
+0.6/kWh +0.6/kWh
Heat consumption by back pressure
power generation, Btu/kWh 4,000 4,000
Black liquor concentration at:
Washers, % 1 7 1 7
Evaporators, % 50 63
Direct evaporators, % 63
Furnace 63 63
Heat consumption in s-effect
evaporation, Btu/Lb 235 235
Exit flue gas temperature, °C (°F) 150 (300) 150 (300)
Flue gas CO 2 content afterboiler, % 18 ] 8
Marginal investment costs for:
Back pressure turbine, $/kW 60
Power boiler (steam generating
section alteration included),
$111) steam 6 6
0il and po er prices and invcstment costs based on data available in Sweden, September, 1973
10-12
-------
TABLE 10-5
CAPiTAL COST COMPARISON OI ’ AMERICAN ANI) SCANDINAViAN BLACK
LIQUOR CONCENTRATION PRACTICES
item American Practice Scand in i vian Praci ice
$ $
Direct contact evaporator 200,000
Larger evaporator plant 600,000
Larger recovery boiler 600,000
Larger back pressure turbine 30,000
Larger power boiler 150,000
Total 350,000 1,230,000
Difference in capital cost 880,000
TABLE 10-6
ANNUAL OPERATING COST COMPARISON OF AMERICAN AND
SCANDINAVIAN BLACK LIQUOR CONCENTRATION PRACTICES
Conversion from North American
to Scandinavian Practice
Item Increases Decreases
$ $
Cost of evaporation 73,300
Evaporator maintenance 6,000
Recovery boiler steam value 324,900
Recovery boiler maintenance 8,000
Power generation value 15,800
Turbine maintenance 600
Maintenance of power boiler 3,000
Total 87,900 343,700
Total decrease in annual operating cost 255,800
Gross margin: X 100% = 29%
10-13
-------
hurnidily from the gas that handling the dust becomes extremely difficult.) Higher
migration velocities COLIId, therefore, be observed in the precipitaLors operating after the
(lireet contact evaporator than those afLer a recovery boiler using a large economizer for
cooling flue gas and using the black liquor which was 1ull ’ concentrated in the
in tii ti 1 )ie-eflect evaporators. ftc characteristics of different tYpes of direct contact
evaporators are discussed in Secti()Ii 10.3, Boiler Design, and in more detail in section 10.6,
l)ircct Contact Evaporation.
10.2 Combustion of l3iack Liquor Dry Solids
10.2. I Arrangement of Combustion
The two main objectives in operating a recovery boiler are to recover the chemicals in the
reduced state (that is, the sulfur should be present as sulfide and not sulfate) and to recover
the heat to generate steam for the process. The value of the chemicals has normally been
much higher than the value of the heat in the dry solids. It has, therefore, been stand arcl
l)rac1ic( when fuel has been inexpensive to reduce the cost of the very expensive recovery
boiler installation by accepting a low heat recovery and producing the necessary steam in a
separate power boiler.
The combustion within a recovery boiler has to be separated into two zones because of the
two different objectives of the recovery boiler operation. The first zone has to be
maintalnc(l under reducing conditions less than the stoichiometric amount of air. The
products of this zone arc a discharge of the chemicals in the molten state with the sulfur
present mainly as sulfide and a discharge of the organic matter as a gas having considerable
heat value. The second zone of the combustion starts with the addition of secondary air.
The amount of secondary air corresponds to the amount of additional air theoretically
needed for complete combustion of the gas, pius an excess of about 10 to 20 percent.
laigure 10-6 is a schematic drawing of two recovery boiler furnaces. The left hand figure is
the Babcock & Wilcox (B&W) design. ‘l’hc nozzles for the primary air (item 4 in the figure)
and the secondary air (item 5 in the figure) arc shown. The right.hand figure shows the
primary and secondary air supply, according to another manu facturer, Combustion
Engineering (CE). In the Babcock & Vilcox (B&W) design. some of the secondary air is
introduced at a higher level (item 6 in the figure) and is normally referred to as tertiary air.
The black liquor is sprayed in rather small drops over the cross section of the furnace
through the flue gases. The intention is to dry the black liquor droplets to a concentration
where the heat value of thc char material with the residual moisture is sufficient to keep a
reasonably stable eomhustion going. The liquor spray nozzles can l)e placed as item 3 in
l ’igure 10.6 indicates. The black liquor dry solids are collected in the bottom of the
10-14
-------
COMBUSTION ENGINEERING
Steam
A
5—•
3—
4—p
2-
LEGEND
I. Furnace
2. Smelt Spouts
3. Black Liquor Spray Nozzles
4. Primary Air Supply
5. Secondary Air Supply
6. Tertiary Air Supply
7. Position of Char Bed Burners for Oil or Gas
8. Normal Configuration of Char Bed
8’. Same at Low Primary Air Flow and Pressure
9. Screen Tubes
10. Superheater
II. Boiler Tube Bank
12. Exit to Economizer
II 12
9
9
- II ’*II
lilt’’
- -3
4—4
SECTION A—A
FIGURE 10-6
PRINCIPLE DESIGN OF AIR DISTRIBUTION TO
RECOVERY BOILER FURNACES
BABCOCK & WILCOX
10-15
-------
reeover ’ boiler furnace in a char bed. The air l)iirns oil Lhc dry solids from the to,) of tile
char bed, which normally has the form indicated in the left-hand part of Figure 10-6.
The primary air is supplied around the circumference of the recovery boiler furnace. rrlilb air
supply should be distributed with a fairly constant ratio to the local supply of black liquor
dry solids. The primary air is normally preheated to a temperature of 150° C (300° F).
Some boilers with higher primary air temperatures have been made. The primary air jets will
penetrate a certain distance into the furnace depending on the size of the jets and their
velocity. The oxygen in the air will be consumed in the combustion during the penetration
of the jet into the furnace. The partial pressure of the oxygen in the border zone between
the flue gas atmosphere and the char lied wilt, therefore, vary over the whole extension of
the air jet. The temperature will also var ’. rrlw oxygen partial pressure and the temperature
are the two main parameters that determine the combustion conditions in a limited spot.
The combustion conditions and the equilibrium conditions will, therefore, vary’ over the
en tire cross section of the iii mace. One Scandinavian manufacturer (C otaverken) introduces
some of the primary air at a higher level and with a wider spacing between the nozzles and
at a considerably higher pressure than for the conventional primary air supply. This is called
high primary air anti the arrangement is shown in the left-hand part of Figure 1 0-7. The
conventional design is shown in the right-hand part of the figure. The char bed will form as
indicated in the figure. Other Scandinavian manufacturers are using secondary air for the
same purpose as high primary air (6).
The flue gases formed during primary and secondary air combustion cool by radiation to the
furnace walls to about 870° C (1600° F) before they enter the convection heating surfaces
(i.e., the screen tubes, the superheater, and the tube boiler bank). The flue gas has a
temperature of about 450° C (850° F) after the tube boiler hank. The flue gases are then
further cooled in an eeonomLzcr to about 400° C (750° F) in the American system with the
direct contact evaporator and to about 160° C (320° F) in the Scandinavian system. Two
typical American recovers’ boilers of modern design with economizers instead of direct
contact evaporators are shown in Figures 10-8 and 10-9 for 13&%V and for CE, respectively.
The reason for having a recovery boiler furnace large enough to allow the flue gases to cool
to about 870° C (1600° F) before they enter the convection surfaces is to allow complete
combustion of entrained organic particles before they can reach the cotter tube surfaces.
i’he deposits from organic matter have been shown to be much more difficult to remove
from the heating surfaces than normal tube deposits.
Burners for oil or natural gas to supply heat to the lower part of the furnace arc arranged as
indicated in Figure 10-6, item 7. These burners are used to:
10-16
-------
HIGH SECONDARY AIR
LOW SECONDARY Al R
HIGH PRIMARY AIR
LOW PRIMARY AIR
AIR SUPPLY ACCORDING
NEW SYSTEM — OLD SYSTEM
FIGURE 10-7
0
TO GOTAVERKEN ANGTEKNIK AB DESIGN
1. Bring the recovery boiler up to operating temperature and to supply heat for
drying the injected black liquor drops,
2. Supply additional heat to increase the temperature in the furnace to stabilize the
combustion during disturbances,
3. Smelt down the bed at shutdowns to allow inspection of the lower part of the
walls and the bottom of the furnace, and
4. Supply heat for additional process steam generation when the black-liquor supply
is not sufficient.
The first three uses of the burner are necessary for the operation of the boiler. The fourth
use is only justified when operating the recovery boiler so that enough heat is supplied for
process steam generation. It coincides in this case with the second use mentioned. Steam in
excess of the normal rating should be generated by load carrying oil or gas burners placed at
10-17
-------
MODERN KRAFT RECOVERY UNIT FROM BABCOCK & WILCOX
a higher level in the recovery boiler. This position will also tend to keep the temperature of
the superheated steam at the correct value since the ratio of flue gas flow to steam
generation is considerably lower for oil and gas than for black liquor.
The operation of auxiliary fuels in the recovery boiler presents an explosion hazard.
Detailed instructions about the installation and operation of such buruers have been made
by the Black Liquor Recovery Boiler Advisory Committee (BLRBAC) (7) (8). Instructions
about emergency shutdown procedures are also given by BLRBAC.
Tertiary Air Ports
& Windbox
FIGURE 10-8
10-18
-------
SECTIONAL SIDE ELEVATION —
FIGURE 10-9
MODERN KRAFT RECOVERY UNIT FROM COMBUSTION ENGINEERING
10-19
- 1 --—-—-
-------
10.2.2 Reactions in the Primary Air Zone
A standard for operating the recovery boilcr to obtain the proper reactions from thc process
was developed liv observing the effects of rlisturbaiiccs in the comhustion process and of
changcs in operation parameters.
A typical example illustrating this state of knowledge is as follows: The combustion in the
hearth COL 1Id, at local spots, have an insufficient rate of combustion and a io t’ temperature.
resulting in a condition sometimes called ‘black out.” rfte remedy was to use a compressed
air lance to blow the black char from the wall and reignite the local spot. Heavy corrosion
was often observed at places where black out conditions were common. Air lancing was
believed to be the reason for this corrosion. Compressed air lances were also used at places
with normal combustion without any resulting corrosion. It was then deduced that the
black out conditions themselves were responsible for the corrosion, but the mechanism was
not understood. Many of the problems that arise in attempting to achieve stable combustion
without forming local black out conditions arc caused by the necessity of reaching a
corn promise between two corn peti ng design requirem ents in providing sufficient reaction
surface between fuel and primary air. The furnace cross section shouLd have a relatively
small size with the present arrangement to allow suitably high temperatures in the primary
combustion zone. At the same time the cross section of the furnace has to exceed a
minimum size to avoid high vertical gas velocity. A high vertical velocity would tend to
entrain burning char pieces and black liquor drops. This carryover would result in buildLip of
deposits on the heating surfaces in the boiler following the furnace, and a higher dust load in
the gas.
The very complicated combustion conditions for sodium-based pulping liquors have been
investigated by Bauer and Dorland (9) 1w means of thermodynamic calculations of Lhc
equilibrium conditions at different temperature levels for a certain composition of black
liquor. Baucr and Dorland used the negative logarithm of the partial pressure of oxygen
(rO = — logi 0 02) as the independent variable and negative logarithms of the partial pressures
of the possible compounds (— logs , P) in the flue gas as the dependent variable in determining
equilibrium conditions. This method was introduced by Sillcn and Andersson (10) for
calcium bisulfite arid magnesium bisullitc liquor. The possibility of complete emission of
sodium or sulfur or both to the flue gas under certain furnace conditions and black liquor
compositions exists (see Figure 10-10). The upper portion of Figure 10-10 shows an
equilibrium diagram for condenscd phases and partial pressures of gases for 727° C
(1,340° F), and the lower portion shows the same diagram for 1,127° C (2,06 10 F). The
kinetics of the possible reactions that can occur were not taken into account in preparing
these diagrams. Rosen has extended the black liquor equilibrium studies to include
variations in the black liquor water content and in the pressure (II).
10.20
-------
rO=—Iog P 0 2
1000°K
1400°K
Upper Diagram - NaOH and Na 2 0 cannot exist at
this or higher temperatures.
Lower Diagram - The partial pressure of metallic
sodium (Na and Na 2 ) is becoming
more and more important as the
temperature increases, and will
lead to high fly ash losses.
FIGURE 10-10
EQUILIBRIUM DIAGRAM FOR CONDENSED PHASES AND GASES FOR A
SODIUM BASED BLACK LIQUOR (9)
10-21
-------
The previously cited work and other observations indicate that the release of sodium from
the bLirning char of I)lack liquor is mainly a function of the temperature and the gas
coticlitions in the bordcr zone between the bed and the flue gas. The total sulfur
concentration and the sodium/sulfur ratio control the sulfur release to the flue gas. Figure
10-I I shows the conditions I or a normal kraft liquor. Sodium is uSc(l to represent the base
in the process. Small amounts of potassium are normally present in Lhe wood and also in the
makeup chemicals. Some potassium will, therefore, be present in the black liquor, but it will
react in the same way as sodium. It has a lower boiler point and its compounds have a lower
melting point. The influence of the potassium can normally be neglected exccpL for mass
balances.
‘l’hc amount of sodium which is distributed to the flue gas increases sharply with the
temperature, following a curve similar to a vapor pressure curve. The release of sodium is
also influenced by the primary air velocity, which determines the diffusion conditions. The
amount of soditim in thc flue gas will increase considerably with increasing air velocity, and
this relationship seems to indicate that the release of sodium could be by evaporation. The
rate of evaporation would depend on the diffusion conditions in the border zone between
the flue gas atmosphere and the bed. The sodium to sulfur ratio in the smelt is strongly
dependent on the temperature. The release of sulfur to the flue gas will, therefore, be a
function of the temperature, the Na 2 IS ratio in the black liquor, and, possibly, of Lhc
al)solutc sulfur content in Lhc black liquor.
Large variations in the Na 2 /S ratio iii Lhc smelt have been observed during firing of bisulfite
liquor of constant Na 2 /S ratio. Variations in [ he smelt sulficlity have been observed to a
lesser extent when firing kraft black liquor. The curves for sodium and sulfur release shown
in Figure 10-]] are made for normal contents of sodium and sulfur, that is, 18 percent and
3.5 percent, respectively. At very low temperatures, 700° C (1,3000 F), all the sLilfur in the
black liquor is released to the flue gas. This actually takes place at the shock pyrolysis as
shown in the research for the Billerud-SCA process. Release of sulfur to the flue gas
atmosphere decreases at increasing temperatures, but eventually increases again until total
release occurs above 1,540° C (2,800° F). There should, consequently, be combustion
conditions with a sufficiently high temperature and velocity to evaporate enough sodium to
combine with all the sulfur released to the flue gas, provided that the molar S/Na 2 ratio is
less than I .0 in the dry solids.
The sulfur may be present in the flue gas over the bed as S, H 2 S or SO 2 . A low temperature
favors the presence of S and H 2 S. Elemental sulfur can also be formed by the combustion of
1-12 S with 02 depending on the temperature and molar ratio. Elemental sulfur in the nascent
statc seems to be responsible for much of the observed corrosion on recovery boiler tubes
(12) (13). Observations and analyses made in the last few years show a good correlation
between the factors that influence the temperature in the furnace and the release of 1-12 S
and SO 2 (14).
10-22
-------
high S—content
w
-j
2ct5
LLJZ
LL
wo
(f)
w
S: Normal sulfur content in dry solids (3.5%)
— — - —— s: High sulfur content in dry solids (5 Q%)
FIGURE 10-11
DISTRIBUTION OF SODIUM AND SULFUR IN A BLACK LIQUOR RECOVERY
FURNACE AS A FUNCTION OF TEMPERATURE
H 2 S is probably present in the flue gas in the primary air combustion zone at about 50 to
200 ppm if the temperatures in the border zone between the bed and the flue gas
atmosphere are optimal for the operation. U 2 S concentrations up to 15,000 ppm have been
observed in this region when black out conditions were present. Small amounts of CH 3 SH
have been found in the primary air combustion zone, but no measurable amounts (<1 ppm)
have been found at the entrance to the screen tubes when there was sufficient secondary air
supply.
The sodium probably evaporates from the bed as elemental sodium and reacts with oxygen
to form Na 2 0 within a very short distance from the bed. The Na 2 0 reacts with CO 2 to
form Na 2 CO 3 at the high temperature in the furnace. Part of the sodium compounds
sublime to dust from the vapor phase. The black liquor droplets and agglomerates will pass
through the primary air combustion zone. The volatile compounds will be partially stripped
from the black liquor during the drying. Decomposition of the black liquor dry solids and
pyrolysis and possibly combustion of the smallest fraction of the drops will start before the
drops have reached the bed. The small amounts of mereaptans, organic sulfides, and
IN FLUE GAS
ABSOLUTE TEMERATURE
10-23
-------
hydrocarbons that arc found may be stripped during the drying or the beginning of
pvrolvsis and are not necessarily products of Lhc combustion itself.
Black liquor drops too small to fall downward can he carried upward by the flue gas and
burn in the secondary air combustion zones. The temperaLure in the reaction zone can then
reach much highcr values than at normal combustion, and a higher than normal release of
sulfur can take place.
The smelt is collected at the bottom of the bed and discharged through the smelt spouts to
the dissolving tank. The melting point depends on the smelt composition (see Figure 10-12).
The inciting point increases considerably with contamination of calcium and decreases with
the presence of potassium and chlorides (Figure 10-13). The melting point and the heat
transmission rate determine the thickness of the smelt layer on the tubes.
[ 0.2.3 Secondary Air Combustion
The final combustion starts immediately alter the introduction of secondary air. The total
amount of primary and secondary air must for most boilers be more than [ 10 percent of the
theoretical air (that is, stoichiometric air). It should, on the other hand, be less than 125
percent to avoid the possible formation of sticky dust, which has a great tendency to foul
the heating surfaces in the economizer and/or the collecting plates in the electrostatic
precipitator. The two air limits correspond to 2 and 5 percent excess 02 in the flue gas for
an average kraft liquor.
B)’ definition, the secondary air is the difference between the total air and the primary air.
Some investigations indicate that the primary air flow should be between 60 and 70 percent
of the theoretical air. The remaining air to be used as secondary air would be as a minimum
40 percent, and, as a maximum, 65 percent of the theoretical air. The secondary air should be
supplied in the furnace so that it mixes with the gas coming from the primary air
combustion zone below. The primary air combustion zone gas may have wide variations in
its demand for oxygen to complete the combustion, depending on local variations in the
bed. It is fortunate that it has been possible to achieve practically complete combustion in
the recover)’ boilers without using very large amounts of excess air.
The secondary air is supplied according to two different methods by the main North
American manufacturers of recovery boilers, as shown in Figure 10-6. The (low) secondary
air is, in the B&W boilers, placed only a few feet above the primary air nozzles, and it
controls, in many cases, the height of the bed in the center of the furnace. This air supply
has, therefore, a mixed function, and acts along the walls as secondary air to complete the
combustion of the gas, but in the center of the furnace as primary air to burii off the bed.
The sum of the primary air and (low) secondary air is normally not sufficient to give
complete combustion. The tertiary air in B&W boilers is supplied above the spray nozzles to
10-24
-------
2156°F
LL
0
LU
a:
a:
LU
1500
1400’ 80°F
I30
Na CO % 100 80 60 40 20 0
2 3 0 20 40 60 80 100 %Na 2 S
COMPOSITION OF SMELT % BY WEIGHT
FIGURE 10-12
EQUILIBRIUM DIAGRAM FOR A Na 2 CO 3 -Na 2 S SYSTEM (15)
give a reasonable amount of excess air to complete the combustion. The nozzles for the air
cannot be controlled, and the velocity of the air is, therefore, proportional to the flow and
to the absolute temperature of the air. This means that changes in the load or in the
distribution will influence the velocity of the different air jets and the resulting turbulence
which is presumed necessary for complete combustion.
The design of CE boilers according to Figure 10-6 uses a tangential air supply to produce a
rotary movement of the gas in the furnace. Only four big nozzles are used, one in each wall.
Each nozzle is divided into compartments which can be shut off individually to control the
velocity in the air jet when the flow is changed. The tangential air supply was previously
placed high in the furnace and the rotary movement tended to load one side of the super.
heater and boiler tube bank more than the opposite side. The air nozzles have recently been
moved downwards in the furnace. This change should increase the temperature in the
2100
1600
1564°F
10-25
-------
°F
1500-
1400-
1300-
SMELT
izoo- COMPOSITION
Na 2 S 50%
IIa 2 CO 3 25%
1100 ,‘ Na 2 SO 4 25%
‘/
I I I I I I I I U U
0 50 100
WEIGHT 0/ NaCI ADDED
TO SYNTHEIC SMELT
FIGURE 10-13
EFFECT OF NaCI ADDITION ON THE MELTING
POINT OF A SYNTHETIC PULP
primary air combustion zone slightly, by radiation, and decrease the sideways influence on
the furnace temperature distribution.
The design of a typical Scandinavian recovery boiler, such as shown in Figure 10-7, has the
secondary air supply split between two levels. The low secondary air is placed above the
primary air nozzles but below the spray nozzles, and on all four walls of the furnace. Most
of the secondary is normally supplied at this level. The rest of the secondary air is supplied
in the upper level of boilers as high secondary air or tertiary air. The high secondary air is
supplied on two or four boiler walls.
All air nozzles are adjustable in the modern Scandinavian design. This feature allows control
of the air velocity with total independence of the flow and allows distribution of the air
between different levels. It also allows adjustment of the air distribution in response to the
liquor spray pattern. This has proved very valuable, especially in cases where the boiler load
was low during the start.up period of mill operation.
10.2.4 Formation of Particulate Matter
The flue gas from a black liquor recovery boiler contains large amounts of dust. The dust
load varies between 40 and 75 kg per metric ton of dry solids (80 and 150 lb/ton). The dust
10-26
-------
is formed by the release of sodium from the bed to the flue gas over it. The amount of
sodium released does not seem to depend on the sodium content in the black liquor dry
solids.
The sodium and sodium salts in the black liquor evaporate at a rate dependent upon their
partial vapor pressures and the diffusion conditions. A few feet above the bed, solid salts arc
present as Na 2 CO 3 even at very high partial pressures of SO 2 . The Na 2 CO 3 reacts later with
502 to form Na 2 SO 3 , which is then oxidized to Na 2 SO 4 by the excess oxygen in the flue
gas.
The excess oxygen and the content of SO 2 in the flue gas show a corresponding decrease in
this temperature field according to complete analyses using gas chromatography on samples
taken before the screen tubes, before the superheater, after the superheater, and after the
boiler tube bank (6).
The dust will contain only Na 2 SO 4 after the boiler ii the content of SO 2 is in excess of
what is necessary for the stoichiomctric conversion of Na 2 CO 3 to Na 2 SO 4 , and SO 2 will be
present in the flue gas at the exit of the economizer. In the opposite case (too little SO 2 ),
Na 2 CO 3 will still be present in the dust after the boiler, and the SO 2 content will be zero or
very low at the exit.
The partial pressures of H 2 SO 4 and SO 3 in the flue gas are a function of the reaction
conditions. The formation of SO 3 and H 2 SO 4 can be increased considerably if fuel oil with
a high content of vanadium pcntoxide (V 2 05) is fired in the char bed oil burners or in load
carrying oil burners. The SO 3 and H 2 SO 4 are probably absorbed on the dust parLicics, and the
dust gets sticky. Dust containing Na 2 CO 3 can also form sticky dust through adsorption of
SO 3 .
Rather high concentraLions of H 2 SO 4 and SO 3 were found when firing of sodium sulfite
liquor. The SO 2 content in the flue gas can be as high as 0.5 percent, that is, more than 10
times that of kraft boilers. Corrosion has occurred in the tube bank on such boilers, but not
on kraft boilers. Some instances of corrosion in the last part of the economizers for the low
odor type of recovery boiler might have been caused by SO 3 and H 2 SO 4 .
Dust sampled from 4000 C (750° F) down to 150° C (300° F) has been analyzed, and
differences in the crystalline structure were found. These differences may explain some of
the differences in bulk weight and handling characteristics of dust in the electrostatic
precipitators at different temperatures, even wheii no sticky dust was observed.
\‘cry small amounts of Na 2 S (0.2 percent) have been found in both large and small dust
particles. Large particles of the range 5-15 pm (2.0-5.9 X iO in) could be disintegrated ash
from the combustion of very small drops or agglomerated sublimation products. Smaller
10-27
-------
particles, down to I pm (3.9 X 1 O in), were analvzccl, and the smaller ones of these
probably could only have formed b sublimation. The smaller particles had about the same
Na 2 S content, which was probably a product of reaction between Na 2 CO 3 and l-l 2 S or
clcmcntal sulfur.
The dust containing Na 2 S has no odor at normal atmospheric conditions. 1-12 S, however, is
generated when the dust is exposed to CO 2 and water or water vapor. This fact can explain
the occurrence of kraft odor aL some distancc from the mill when there is almost no odor at
the mill site.
NaCI may be present in the dust lithe black liquor contains chlorides. The 1-ICI and chloride
in the flue gas will probably react with Na 2 SO 4 in the lower temperature range. It is
possible to collect a substantial amount of HCI by scrubbing the flue gas with water. Sodium
chloride condenses at lower temperature than Na 2 SO 4 , and this dust is, therefore, likely to
have a smaller diameter than the Na 2 SO 4 particles in the dust.
Iron can be found in the dust deposits on the tubes at what would be an alarming
concentration if it were caused by corrosion in the boiler. Iron compounds exist, however,
with a relatively high partial pressure at the temperature of the primary and secondary
combustion zone and their evaporation is possible, with condensation Laking place on the
cooler surfaces in the boiler and economizer.
10.3 Different Recovery Boiler Designs
The present design of recovery boilers was developed in North America essentially by B&W
and CE. Manufacturing overseas was by allied companies or licensed boiler manufacturers
and to some extent by independent boiler manufacturers. The main differences between the
two American boiler types are in the air supply, the spraying of the black liquor, the smelt
discharge, and Lhe design of the direct contact evaporator for the final concentration of the
black liquor.
Another approach was often chosen in Scandinavia, with no direct contact evaporators for
recovery boilers. The black liquor was evaporated to the final concentration in multiple.
effect evaporators, and large economizers were used for cooling the flue gas before the
electrostatic precipitators. The reason for this approach was the high cost of fuel and power.
Changes in the American design were necessary to produce odor free operation. The high
ratio of the value of the recovered chemicals to recovered heat, and the very low price of
fuel in North America made it economically feasible to run the recovery boilers with
incomplete comhustion. The deficit in steam generation was made up by oil- or natural
gas-fired boilers at a comparatively low capital cost. The incomplete combustion sometimes
10-28
-------
caused very high emissions of H 2 S. The direct contact c ’apor tors were another source of
H 2 S, which was formed by the reaction:
2Nal-IS + CO 2 + 1120 Na 2 CO 3 + 2112 S
The latter source could, however, be elimindted by oxidation of the black liquor, that is, by
converting the Na 2 S to Na 2 S 2 O 3 by air or molecular oxygen in contact with the black
liquor ahead of the multiple-effect evaporation. This practice, however, would not eliminate
the generation of l-l 2 S in the boiler itself. It would instead have a slight effect in the
opposite dircction by decreasing the heat value of the hlack liquor. This would lower the
temperature in the primary air combustion zone and move the chemical equilibria toward
forming more H 2 S and S, and less SO 2 , from the bed.
The two North American manufacturers proceeded in different ways to achieve odor-free
operation in the cases where complete oxidation was not used. B& V adopted the design
with a large economizer instead of the direct contact evaporator. CE has also tried other
approaches to the problem. The Scandinavian manufacturers have generally kept their
design, with small refinements.
10.3.1 Babcock & Wilcox (B&W) Recovery Boilers
B&W recovery boilers used two different types of direct contact evaporators prior to the
low Odlor era. These were cyclone evaporators and venturi evaporator-scrubbers.
A B&W recovery boiler with a cyclone evaporator is shown in Figure L0-]4. The primary air
is supplied through air nozzles placed around the circumference of the boiler at an almosL
constant height over the slightly inclined bottom of the furnace. The air nozzles are
arranged in groups of normally 4-5 nozzles at the same height. The air flow pressure in each
group can be controlled by a damper. The smelt is discharged at the front wall through
watercooled smelt spouts.
Secondary air is supplied at the distance of about 2 m (6 ft) above the primary air nozzles
(measured at the center) by a smaller numbcr of nozzles. Secondary air is supplied on all
four walls. Spray nozzles for the black liquor arc placed in the front and rear walls above the
secondary air level.
The nozzles can be tilted up and down, and they also have a sideways swinging motion.
Tertiary air is supplied above the spray nozzles for the black liquor. The air to the three
windboxes is supplied by one or two forced draft fans and steam coil air heaters with the
pressure adjusted to the demand for the secondary or tertiary air. The air is throttled to the
primary air windboxes to adjust the pressure to the correct level. Measurements of the air
flows to the different windboxcs are made after the steam coil air heater.
10-29
-------
FIGURE 10-14
BABCOCK & WILCOX RECOVERY
BOILER WITH CYCLONE EVAPORATOR
The furnace is arranged with a rather large arch in the rear wall to distribute the gas in the
superheater. Screen tubes are used ahead of the superheater to adjust the flue gas
temperature to give the correct temperature of the superheated steam. The flue gas then
passes the boiler tube bank and the economizer before exiting to the cyclone evaporator.
The cyclone evaporator is shown in Figure 10-15. The flue gas enters the cyclone
tangentially near the bottom. Black liquor is sprayed across the gas inlet. The liquor drops
are separated from the gas on its helical path to the outlet on the top of the cyclone. Black
liquor is recirculated to nozzles at the top for wetting the walls. This wetting flushes the
black liquor drops and dust, which have separated from the gas, to the sump tank in the
bottom of the separator. Some control of the exit gas temperature or the final
concentration of the black liquor can be exercised by adjusting the black liquor flow to the
spray in the gas inlet. The operation resembles a low pressure drop venturi scrubber. The
liquor is heated in a direct steam heater before it is sprayed into the furnace. The heater
steam condensate dilutes the black liquor to a lower concentration than it had when leaving
the direct contact evaporator.
The dilution of the black liquor is disadvantageous for heat economy, and it represents an
explosion hazard if the black liquor injection into the furnace is interrupted temporarily and
RECtRCULATING
PU APS
EVAPORATOR
RECIRCULATiNG
PUMPS
10-30
-------
FLUE GAS
TO STACK
WALL.WETTING
HOWlS
CYCLONE
EVAPORATOR
CONTROL
FIGURE 10-15
CYCLONE EVAPORATOR
the steam control to the direct heater should malfunction. Reheating of the black liquor
using indirect hcat exchangers has been done in Finland.
B&W introduced a boiler design with a large economizer to meet the demand for low odor
generation (see Figure 10-8). The main difference from the previous design is that the height
of the furnace was increased resulting in a proportional increase in the retention period for
the combustion with tertiary air. Also, a large economizer was supplied for cooling of the
flue gas to about 200° C (400° F) before the exit to the electrostatic precipitator. The
economizer consists of long vertical tubes with a number of baffles arranged to give
substantial crossflow for the flue gas. This arrangement increases the gas velocity and
achieves better heat transmission than parallel flow. The crossflow arrangement seems to
create some pockets with low gas velocity and to create some problems with the ash disposal
with sootblowing, which is done with conventional retractable $ootblowers.
10.3.2 Combustion Engineering (CE) Recovery Boilers
The CE recovery boilers were equipped with cascade evaporators before the requirement for
reduced odor levels. A typical design is shown in Figure 10.16. The primary air was supplied
MECHANICAL
POWER
MX TANK
kACK.UcuOR
INLET SPRAYS
FLUE
GASES
10-31
-------
FIGURE 10-16
COMBUSTION ENGINEERING RECOVERY BOl LER WITH
CASCADE EVAPORATOR
at all four walls with primary air nozzles about 1 m (3 ft) above the bottom of the furnace.
All primary nozzles were placed in a common windbox. The smelt was discharged through
smelt spouts about 0.3 m (1 ft) above the bottom decanting hearth, which is horizontal.
Black liquor spray nozzles were placed about 6 m (20 ft) above the primary air nozzles on
all four walls, with several nozzles per wall. The nozzles could be tilted up and down, but
not swung sideways.
Except for the very first such boilers built, the supply of secondary air has been introduced
at a distance of about 2.4 m (8 ft) above the level of the liquor sprays. The secondary air
nozzles were divided into sections which could be shut off individually to allow for adjust.
ment to the secondary air flow. Oil burners were, in some cases, placed in the secondary air
windboxes for generation of steam in excess of that generated from black liquor combustion.
— uc,.o
10-32
-------
The furnaces were built originally without a nose at the rear wall under the superheater. A
nose was added around 1960 to distribute the flue gases through the superheater. The
superheater was made of panels with a side spacing of 30 cm (12 in) to improve the
operating conditions of the superheater.
The gas passed the screen tubes, the superheater, the boiler tube bank, and the economizer
to one or two cascade evaporators of the double rotor type. A cascade evaporator is shown
iii Figure 10.17. The rotors are constructed of tubes between the end disks and carr\’ black
liquor up through the flue gas pass. The gas velocity and the temperature determine the
diffusion conditions and the evaporation rate to the flue gas. The time of exposure to the
gas is considerably longer on the outer part of the rotors than on the inner part and may
cause overdrying of black liquor at the outside of the rotors. This can be decreased by
increasing the speed of the rotors. The residence time in the cascade is rather long (about
four hours) compared to the cyclone evaporator. The variations in black liquor
concentration caused by changes in the flue gas conditions should be damped by the large
volume in the cascade. The concentration is normally controlled by bypassing the
economizer with more or less flue gas through a damper or by diluting the black liquor with
weak black liquor.
The demand for low emissions was met by CE with the introduction of the air cascade (see
Figure 10-18). The direct contact with the flue gas was eliminated by using the combustion
OPEN VIEW
FIGURE 10-17
OF CASCADE EVAPORATOR
10-33
-------
LEGEND
-4 ROUTING OF AIR
ROUTING OF FLUE GASES
I. LAMINAIRE AIR HEATER
(COOLING FLUE GAS)
2. AIR CASCADE EVAPORATOR
3. FORCED DRAUGHT FAN
4. SECONDARY AIR
5. PRIMARY AIR
FIGURE 10-18
COMBUSTION ENGINEERING RECOVERY BOILER WITH
CASCADE EVAPORATOR ACE SYSTEM
air for the direct contact evaporation. This system is called ACE. The air was heated to
400-425° C (750.800° F) by rotating air heaters of the Ljungstrom type. This system had
the disadvantage that the air to the furnace carried all the evaporated water vapor from the
black liquor and so increased the humidity ratio by about 0.1 kg H 2 0/kg dry air (0.1 lb
11 2 0/lb dry air). The increase in the partial pressure of the primary air increases the
endothermic reaction between CO and 1120 to form CO 2 and hydrogen (H 2 ) at the contact
with the bed. The temperature in the primary air combustion zone would, consequently,
decrease, and tend to increase the emission of sulfur to the gas and increase the ratio of H 2 S
to SO 2 . The lower temperature also would tend to decrease the reduction of the smelt.
A more recent development is shown in Figure 10-19. Recovery boilers of this type use large
rotary air heaters to cool the flue gas to a temperature suitable for the electrostatic
precipitators. The combustion air is heated to 315° C (600° F) to increase the temperature
in the furnace, which increases the release of sodium, reduces the emission of sulfur from
the bed, and decreases the ratio of H 2 S to SO 2 . This system is called LAH. The forced draft
fans are placed after the heaters and are common for both primary and secondary air. This is
disadvantageous in increasing power consumption and in decreasing the accuracy of air
measurement.
10-34
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LEGEND
ROUTING OF AIR
—p ROUTING OF FLUE GASES
I. LAMINAIRE AIR HEATER
2. INDUCED DRAUGHT FAN
3. SECONDARY AIR
4. PRIMARY AIR
FIGURE 10-19
COMBUSTION ENGINEERING RECOVERY BOILER WITH LAMINAIRE AIR
HEATER & COMPLETE MULTIPLE EFFECT EVAPORATION L.A.H. SYSTEM
The final evaporation stage is made in a special type of concentrator following the multiple
effect evaporator to reach 65 percent dry solids before the hopper and filter ash are added.
The influence on the emissions at the high air temperature was very favorable. The theory
predicts a greater evaporation of sodium and a decrease in the H 2 S/SO 2 ratio above the bed.
The results of the operation seem to verify this prediction; however, difficulties were
encountered with the cleaning of the air heater because of narrow spacing between the
regenerator elements. Frequent water washing was required. The air heaters were
dimen sioned for operation with 80 percent boiler load on one air heater, while the other
heater was out for washing. The water washing may prove to be too troublesome for the
operation and may cause some extra corrosion problems similar to those experienced in
Scandinavian water washing of gilled tube economizers used before 1960.
CE now has designed recovery boilers (see Figure 10-9) that include large vertical steel tube
economizers. The economizers are built in two or three passes with an open space between
two consecutive passes. The flue gas passes downi srard parallel to the tubes and upward in
the empty space between the tubes. The pressure drop is relatively low even with high gas
velocities, and the steam used for sootbiowing can be kept at a minimum. Finned tubes are
often used with this design to enlarge the heat transmission surfaces of the tubes. This type
of economizer has previously been used in Scandinavia with very good results.
0
10-35
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10.3.3 Scandinavian Recovery Boilers
The design of recovery boilers iii Sweden and Finland beginning with the first LnJeCtiOil-type
boiler included an economizer followed by an air heater for the cooling of the flue gases to
130° C (260° F). Such cooling was achievable when the heating surfaces were “technically
clean.” This design was chosen because of high fuel prices. The large dust load of the gas
meant that the economizer and air heater often had to be water washed, even with
continuous shot cleaning. Washing periods between 5 and 21 days were standard. The
differences in the cleaning cycles were probably (tue to acid dust caused by high excess air;
however, the influence of acid dust was not known at the time. The water washing damaged
the economizers and they had, on the average, to be rebuilt every l0-L5 years. T he
recirculation air heater, Figure 10-20, was used to decrease the corrosion of the air heaters
in the “cold Corners.” The introduction of the hot precipitator in 195? (placed between the
boiler tube bank and the economizer) solved the problem of keeping the economizers clean
without water washing, but it was more difficult to keep the precipitator in operation for
long periods without shutdowns for cleaning. The long steel tube economizers were chosen
after 1967, when the precipitators had to be designed for high collecting efficiency for
particulate emissions and trouble free operation.
A typical Sc:ndinavian recovery boiler is shown in Figure ]0-21. It is manufactured by
Gotaverken Angteknik AB (G.V.), Sweden, a licensee of B&\V, England. Other manufac-
turers of recovery boilers in Scandinavia include Svenska Maskinwerken AB, Sweden, Oy
Tampella AB, Finland, and A. Ahlstrdm Osakeyhtio, Finland.
A comparison between the North American (Figures 10-14 through ]0-19) aiid the
Scandinavian standards (Figures 10-20 and 10.21) are shown in the preceding illustrations.
The outstanding features in the Scandinavian design are as follows:
I. The primary and secondary air flows are handled by separate fans and air heaters,
A,B.
2. The suction ducts to the Fans are conveyed from the top of the building to give
straight duets, C, D, for an accurate measurement of the gas flow. rrhis also is
advantageous for ventilation since all recovery boilers in Sweden and Finland are
built indoors because of the cold climate.
3. The primary air is split between low primary air, E, in the conventional
windboxes and high primary air, F, in windboxcs at a higher position and at a
higher pressure through a booster fan (Gotaverken).
4. The secondary air is split between low secondary air, H, on all four walls below
the sprayer level and high secondary air, I, (Tertiary air) above the sprayer level.
10-36
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BOILER
Arrangement of “hot” precipi-
tator
Arrangement of “warm” pre—
cipitdtor
FIGURE 10-20
Arrangement of heating surfaces
in economizer—air preheaters. Left,
old system, right, new system
(after 1956)
RECIRCULATION AIR HEATER FOR SCANDINAVIAN RECOVERY BOILER
854
PRIMARY- SECONDARY
AIR
t 85
PRIMARY—SECONDARY
AIR
-------
PRIMARY
AIR
m
—
E
AIR DISTRIBUTION
H IGH
SECONDARY
AIR
LOW
[ SECONDARY
AIR
F
HIGH PRIMARY
FIGURE 10-21
0
TYPICAL GOTAVERKEN 4NGTEKNIK RECOVERY BOILER
10-38
FLOW METERS
M
¶ -iI j--
p.
1 -
-------
5. The cooling of the flue gases after the boiler tube tank is made with a long vertical
sLed tube economizer in several passes.
6. The gas flow is from the top of the tube baffle down to the bottom, which
facilitates the ash transport at sootblowing (same as CE).
7. The flue gas is normally cooled to 1600 C (320° F). This has been found to be the
economic temperature for steam generation, the cost of the economizer and cost
of the precipitator, and the power consumption in the induced draft fan.
8. ‘ [ ‘he ash handling in the hoppers of the recovery boiler tube bank and the
economizer banks, and in the precipitator, is of the dry type. Drag chain
convcyors of a special design (Redler) are used. Rotary valves are used for sealing
of the gas passage at the discharge of the ash. The ash from the hoppers is
conveyed to the mixing tank with vertical dust chutes.
9. To avoid leakage of water vapor up inLo the dust chutes, which causes clogging,
small screw convcyors arc inserted before the mixing tank and equipped with air
jet seals.
10.3.4 Rebuilding of Old Recovery Boilers to Low Odor Design
An existing recovery boiler with a direct contact evaporator can be rebuilt to the low odor
design by installing a feed water economizer, possibly combined with an air heater. This
addition will eliminate the emission of odorous compounds from the direct contact
evaporator, but not from the furnace when the furnace is overloaded.
The air supply system can be revised to achieve better control of the air supply and thus
increase the capacity for complete combustion of black liquor dry solids.
The evaporation plant has to be equipped with a concentrator to increase the conc’entration
of the black liquor to the recovery boiler department to 62-65 percenL. The electrostatic
precipitaLor capacity must be increased to correspond to the reduced emission of particulate
matter allowed by regulations. The higher black liquor concentration will tend to decrease
the flue gas flow, but the higher flue gas temperature increases the velocity of flow. The
result is often an increase of about 10 percent in required precipitator capacity.
Furthermore, the migration velocity of the dust is lower primarily because of the lower
humidity of the flue gas. The dust collecting .efficiency has to be upgraded to meet the
regulations. A new full size precipitator or at least an additional precipitator has to be
included in the rebuilding program. The additional dust collecting capacity may be achieved
with a scrubber if there is a bleaching plant which requires hot water and if the stack plume
caused by the lower exit gas temperature is acceptable.
10-39
-------
These additions can be erected with the old equipment in operation, and thcy can be dueLed
to the existing small economizer and the ID fan in a few days, if local conditions arc
favorable. The costs for lost production caused by downtime will then be limiLcd to
acceptable figures. This approach is probably the least expensive way to get low odor
operation (16).
The usage of an existing recovery boiler for bark- and oil-firing after suitable changes is
possible if a new modern recovery boiler is built in combination with an increase in
production. The losses are Ilten restricted to the special equipment for black liquor firing
and the green liquor system. The life expectancy for a boiler converted to oil and/or bark
firing is, in many cases, greater than if the boiler were kept on-line as a recovery boiler.
10.4 Process Variables
10.4. I Objectives of the Recovery Boiler Process
The main purposes of the recovery boiler operation in the ordinary kraft process are:
I. Recovery of chemicals in the black liquor for dissolving into green liquor with all
sulfur in the reduced state,
2. Generation of steam for the process in the mill,
3. Generation of steam with high pressure and high temperature to allow generation
of power for the mill to the maximal extent,
4. Control of the combination to avoid emission of malodorous gases, SO 2 , and
particulate matter,
5. Allowing the use of the most inexpensive makeup chemicals available, and
6. Reliable operation with minimal capital, operating, and maintenance costs.
Exceptions to these general objectives arc possible under certain circumstances. An example
is cross-recovery between bisulfitc or NSSC and kraft. The flue gases can, in this case, be
scrubbed for recovery of SO 2 for production of the cooking acid for the bisulfite or NSSC
process. The recovery process should then be operated Lo give sufficient SO 2 in the flue gas.
This operation will decrease the sulfidity of the green liquor (and the white liquor from the
recausticizing), and normally this decrease is favorable for a mill with small sulfur losses.
Some of the objectives given above arc contradictory. For example, the generation of steam
with a high pressure and temperature for the generation of power increases low capital and
10-40
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maintenance costs. The capital cost increases with pressure aud temperature and so do the
dangers of corrosion and incrcascil maintcnancc costs. Anothcr contradictory pair are
avoidance of emission of 502 and low opcrational costs. Sulfur dioxide emissions can be
avoided by evaporation of sufficient amounts of sodium to combine with the 503 in the
gas, but this increased evaporation will increase the (lust load in the gas and possibly increase
the demand for sootblownig steam. It will also increase the size of the eLectrostatic
preeipitators needed.
Clearly then, the design of the recovery boiler must be the result of a balance between the
different objectives and conditions, not only in the recovery boiler but also in the different
departments which are influenced by or which influence the operation of the recovery
boiler. Some of the operation and design parameters for normal kraft recovery boiler
application are discussed in the following section.
10.4.2 Firing Rate for Dry Solids
Combustion of a given amount of black liquor solids produces a certain heat input in the
recovery furnace. The combustion requires sufficient air for complete combustion and will
produce an amount of flue gas that depends on the OS composition and amount of excess
air. The temperature in the furnace will depend on four factors, the heat input from the dry
solids, the preheated combustion air, the furnace dimensions and the operating pressure.
The first depends on the production, on the type of wood species used for the mill, on the
pulping yield, on the efficiency of the brown stock washing, and on possible oxidation of
the black liquor. The second factor in the heat input can be varied from ambient air
temperature up to t50° C (3000 F) by heating the air with steam of 0.45 MPa and 1. t MPa
(50 and 150 psig) pressure. The air can be heated further by high pressure steam from the
boiler to 230-260° C (450.500° F), depending on the pressure. This will give a slightly
higher temperature in the furnace for improved operation or to compensate for a decrease in
load. Heating the combustion air from the flue gas with indirect air heaters will increase the
cost of the air duets and would have an adverse effect at load changes (i.e., the air
temperature would decrease when the black liquor heat input decreased). The air
temperature will, on the other hand, increase with increased load; in which case decreasing
its temperature, in some cases even to the extent of using ambient-temperature air, can offer
a favorable solution.
The furnace dimensions must be chosen with some regard to the flue gas flow from
combustion of the dry solids. The amount of dry flue gas from the organic content in the
dry solids can vary considerably with different species and probably also with the rate of
growth and the age of the wood and chip storage time. Changes in the inorganic content and
the causticizing efficiency and reduction will cause only relatively small corrections. The
furnace dimensions, therefore, depend not only on the total amount of dry solids bLit also
10-41
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on thc composition of the black liquor. Tliesc factors should be takcn into consideration if
carryover of sprayed liquor droplets arc to be avoidcd.
Sufficient retention timc in the furnace and excess air arc needed to achieve completc
combustion before the flue gas enters the cooling tubc surfaces. On the other hand, the
oxvgcn content should not be too high to avoid the possible formation of sticky dust.
These factors and the conditions for the chemical reactions mean that it is impossible to
operate a recovery boiler over a wide firing range. Incomplete combustion and further
generation of l-1 2 S will result if the boiler is o crloaded, and a low reduction of the smelt
and a high emission of SO 2 will result at a low load with low temperature in the furnace.
rI he latter case can l)e avoided to some extent liv increasing the heat input by firing an
auxiliary fuel such as oil or natural gas in the char lied burners.
The average firiiig rate at which complete combustion can be achieved without frequent
peaks above I ppm l-1 2 S (with the black Liquor concentrated to at least 62 percent dry
solids) is, according to present observations, as follows;
Dry .solids load per unit of furnace cross section, standard liquor conditions
1. American black liquor, 14,700 kg/rn 2 /day (3,000 lb/sq ft/clay)
2. Scandinavian black liquor, 13,200 kg/rn 2 /day (2,700 lb/sq ft/clay)
This difference is of great importance when the results from combustion in different
countries are compared based upon economy and emissions of odor and particulate matter.
There are, however, a number of cases where higher loads have been obtained with
satisfactory combustion results. These results might depend on a greater attention to the
operation of the boiler than can be expected as an average.
A great number of black licluor analyses have been investigated, both from North America
and Scandinavia. The oxygen demand and the flue gas from the dry solids (the vapor from
the black liquor has been excluded) at theoretical complete combustion without exeess air
varies considerably more for North American liquors than for Scandinavian:
Va nation in Oxygen Demand and Flue Gas for Krafl Black Liquor Dry Solids
I. North American, ±19 percent
2. Scandinavia, ±8 percent
10-42
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Normally there is a raLlier consistent linear relationship between the oxygen demand and the
Bl-l\’ for all fuels and organic compounds. The Bl-l ’ is, however, determined with complete
Combustion to the oxidized slate for all the elements and with the water vapor in the
condensed state. The heat that is released in the recovery boiler is less than the l3 1-l\’.
Corrections have to be made for the reduction of the sulfur compounds, the heat of
evaporation for the water in the black liquor, the water formed by the combustion of the
hydrogen in the dry solids, and the heat of fusion for the smelt. It is, therefore, unlikely
that there should be a consistent relation between the heat input steam and the air
consumption, or between flue gas flow and the amount of black liquor dry solids fired.
The possibilities of reducing emissions by varying the dry solids firing rate in a recovery
boiler are limited. The rated capacity for firing is determined after observation of the heat
value and characteristics of the black licluor. The primary air flow for the rated capacity
must be determined to give the correct temperature in the recovery boiler furnace in order
to achieve the right proportions of sodium and SO 2 in the flue gas. This procedure probably
leaves a rather small range of firing rates at the maximum load within which complete
combustion without release of 1-12 S can be achieved. ‘l’his range is about 5 to 10 pcrce 1 t and
limits the possibility of variations of the operation on the tipper side. rlemporLir l load
changes of greater magnitude must be covered by tank storage for black liquor and green
liquor.
The range of load changes is also limited on the lower side. The relative amount of primary
air and/or the temperature of the primary air has to he increased at a lower firing rate of dry
solids to avoid emission of SO 2 . Such corrective measures might allow operation down to
80% of the normal firing rate. The temperature in the primary air combustion zone will be
too low at still lower loads without additional supply of heat by firing of oil or natural gas
in the char bed burners. I-ugh emissions of H 2 S (and elemental sulfur) can also take place in
the primary air zone at low temperatures. Results from test runs indicate that if the H 2 S
release in the primary air zone is considerably above 200 ppm, even locally, it will be ‘cry
difficult to achieve an H 2 S concentration below I ppm in the exit from the boiler. It should
be possible to get down to 65% of the normal dry solids firing rate with accepLablc
combustion and emissions with a substantial firing of additional fuel like oil or gas in the
char bed oil burners to increase the temperature in the primary air zone.
Operation at reduced load of black liquor dry solids with additional fuel requires an even
distribution of the additional heat input. Concentration of the heat supply from one or two
burners can distort the configuration of the bed and cause carryover of organic material and
combustion in suspension. Means to control the air to fuel ratio for the additional fuel
should be provided.
The possibilities of operating recovery boilers at low load have not been investigated
thoroughly. The high investment cost for the recovery boiler makes low load conditions
10-43
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rather unusual. The minimum load which can be recommcndcd without causing unstable
conditions in the primary air combustion zone, and without additional heat input from oil
or natural gas, would be 80-100% of the maximum load with respect to SO 2 emissions and
65-100% with respect to H 2 S emissions.
10.4.3 Black Liquor Characteristics
The most important characteristics of the black liquor in the design and operation of
recovery boilers are:
1. Heat value and oxygen demand for complete combustion,
2. Viscosity and surface tension,
3. Organic/inorganic ratio and S/Na 2 ratio, and
4. Boiling point rise.
The heat value and oxygen or air demand for complete combustion arc of utmost
importance for both the proper design and proper operation of a recovery boiler. Attempts
to achieve a usable formula for the heat value as a function of the elemental analysis have
met with only limited success. The best solution seems to be to conduct laboratory cooks to
the desired degree of delignification and bomb tests on the test liquor.
The variability in the black liquor composition with different wood species is related to the
relative proportion of lignin and carbohydrates, and the type of carbohydrates, present in
the different speci s (17, 18). Appendix 10-2 shows the composition and the BHV for
softwood lignin, hardwood lignin, and carbohydrates, and the oxygen demand required for
complete combustion. The composition and heat values have then been approximated for
dry solids with a normal content of chemicals.
The. amount of heat that will be released in a reducing atmosphere, that is, the “efficient
heat value in reducing atmosphere,” and the “resulting heat value of black liquor” (the
total released heat used for steam generation) assuming no losses from the flue gas, have also
been calculated. The values are shown in Figure 10-22 as a function of the oxygen demand.
Three lines through the origin are shown. The deviation from a linear relation is shown as
a percent. The BHV show the smallest deviations; however, these deviations are of the same
magnitude as the allowable variations in the excess oxygen. The two lower lines in Figure
10-22, which show the heat released in a reducing atmosphere and the Lotal heat available
for steam generation, show a very good linear relationship, but the carbohydrates have great
deviations from the lines through the origin.
10-44
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Bomb heat value
Efficient heat value in reducing
atmosphere
Resulting heat value of black liquor
Deviation from linear relation
between heat value and ox gen
demand and vice versa in percent
of foctinI value
FI(iUI ft 1U-12
HEAT VALUE VS. OXYGEN DEMAND AT COMPLETE COMBUSTION
A softwood hgr n
B= hardwood ligr n
C carbohydrates
A 1
B 1
ci
Il JOO
lOpOO
oOo.
8,000
Ditto, adjusted for normal
content of unot anics
B
+ 1.4
C.
.0
I—
I i i
D
-j
I-
L i i
I
App. 102
B 1
0
A
+-
OXYGEN DEMAND, lbmoles/ lb DRY SOLIDS
10-45
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\‘ariations in the lignin content in the wood, the amount of extractives and differences in
the cooking process, variations in climatic conditions and the age of the trees before
harvesting, and other variables might explain the great variations in the composition of the
black liquor dry solids from different mills. Consideration of 1)0th the variations in the dry
solids content and the Bl-l V when comparing various firing rates seems to be more accurate
Lhan considering dry solids content alone.
The viscosity of black liquor as a function of the temperature is shown in Figure [ 0-23 (19).
Very great variations were observed between different mills using the same species of wood.
The ‘iscosit ’ is of great importance for the spray pattern. Liquor from eucalyptus pulping
has, as an example, such high viscosity that the black liquor concentration is often limited
Lo 60 percent to achieve a reasonably good, stable distribution of spray.
The surface tension seems to have little influence on the spray pattern. Investigations made
wiLh additives to decrease the surface tension in order to improve the spraying showed
without significant changes.
The inorganic/organic ratio affects the heat value per pound of dry solids and complicates
the operators efforts to keep the load of the recovery boiler constant. The variations in the
inorganic-organic ratio cause variations in the smelt flow and green liquor production. Such
variations may decrease the reduction, as a high inorganic content often is a result of a
tern porarily low eausticizing efficiency and reduction.
The molar ratio S/Na 2 , along with the concentration of sulfur and the temperature,
probably determines the release of sulfur from the bed. The ratio should, therefore, be kept
as nearly constant as possible. In some plants, both the supply’ of makeup salt cake and also
the spent acids from the dO 2 generation and the tall oil plant are added continuously at a
constant rate. The ash transport from the ash hoppers in the boiler and the economizer can
be equalized to some extent by adjusting the sequence of the sootbiowers to give a more
steady supply of ash from different parts. The scraper eonveyors used in Scandinavia, as
compared to the common black liquor flushing system used in North America, tend to give
a steady supply of ash.
The boiling point rise is a function of the inorganic/organic raLio and the concentration. ‘l’hc
black liquor is normally sprayed into the furnace with a temperature near the boiling point.
Changes in the spray pattern will follow a decrease in the boiling point rise if the
temperature is kept at a constant value. An increase in the fraction of fine drops, which will
cause carryover, will be the result of spraying at a temperature above the boiling poimit.
The boiling point rise is sometimes used for determination of the black liquor
concentration. This method gives inaccurate results because of the variations in the
inorganic/organic ratio.
10-46
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I—
C l )
0
>-
I-
Cl)
0
0
Cl )
>
0
w
TEMPERATURE, OC
*Where OS is dissolved solids
FIGURE 10-23
VISCOSITY OF BLACK LIQUORS (19)
10.4.4 Black Liquor Oxidation
Black liquor oxidation was discussed in Chapter 9. Oxidation of thc black liquor with air
converts the sulfide sulfur to thiosulfide. The importance of this oxidation to the recovery
boiler process is that remaining sulfides react with CO 2 in the flue gas to form H 2 S on
contact. H 2 S is also produced from sulfides on contact in a wet precipitator and a wet ash
conveying system. This latter release, however, is insignificant compared to that from direct
contact evaporation. The generation of H 2 S on direct contact with the flue gas is totally
eliminated with complete oxidation; however, oxidation tends to increase slightly H 2 S
generation in the furnace.
The decrease in heat value of the dry solids caused by black liquor oxidation leads to a
decrease in the temperature in the furnace. The lower temperature decreases the ratio of
H 2 S and sulfur to SO 2 , but increases the total sulfur release.
30 40 50 60 708090 110 130
10-47
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Investigations of the influence of oxidation are difficult because not only the sulfide sulfur
but also the organic matter is oxidized. Some volatile compounds are strippcd from the
liquor, and the addition of oxygen increases the weight of the dry solids. Dry solids with 4
percent sLilfur content increase in weight at complete oxidation by 2.5 percent.
Water is evaporated from the black liquor on contact with the air during oxidation. The heat
consumption for this evaporation is quite high, about 3270 kJ/kg, (1400 BTIJ/lb) as
compared to 486 kJ/kg (210 BTLJ/lb) in a conventional 6-effect evaporator. Furthermore,
the black liquor is cooled in the direction of the wet bulb temperature, and this heat has to
be replaced in the evaporating plant. The loss in heating valve causcd by the oxidation of
Na 2 S to Na 2 S 2 O 3 must be replaced in the furnace.
The loss in heat value per unit mass of original dry solids almost doubles if the oxidation
efficicncy is increased from 90 to near 100 percent (20). The heat losses in steam generation
can he calculated as 1] 6 kJ per kg of dry solids (50 BTU/lb) per each percent sulfur for 90
percent oxidation and 209 kJ per kg of dry solids (90 BTU/lh) per each percent sulfur for
100 percent oxidation. Allowances arc made for the evaporation in the oxidation towers as
compared to evaporation in a 5-effect evaporator. This heat loss might gain in importance
with the increased sulfidities resulting from recovering the malodorous gascs and with
increased fuel prices.
The oxidation decreases the air consumption by 1-2 percent but the heat input has a greater
influence on the furnace temperature. The primary air temperature must be increased to
compensate for the lower heat value and to keep the reaction condition constant in the
primary air zone with use of fully oxidized liquor.
10.4.5 Air Distribution and Air Temperatures
The air distribution, the air temperatures, and the spraying pattern are the main
independent variables by which the operation .of the recovery boiler can be changed. The
specific load, expressed as dry solids per unit cross section of the furnace, depends on the
dimensions of the recovery boiler and the production of the pulp mill. The heat value of the
dry solids of the black liquor depends on the type of pulp and the species of wood used for
production.
The total amount of air which can be used for the combustion under the present technology
has both lower and upper limits. The lower limit is set by the condition that at least 10
percent excess air has to be used to avoid unburned matter in the flue gas and formation of
H 2 S. The upper limit of about 20-25 percent excess air is necessary to avoid the possible
formation of SO 3 in the flue gas, and the resultant sticky dust caused by SO 3 adsorption on
the dust. A certain amount of the total air must be used for the primary air combustion
zone to give the proper oxygen concentration in this zone and to give the correct
10-48
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temperature. With the present recovery boiler designs, the correct’amount for primary air is
approximately 60 to 70 percent of the theoretical amount of air for complete combustion
without excess air. The reduction of the sulfur compounds in the smelt is probably
enhanced by a low oxygen concentration and a high temperature, within certain limits. An
increase in the amount of primary air increases both the oxygen concentration and the
temperature. The two parameters of oxygen and tcmpcrature are evidently coupled
together, and some combination of them must exist that is the most favorable.
The air is currently heated either by steam from the backpressure turbine at pressures of
1.13 MPa (150 psig) and 0.45 MPa (50 psig) or with feedwater recycled from the top of the
economizer through air heaters and fed back to the inlet of the economizer (Figure 10-20).
The most economic choice depends on the prices of fuel and power at the specific site. The
steam air heaters give an increased flow of high pressure steam through the turbine and
increase the back pressure power generation. The recirculation air heater permits a lower
capital cost for the economizer for the same flue gas exit temperature. This latter system is a
very clean and easily operated arrangement, but limits the air temperature to about 150° C
(300° F). Possible ways to increase the air temperature are either to return to using the
indirect type of air heater with cooling of the flue gas by air or to use a rotary Ljungstrom
type air heater. The Ljungstrom type air heater requires rather extensive dueting of the hot
air from the heaters to the furnace. Direct heating by firing auxiliary fuel after the steam
coil or reeireulation air heaters is possible. The water vapor from the additional fuel will
change the equilibrium conditions for the cndothcrmie water gas reaction, which is only
partly compensated by the increase in the partial pressure of CO 2 . This method works the
same way as firing oil in the furnace with the char bed burners, which is favorable at low
load conditions. The primary air velocity is increased, tending to increase the evaporation of
sodium.
Advantages in raising the primary air temperature are:
I. Velocity for the same flow of air increases and therefore improves the sodium
evaporation from the bed,
2. Release of sulfur from the bed decreases,
3. Distribution of the released sulfur changes in the direction of more SO 2 and less
H 2 S, and
4. Decreasing the primary air flow is possible, and consequently, more air is available
for the secondary air supply.
Adjusting the velocity and distribution of the primary air flow within the furnace is
probably valuable so that the flow corresponds to the black liquor distribution. B&W
10-49
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designed a system with an adjustable nozzle, which was used around 1960 in the United
States. but it was abandoned. This system has, however, been used in Scandinavia (see
Figure 10-24) as standard equipment.
The difference between the total air flow .and the primary air flow is used as a computation
of secondary air to complete the combustion. Assuming an average figure 01115 percent of
the theoretical air as the total combustion air, and 70 percent used as primary air, only 45
percent is available to complete the combustion. This small amount has to cover the whole
cross section of the furnace and effects good mixing of the combustible gas with the air. The
suppliers of recovery boilers all use different ways of achieving complete combustion. They
all seem to work provided that a reasonably low amount of 11 2 S is present in the gas
mixture coming from the primary air combustion zone.
Jones, Brink, and Thomas suggest that another method of controlling the combustion
conditions is to supply oxygen to the air (21, 22). This tends to reduce the volume of the
eombustion chamber and to allow the temperature to be increased to the desired level. But
FIGURE 10-24
ADJUSTABLE AIR PORTS
PRIMARY AIR
PORT REGISTER
PRESENT ADJUSTABLE
DESIGN
10-50
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there arc some economic drawbacks. The amount of flue gas per unit of steam generated is
less. The superheating of the steam, which is of extreme importance for the back pressure
power generation, is considerably more expensive and more complicated than with the
present type of superheaters, because the flue gas temperature at the entrance to the
superheater must be kept low enough to avoid very difficult slagging problems.
The greatest problem with oxygen addition is probably that of achieving a sufficient mixing
of the gas from the primary air combustion zone with secondary air.
10.5 Diverse Obnoxious Compounds
CH 3 SI-l, CH 3 SCH 3 , CH 3 SSCI-L 3 , carbonyl sulfide (COS), and earbonyl hydrogen sulfide
(COSH) were measurerl in addition to H 2 S, SO 2 , and S in the lower part of the recovery
boiler in the gas phase. They were not found to any measurable extent, however, in the
upper part of the furnace in the neighborhood of the screen tubes where the boilers were
operated with 2 percent excess oxygen. Some tests indicate that I I2 S is present in rather
high concentrations when CO and H 2 arc also present in high concentrations. A very good
correlation seems to exist between H 2 S and H 2 content. Both H 2 S and SO 2 emissions can
be controlled to a satisfactory degree by applying present technology to the design and
operation of recovery boilers as previously discussed (23).
The obnoxious compounds, according to available test results, present no problem in a
reasonably well operated recovery boiler. CH 3 SI-I, Cl-I 3 SCH 3 , and CH 3 SSCH 3 can, however,
be stripped from the black liquor in the direct contact evaporators.
The NO emission from recovery boilers are probably of minor importance because of the
low combustion temperatures that are reached even locally in comparison to the
temperatures which arc reached in the flames of oil and pulverized coal burners. The
relatively high NOx content which has been found after combustion of ammonium bisulfite
(NH 4 HSO 3 ) liquor, is probably caused by oxidation of the monatomic nitrogen produced
by the cracking and combustion of ammonia.
10.6 Direct Contact Evaporation
Direct contact evaporation of black liquor is performed at most kraft pulp mills in the
United States for concentrating the liquid from 40.50 percent to 60-70 percent solids (ash
from precipitators and ash hoppers not included) to facilitate combustion in the recovery
furnace. The normally used direct contact evaporators are the cascade, cyclone, and venturi
evaporators. The direct contact evaporator can act as an air pollution source, but also in
some aspects as an air pollution control device. A development by CE was the ACE system
in which the black liquor is concentrated by the combustion air in a direct contact
evaporator. The air is preheated by Ljungstrom air heaters to 400-425° C (750.800° F). The
10-51
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water vapor and any compounds released from the black liquor are carried with the
combustion air to the furnace. This arrangemcnt is called an indirect contact evaporator.
The relative suitability of direct contact evaporation, as opposed to multiple-effect
evaporation of black liquor, is based on operation and capital costs, flexibility with respect
to capacity, and environmental aspects.
1 0.6.1 Specific System Characteristics
Direct contact evaporation can be performed in one of three different tYpes of systems. One
of the systems, the high pressure drop venturi scrubber with strong black liquor as the liquid
medium, has the dual function of water evaporation and particulate removal. rl he venturi
type evaporators are gradually being complemented with additional dust collectors or
replaced because of their inability to meet particuh tc air pollution emission standards. The
different types of direct contact evaporation systems employed are shown in Figures 10-14
through 10.17. The operating characteristics of recovery boiler flue gases, as indicated by
moisture content, gas temperatures, and particulate loadings for the different types of direct
and indirect contact systems for concentrating strong black liquor, are presented in Table
10-7.
TABLE 10-7
RECOVERY FURNACE EXHAUST GAS PROPERTIES FOR DIRECT AND
INDIRECT CONTACT EVAPORATION SYSTEMS
Direct Contact Evaporators Indirect
Property Cascade Cyclone Venturi ACE
Pressure drop, in. w.g. 2.4 2-4 15-30 2-4
Flue gas temperature at:
economizer exit, °C 315-370 315-370 3 [ 5-370 3 15-370
(°F) (600.700) (600-700) (600-700) (600-700)
evaporator exit, °C 150-163 132-163 88-1 10
(° F) (300.325) (2 70-325) (190-230)
Flue gas moisture content at:
economizer exit, g/m 3 5-9 7-1 I 7-Il 7-li
(grief) (24) (3-5) (3-5) (3-5)
evaporator exit, g/m 3 2-7 5-9 0.9-LO —
(grief) (1-3) (2-4) (0.4.0.8) —
precipitator exit, g/m 3 0.1-1.1 0.1 -1.1 0.07-0.2
(gr/ef) (0.05-0.5) (0.05-0.5) (0.03-0.10)
10-52
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The cascade- and cyclone-type direct contact evaporators arc low pressure drop gas-liquor
contact devices that are uscd for concentration of strong black liquor by evaporation.
Electrostatic prccipitators, located downstream of the direct contact evaporator, provide
particulate emission control and chemical recovery with thesc low lrcssurc drop systems.
Cascade evaporators employ a rotating cylindrical drum with attached tubular wheels
perpendicular to the direction of the gas stream for gas-liquid contact and are normally used
with recovery boilers manufactured by CE. Cyclone evaporators arc basically low pressure
drop cyclonic serubbers for gas-liqLiid contact and are normally used with recovery boilers
manufactured by B&W.
The direct contact evaporator serves several functions besides the concentration of black
liquor to 60-70 percent solids. These functions include:
I. Reducing the inlet gas temperature to the electrostatic precipitator, where the
lower gas temperature results in a reduced volume flow rate, which allows a
smaller precipitator to be constructed at a lower capital cost;
2. Reducing the inlet particulate loading to the electrostatic precipitator by 20-40
percent by weight, primarily by scrubbing of large particles emitted from the
furnace, as reported by Hisey (24);
3. Absorbing about 75 percent of the SO 2 emitted from the recovery boiler (25,
26), and nearly all of the sulfur trioxide (27); and
4. Absorbing H 2 S emitted from the recovery boiler (28, 29, 30) tinder conditions of
high black liquor pH and low sodium sulfide concentration in the strong black
liquor.
10.6.2 Air Pollution Control
Direct contact evaporation has the potential for liberating substantial amounts of
malodorous sulfur gases from the black liquor or for absorbing sulfurous gases generated
from the recovery furnace. Considerable amounts of H 2 S and lesser amounts of CH 3 SH can
be released from the black liquor by acidifying Na 2 S and sodium mcrcaptidc (CH 3 SNa) by
the action of acidic flue gas constituents, such as C0 2 , SO 2 , and SO 3 . Organic sulfur gases,
such as Cl-I 3 SCH 3 and CH 3 SSCI-l 3 , and malodorous organic nonsulfur compounds can be
evolved from the heating of the black liquor by contact with recovery boiler flue gas (31).
Major variables affecting the potential for release of malodorous sulfur compounds from
black liquor during direct contact evaporation include inlet liquid composition, liquid pH
and alkalinity levels, inlet liquor and flue gas temperatures, and the degree of gas-liquid
contact. Recent studies (28, 32) indicate a substantial increase in reduced sulfur emission
10-53
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levels to the stack, if sodium sulfide concentration is increased in the strong black liquor
entering the direct contact evaporator. Therefore, a high black liquor oxidation efficiency
must be obtained to reduce the Na 2 S to between 0.01 and 0.1 gIl to minimize malodorous
sulfur gas generation during direct contact evaporation. See Table 10-8 (33).
TABLE 10-8
EFFECT OF BLACK LIQUOR OXIDATION ON SULFUR GAS EMISSIONS
DURING DIRECT CONTACT EVAPORATION (33)
Sulfur Gas Unoxidized Liquor Oxidized Liquor
kg/ i . lb/ton kg/ i . lb/ton
U 2 S 2.5-15 5.0-30.0 0.05-1.0 0.1-2.0
CH 3 SH 0.15-1 .0 0.3-2.0 0.025-0.10 0.05-0.20
(CU 3 ) 2 S 0.025-0.075 0.05-0.15 0.005-0.025 0.01-0.05
(CH 3 ) 2 2 0.05-0.15 0.10-0.30 0.005-0.075 0.01-0.15
These findings have been verified in subsequent studies conducted by Martin (34).
Murray and Rayncr (29) have shown that the liquid pH of the incoming black liquor can
have a considerable impact on U 2 S emissions during direct contact evaporation. Increasing
the liquid pH reduces the rate of l l2 S formation at any given level of sodium sulfide and
results in a reduction in the amount of 1-125 generated in the direct contact evaporator. See
Table 10-9 (33).
Two additional variables arc liquid alkalinity and gas-liquid contact. The presence of
substantial proportions of carbonate ion in the black liquor at high pH of 12.0 or more gives
the liquid a large potential buffering capacity against pE-l reductions caused by contact with
CO 2 from the recovery furnace flue gas. Increasing the degree of gas-liquid contact may
result in either a substantial increase or decrease in malodorous gas emissions, depending on
the characteristics of the black liquor and the flue gas temperatures.
Under certain conditions, the direct contact evaporator can act as an air pollution control
device to absorb 1 -12 S from the recovery boiler combustion zone flue gases (28, 30, 32). A
particular advantage in direct contact evaporation, whcrc high degree black liquor oxidation
is practical, is absorption of H 2 S from the combustion zone during periods of recovery
boiler upset.
10-54
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TABLE 10-9
EFFECT OF BLACK LIQUOR pH ON H 2 S EMISSIONS
DURING DIRECT CONTACT EVAPORATION (4)
1 - I 2 S Concentration 4
Na 2 S Inlet OuLIct Change
gil
12.6 14.2 11 35 +24
12.3 18.3 27 L22 +95
12.1 15.3 24 180 +156
Com 1 nited in ppm by voluinc at 00 C and 760 mm Hg (32° F and I 0 atm).
10.6.3 Complete Multiptc.Effcct Evaporation and indirect ContacL Evaporation
Comparison
Indirect contact evaporation is used to concentrate black liquor from 50 to 60-65 percent
solids to eliminate thc possibility of odorous gas release during direct contact evaporation.
Systems employing complete multiple-effect evaporation and indirect con tacL evaporation
have been installed at new kraft mills in the United States and Canada (35, 36).
The multiple effect evaporation to the dry solids concentration used for injection into the
boiler eliminates a potentially large source of malodorous sulfur gas emissions resulting from
direct contact evaporation. The technique has proved successful in extensive experience in
Scandinavia (37). Multiple effect evaporation to virtually eliminates the recovery furnace as
a source of malodorous gas emissions. It is then only necessary to operate the recovery
boiler so as to minimize sulfur gas emission. The Scandinavian system also has a lower
moisturc coriteni in the exit gases from the recovery boiler than systems using direct contact
evaporators, therefore reducing plume opacity caused by condensed water droplets.
Indirect contact evaporation can also reduce sulfur emissions, but by introducing the water
vapor from the air cascade into the furnace, the combustion equilibrium conditions change
and the temperature decreases in the primary air combustion zone. The air-cascade
evaporation system may cause corrosion and particulate plugging of the rotary heat
exchanger.
Direct contact evaporators tend to have greater heat economy than air cascade evaporators,
but lower heat economy than complete use of additional multiple-effect evaporators (38).
The difference in heat economy grows in importance with increasing fuel costs. An
10-55
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additional factor in direct contact evaporation is that high degree black liquor oxidation can
rcducc the heating value of black liquor by as much as 5 to 10 percent (39).
10.7 Flue Gus Scrubbing for Gaseous Emissions
Flue gas treatment with absorption of the malodorous gases would be an economical
mcthod of converting old recovery boilers to meet present standards for air pollution if it
were not possible to rebuild the plant to the low odor system. This equipment can be
erected after the precipitator during the operation of the boiler and connected in a
reasonably short time if space is available. Any changes requiring a long downtime for the
recovery boiler are prohibitive because of production loss. With fuel prices below S I .42 per
million kJ ($1.50 per million I3TU), it is almost never feasible to rebuild an existing unit to
gain better heat economy even though a new design would be feasible for a new mill.
In an existing recovery boiler, it ma ’ not be possible to operate without release of SO 2 and
1-12 S. The existing direct contacL evaporators can be used to absorb SO 2 . The absorption of
l-1 2 S may also be possible, but in most cases insufficient to meet the regulations regarding
emissions of this gas.
Methods and equipment have been designed for the absorption of 1 12 S. Three of these
systems are commercially available. Only the last type has been installed in more than one
mill. These systems are:
The B. C. Research Council absorption scrubber,
The TRS Weyerhaeuser absorption scrubber, and
The Gadelius.Misu bish i absorption scrubber.
The B. C. Research Council’s (B.C.R.C.) method uses a rather concentrated Na 2 CO 3
solution for absorption of the H 2 S (29, 40). The principle is that the carbonate
concentration should be in equilibrium with the CO 2 partial pressure in the gas to avoid
excessive loading with Na 2 CO 3 and to avoid an increase in the lime consumption in the
reeausticizing department. The carbonate solution is sprayed at the top of a packed tower,
and the liquor moves downward countercurrent to the flue gas.
The liquor is extracted at thc bottom of the scrubber and is pumped to an oxidizing unit
where Na 2 S is converted to Na 2 S 2 O 3 . Enough iron compounds are normally available in the
remaining dust after the precipitator to act as a catalyst for a rapid conversion to Na 2 S 2 03.
This conversion was observed during [ he operation of scrubbcrs for generation of hot water
by heat recovery from flue gas (41).
10-56
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The method is the property of B.C.R.C., and SF Products Canada, Ltd., manufactures and
markets the equipment. It has been discussed in combination with heat recovery for
generating hot water for the bleach plant, in which case (lie feasibility of the installation
seems satisfactory, especially at the present level of fuel prices. The pressure drop on the
flue gas side depends on the concentrations of 1-125 and Cl-I 3 51-I in the flue gas, as these
determine the number of exchange units in the packed tower. One application was
calculated at 750 Pa water (3 inch w.g.) pressure drop for reduction from 600 ppm to
5 ppm, and the same pressure drop for the heat recovery section. Compressed air is used for
the oxidation unit. A bleed-off of the scrubbing liquor and fresh alkali make-up is necessary
to maintain the correct liquor composition. A typical arrangement is shown in Figure 10-25.
The TRS System was developed and patented by the Weyerhaeuser Company (42) and uses
a Nalco water solution of chelated ferrie chloride in a proprietary formulation. The
absorbed 1-12 S forms elemental sulfur, and a special packing is used to avoid plugging. This
package was developed and is marketed by Fritz W. Glitseh & Sons, Inc.
The IRS-scrubber absorbs up to 99 percent of the 1-12 S and collects about 85 percent of the
particulate matter. The resulting slurry of colloidal sulfur, salt, and other materials is passed
through a thickening and washing operation for recovery of the sulfur. Soda ash and air are
used to keep the solution neutral and to oxidize the Nalco water solution. The solids
concentration is kept at about 20 percent by bleeding to the green Liquor. The sulfur, salt,
and Naleo solution are recovered. Several scrubber units are used in parallel; one can be
taken out for cleaning, or the packing elements can be removed for cleaning. The pressure
drop with all units in operation on the gas side is about 1250 Pa water (5 inch w.g.). A
typical arrangement is shown in Figure 10-26. The above method has the advantage of
removal of sulfur from the kraft process. Such removal may become a necessity in the
future to keep the sulfidity of the white liquor in a range such that corrosion effects on
equipment remain tolerable.
The Cadelius-Mitsubishi method uses one or two stages of spray nozzles in countercurrent
arrangement to the flue gas flow for the absorption of l-1 2 S in Na 2 CO 3 or NaOH. Fresh
alkali is used for the niakeup of the absorbing solution. Several units are in operation in
Japan. This scrubber has very low pressure drop on the gas side, about 245 to 375 Pa (1 to
1.5 inch w.g.). [ None of these methods will have a high absorption efficiency for Cl-1 3 5H,
CH 3 SCH 3 , or Cl- I 3 SSCH 3 . The absorption efficiency for the other sulfur compounds must,
therefore, be quite great in some eases to bring the TRS down to the limits of the
regulations. This fact should be considered at the design stages for flue gas scrubbing
equipment.
Other methods for removing reduced sulfur compounds, such as scrubbing with an alkaline
suspension of activated carbon, were suggested after laboratory studies of absorption (43,
44). The recently recognized importance of both increasing the furnace temperature in the
10-57
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Possible Heat
Recovery Section
Demi stei
Spray Nozzles
Absorber
Flue Gas
Water For
Cooling & Saturation
Possible
Dust Recovery
clrculatloj Discharge To
Pump Pulp Mill
Section ..- Na 2 CO 3
N a 2 S 2 O 3
Na 2 S
NaSH
FIGURE 10-25
BRITISH COLUMBIA RESEARCH COUNCIL DESIGN FOR
H 2 S ABSORPTION SCRUBBER
,I H 2 5—”Free” Flue Gas
Alkali
Make Up+
(Water)
(Air— 0 2 )Exit
Oxidizer
10-58
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SALT CAKE
RECOVERY
S
Return
Scrubbing
Liquid
.
. .. . S
Evaporotion :
Crystal— • s .ii •. . SI . .. .. . 555 5555S••ll
lization
units REGENERATION TANKS
‘NaA V
MAKE-UP RECYCLE
FIGURE 10-26
ABSORPTION TOWERS
Gases
from
Furnace
Evaporator
S
.
S T C K
Spent ScrubbingLiqud
SI S
....S.S...•. •SISS.SSSSSISS
• S
SI S S S
SULFUR
RECOVERY
GLITSCH-WEYERHAEUSER DESIGN FOR A TRS SCRUBBING SYSTEM
-------
PSYCROMETRIC CHART .
Flue Gas Conditions
Ambient Air
DRY BULB TEMPERATURE
The mixing lines show extention
HUMIDITY
RATIO
I I
Ia Ia
Ambient Air Condition
Condensation Yes
From ® to ®
3 3a 3 3a
No No No
FIGURE 10-27
FORMATION OF VISIBLE PLUME THROUGH
CONDENSATION OF WATER VAPOR
Flue Gas Condition
10-60
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primary air combustion zone to reduce the emissions of sulfurous gases (9, 11, 13, 14); and
increasing the residual alkali to reduce the formation of some of the malodorous
compounds, possibly needs further examination since this method can probably reduce the
emissions of malodorous gases. These changes probably arc easier to adopt operationally
than an absorption system.
502 can be removed from the flue gases by any of several designs of simple scrubbers. Two
of these are the SF Scrubber-Moclo System and the Warkaus scrubber (45), which is a
double venturi scrubber arranged with the gas and scrubbing liquid in parallel flow with
practically no pressure drop on the gas side.
One major problem, when using a scrubber installation, is that the flue gas becomes
saturated and the visibility of the plume from the stack increases considerably. This problem
is of less importance if the scrubber is used also for hot water generation for the bleach
plant. The humidity of the flue gas is then reduced, therefore, reducing the plume. The
plume can be virtually eliminated if cold water is available to reduce the humidity ratio still
further. (See Figure 10-27.) Cooling of the flue gas adversely affects the plume buoyancy.
10.8 Collection of Particulate Matter from Recovery Boiler Flue Gas
10.8. t General Conditions
Formation of the particulate matter in the flue gas was explained briefly in 10.2.4. The
amount of dust, its particle size distribution, and its handling characteristics depend on the
reaction conditions in the recovery boiler.
The dust load varies between 40 and 75 kg per metric ton of dry solids (80 and 150 lb/ton).
This gives a range from 140 to 680 kg per metric ton of pulp (280 to 1360 lb/ton) for
extreme combinations of operating conditions of recovery boiler operation and cooking
yield.
The dust concentration in the flue gas calculated on a dry gas basis at standard conditions of
15.5° C and 760 mm Hg (60° F and 29.92 in. Hg) will vary correspondingly from 9 to
25 g/m 3 (4 to 1 1 grldscf). Only the dust that requires high efficiency collection of small size
particles is included in these figures. Coarse dust from the sootb lowing, which will settle by
sedimentation at normal velocities in a settling chamber, is not included since this will be
separated in the distribution plenum before the collector elements and will not likely
present any burden to the dust collector.
The dust load can increase above the given figures if part of the dry solids is burnt in
suspension. The dust from the kraft process consists chiefly of Na 2 SO 4 and Na 2 CO 3 . The
concentration of Na 2 CO 3 depends mainly on the ratio between sodium and sulfur in the
10-61
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flue gas. Traces of NaCI and Na 2 CO 3 arc normally found at mills operating with scalogged
wood. The concentration of NaCI can be considerable. The concentration of dust in the gas
might also bc increased by pulping of sealogged wood. The dust particles can contain small
amounts of Na 2 S.
the present trend of increasing the temperature in the combustion zone to achieve an exit
flue gas from the boiler that is virtually free from H 2 S and SO 2 also probably increases the
dust content in the flue gas, as compared to recovery units of similar, but earlier design.
Increases of about 40 percent are possible for a new unit as compared to an old Linit at the
same mill and operating with the same kind of liquor. The most important data for the
design of a dust collector are the gas flow, the dust load, and the particle size distribution.
Design data from the recovery boiler manufacturer should be used with appropriate safety
margins for the dust collector to cover upset conditions.
The small particle size of the dust makes mechanical dust collectors unsuitable for cleaning
of flue gas from recovery boilers. Baghouses are not suitable because of the handling
characteristics of the dust. Electrostatic prccipitators are, therefore, used for recovery boiler
installation. The economic collection efficiency based on the price of the recovered
chemicals, capital and operation costs, and payoff time are estimated at between 92 and 97
percent for different conditions. Such a collection efficiency generally will not control
particulate emissions to the degree needed to meet air pollution control regulations.
Wet scrubbing of the flue gas with water or with thin black liquor, which was previously
oxidized at an efficiency of 99 percent or more and stripped of methyl mercaptans and
organic sulfides, can provide sufficient collection efficiency to meet air pollution control
regulations. This alternative is, however, less attractive as the chemicals are present in an
aqueous solution. The discharge back into the process causes an increase of the inactive
chemicals, operational difficulties, and slightly increased losses in other departments. This
method may, however, prove the most economical for an old mill under certain conditions.
A combination of an electrostatic precipitator with about 95 percent collection efficiency,
followed by a scrubber to achieve a total efficiency of 99.5 percent, is economically
favorable if used in combination with heat recovery from the flue gas. The application of a
tail end scrubber must be investigated thoroughly for particle size distribution. The
collection efficiency is high for dust larger than 1 pm (3.9 X 10 in), even with a low
energy scrubber (probably because of the hygroseopicity of the dust), but decreases rapidly
with decreasing particle size.
10.8.2 Electrostatic Precipitators
The functioning of an electrostatic precipitator is based on movement of charged particles
of dust in an electrostatic field. The emission electrodes arc given a negative potential
10-62
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ranging from 30,000 to 80,000 volts depending upon operating conditions. Thcy emit
electrons that charge the dust particles, and at the same time they form, together with the
grounded collecting electrodes, an clectrostatic field.
The high tension negaLivc current is achieved with a transformer and a set of rectifiers,
normally forming one unit and monitored by a spark rate control unit, designed to give a
certain number of flashovers per second. The theory and technology for electrostatic
precipitators are comprehensively treated by White (46) and Oglcsby (47).
Development work to increase the reliability of the operation and to decrease the capital,
operating, an(l maintenance costs is still in progress. The engineer must consider not only
the influence of the gas and dust characteristics, but also the process to which the
precipitator has to be applied. The dust and gas characteristics for the recovery boiler
process are more favorable for the precipitator operation with the North American recover)’
boiler system than with the Scandinavian system. This difference is mainly because of higher
water vapor content in the flue gas from American direct contact evaporators. This fact
shoLild be recognized when new methods for achieving a high black liquor dry solids
concentration arc considered, or when a change from sootbiowing with steam to
sootbiowing with compressed air is considered.
The negatively charged dust particles move to the grounded collecting electrodes (plates),
where they transfer a certain part of their charge, depending on the resistivity of the dust.
The dust particles arc kept on the plate by the electrostatic ficid and the remaining charge.
The dust is removed from the plates by rapping the plates. The acceleration in the surface of
the plates during the rapping must be sufficient to dislodge the dust from the plates by
shearing action. The dust falls downwards, mainly following the plates like a web at ideal
conditions, and is collected at the bottom of the precipitator. The ideal rapping system
should dislodge the whole dust layer with one shock wave passing through the plates.
During the shock waves, some flakes of dust that arc near the plates retain little charge.
These flakes arc shaken loose and are very easily entrained in the gas flow. The dust layer
should be allowed to build up to a certain thickness between rappings to minimize
reentrainment.
The discharge or emitting electrodes collect dust particles with a positive charge and
therefore need rapping. The dust collected on the emitting electrodes can vary between
needle-like deposits to thick layers of dust if they arc evenly distributed. The former type is
often found if the dust contains large amounts of chlorides, and the latter if the normal
Na 2 SO 4 .Na 2 CO 3 dust is sticky. The acceleration during the shock waves at the rapping
usually is about 20.40 times the acceleration due to the earth’s gravity (g’s) for normal dust.
It is very difficult to get an even distribution of the acceleration forces. The acceleration
must be increased to above 200 g’s if sticky dust is generated at combustion. High stresses
are caused above 200 gs on the components of the emitting and collecting system and
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decrease the periods bctween major maintenance. These figures relate to parts of the
collecting plates where the acceleration is at a minimum. \‘alucs considerably higher can
normally be measured at the points where the rapping forces arc applied.
The dust, collected at the bottom of the precipitator chambers, is discharged in different
ways. The dry bottom design Lised prior to the l95O s in North America was changed to a
wet bottom design because of required maintenance of the dry bottom conveyors. The
Scandinavian precipitators used dry and wet bottom designs at the start, buL changed to
mechanical convevors after experiencing corrosion with the weL bottom design. The
difficulties with the convcyors were eliminated by using a heavier steel bolted chain with
larger pitch (about 15 cm (6 in)) and bearings of graphite for the shafts to reduce
maintenance. Screw conveyors for the transverse transport at the end were changed to
Buhler convcyors of the same type as used for cement kilns.
The American trend has been toward dry bottom design with the low odor concept, but the
problems associated with heat distribution and heat insulation seem to have been
overlooked. The bottom of the precipitator is heated only by the collected dust as the ideal
gas velocity below the electric fields is zero. The same conditions are valid for the walls of
the precipitator chamber at the sides of the electric fields that are heated only by radiation
from the collecting plates nearest the walls. Any gas passing below the fields and at the sides
of the fields will decrease the collecting efficiency by “sneakage” and is almost intolerable
in high efficiency precipitators.
The collection efficiency is easily calculated assuming an even gas distribution (i.e., the same
velocity in all parts of a cross section of the precipitator fields), an even dust distribution in
the gas entering the precipitator, and an instantaneous mixing of the gas over the cross
section. Deutsch’s formula gives the collection efficiency as:
17 = I — exp (— wL/Ftc) Eqn. 10.1
where:
the collecting efficiency of the precipitator
w = the overall migration velocity of the dust in rn/s (ft/see),
L = the total effective length of the electric fields in m (ft),
R = the distance between the emitting wire and the collecting plates in m (ft),
c = the gas velocity in rn/sec (ft/see).
10-64
-------
Dcutsch’s formula docs not take into account the possible variation in w with the velocity,
the possible rcentrainmcnt of particles, or the influence of particle size distribution. This
formLlla is valuable to calculate the influence of limited changes (±15 percent) in gas load on
precipitator efficiency.
Attempts were made to refine the Dcutsch formula to apply it to calculating the difference
in performance of particles of different size and therefore cvaluate results from test
precipitators. Deductions based mainly on results of operation with sodium sulfate, dust
from pulverized coal firing, and certain metallurgical processes (48) show that the collecting
efficiency is:
T i I — exp (— wL/Rc)k Eqn. 10-2
where the K (dimensionless exponent) normally varies between 0.5 and 0.8.
Equation 10-2 gives less increase in the collection efficiency for a given relative increase in
collecting surface than the original Deutsch’s formula.
The above formulas can be written in another form that might give a better illustration of
the relationship between precipitator size and the gas flow versus the collection efficiency.
Observing that the velocity, c, is:
Eqn.10-3
V=AXBXL Eqn.10-4
where:
C = gas flow, m 3 /sec (ft 3 /sec)
A = effective height of fields, m (ft)
B = effective width of fields, m (ft)
V effective volume of electric fields, m 3 (ft 3 ).
Equations 10-1 and 10-2 can be written as:
1 - exp (-wV/RG) Eqn. 10.5
= I - exp (_wV/RG)k Eqn. 10-6
10-65
-------
in addition:
B2XRXn Eqn.l0-7
C=2 >Cn XA X L Eqn. 10-8
where:
n = number of passages in the field,
n + I = number of collecting plates in each field, and
C total effective collecting plate area, m 2 ( it 2 ).
Equations 10-5 and 1.0-6 can be rewritten as:
1— exp(-wC/G) Eqn. 10-9
= 1 - exp (- rC/G)k Eqn. 10-10
Based on a reasonably large number of tcsts from prccipitators on Scandinavian type
recovery boilers, equation 10-i 0 is recommended for use with k = 0.7 for estimating
purposes. For American low odor units that have a higher exit flue gas temperature, 205° C
or higher as compared to 160° C (400° F vs. 320° F), a value of k = 0.6 is recommended.
The relative changes in precipitator size with varying collection efficiencies are shown in
Figure 10-28.
Comparative tests with different designs of components for electrostatic precipitators must
be executed with very accurate control of the flue gas and dust conditions for the tests to be
of real value.
The following design parameters arc available for the design of electrostatic precipitators.
These arc:
I. Flue gas flow,
2. Flue gas temperature,
3. Flue gas composition of dry gas,
4. Flue gas moisture content,
10-66
-------
80
FILTER PLATE AREA, m 2
CON DITIONS
GAS VOLUME 5400m 3 /tD.S.
CO 16 - 17% by VOLUME
SOLIDS CONC. 60%
FIGURE 10-28
ELECTROSTATIC PRECIPITATOR SIZE AS A FUNCTION
OF COLLECTING EFFICIENCY (37)
5. Flue gas dust content,
6. Particle size distribution of dust,
7. Operating pressure for precipitator,
8. Variations in above data because of temporary overloading, sootbiowing, and
possible additional gas flows as from dissolving vent stacks, and mix tank vent
stacks, and
85
90
92
94
96
97
98
98.5
99.0
99.2
99.4
0
z
L i i
0
L I-
LI-
LIJ
2
0
F—
0
L ii
-J
-J
0
0
— WARM PRECIPITATOR —
— GAS TEMPERATURE -.
— 190°C
— 170°C
150°C
— 130°C — ...._
HOT PRECIPITATOR —
380°C —
I
GAS TEMPERATURE —
360°C
‘340°C
20 40 60 80 100 200 300 500
1000 1500
10-67
-------
9. Maximum possible variations because of changes in the pulp production
parameters. such as species and pulp yield.
The above list represents more than what was previously available for the design of a
precipitator. The demand for accurate information, however, has increased because of the
high efficiency currently required. A safety margin above the boiler manufaeture(s figures
for factors such as gas flow and temperature, should be allowed to accommodate the errors
caused by the measurement errors.
The following parameters can be determined for the electrostatic precipitator itself and the
regulations for particulate emission. These are:
1. Dust collection efficiency,
2. Number of parallel precipitator chambers,
3. Number of electric fields,
4. Number of transformer-rectifiers,
5. Gas inlet arrangement,
6. Type of rapping and rapping frequencies,
7. Maximum gas velocity in the preeipitators,
8. Type of dust discharge, wet or dry bottom,
9. Material of shell for precipitator chambers, and
10. Heating of shell and heat insulation.
The dust collection efficiency must be chosen with some consideration for the degeneration
that often takes place in the physical condition of an electrostatic precipitator. The
alignment of the emitting electrodes with respect to the collecting plates is important. The
alignment is easily upset during the exchange of the high voltage insulators if poor guidance
is provided for the adjustment. Because the collection efficiencies can decrease if the
alignment is faulty, performance guarantees, over a two-year period from startup, arc often
requested. Curves for changes in efficiency as function of various parameters, such as gas
flow and gas temperature, should be included in the guarantees.
10-68
-------
The number of fields and chambers must be determined, taking into consideration the
decrease in collection defficiency, if and when one field goes out of operation, causing a
subsequent change in outlet dust concentration.
The number of trans lormer.reetifiers must be determined for possible variations in the
specific dust load, temperature, and moisture content at the precipitator inlet. \‘arying dust
loads in different chambers distort the distribution of the current to different chambers if
the chambers are coupled in parallel. The more dust in the gas, the lower the vol tage and the
lower the emission (i.e., the chamber with the highest dust concentration will determine the
voltage on all parallel fields). But this chamber will emit less current than the other parallel
chamber(s). The efficiency of the fields will, therefore, decrease, especially in the field with
the high dust load. This will disturb the next field downstream. The first two fields in each
chamber should, therefore, have separate transformer-rectifiers, but coupling the third or
later fields in parallel to the same transformer probably is justified if the dust concentration
is very low in these fields.
The gas inlet arrangement should allow isolating the precipitator for maintenance, achieving
a good gas distribution, and avoiding buildup of dust in the lower part of the inlet plenum,
which will eventually disturb the gas distribution. Extension of the bottom conveyor to
cover the bottom of the inlet duct is a good solution if sufficient baffling is arranged to
avoid sneakage of gas below the fields. Guide vanes and gas distribution plates must be
rapped efficiently to give satisfactory performance for a prolonged period.
The rapping mechanisms should be sufficient to clean the emitting electrodes and collecting
plates without being actuated for extended periods. The main part of the dust layer
probably is discharged in the beginning of the rapping and extended rapping tends to break
away flakes of dust, causing “snowliaking.” Tests show large increases in dust losses with
increases above the optimum frequency of rappings.
The gas velocity in an electrostatic precipitator must be limited to avoid snowflaking. A
velocity not exceeding I .0 mIs (3.5 ft/see) is recommended for American design (47). This
normally is the practice also for Scandinavian precipitators, except where the precipitator is
followed by a scrubber for heat recovery, in which ease higher velocities arc acceptable.
The dust discharge can be accomplished with black liquor pumped over the bottom either
continuously or intermittently. Some designs use impellers in the bottom to stir the dust
into the black liquor. The black liquor is then discharged to the cascade evaporator or to a
mixing tatnk. The gas velocity between the baffles that prevent sneakagc of gas below the
electric fields is very low. Here the gas acts, to some extent, as a revolving gas volume at
approximately the wet bulb temperature. Corrosive conditions are easily reached in the
lower part of the fields, especially if the black liquor contains a high concentration of
10 69
-------
chlorides. Such conditions adversely affect the lifetime of the emitting electrodes and, in
some cases, have also damaged the lower parts of the collecting plates.
Using a dry bottom avoids these difficulties, but proper chains and scrapers with sufficient
sLreiigth and stability must be provided. The transverse conveyor, preferably at the inlet end
of the precipitator, should have sufficient capacity to accommodate the uneven feeding
from thc different fields. Drag chain convcyors of the same type as normally used for
cement kilns. give outstanding service. II scrcw convevors are used, they must have a
relatively large diameter and sufficient stiffness to avoid vibrations. The troughs for the
transverse conveyor should not have more depth than is the necessary minimum.
The reason for minimizing the depth is to keep the temperature high and avoid
condensation of water vapor into the hygroscopic dust. The troughs must be wcll insulated.
The (lust S preferably discharged via a roLary valve to seal against air leakage into the
precipitaLor chamber. (The induced draft fan operates better if it is placed on the clean side
of the precipitator.) Air leakage can cause local corrosion, which can disturb the gas
distribu Lion.
The rotary valve should be isolated from the mixing tank by a short screw conveyor with an
air screen to prevent diffusion of water vapor up into the rotary valve and the dust
conveyor. This arrangement has proved reliable in avoiding plug-ups of the rotary valve and
dust conveyors. The necessary amount of air for the air screen is less than 1.7 m 3 /min
(60 ft 3 /min) per precipitator chamber, and this air normally passes out through the vent on
the mixing Lank.
The normal precipitator housing in North America consists of tile walls and concrete roof
and little heat insulation. This design is susceptible to cracking with subsequent air leaks and
inside corrosion. The concrete housing used in Scandinavia was designed to avoid cracks by
proper consideration of heat conduction. This design was used as long as the gilled cast iron
economizer was still in use. The precipitator must lie water washed at intervals of a few days
to two weeks. The temperature of the flue gas to the precipitator can be as low as 1100 C
(230° F) for about one work shift immediately following a water washing. Hot precipitators
with steel plate chambers placed between the boiler outlet and the economizer that
operated at 400° C (750° F) or slightly lower, came into use in 1957. This design was
adopted for usc with the current long vertical tube economizer. The exit flue gas
temperature is around 160° C (320° F). Experience seems to indicate that a steel plate shell
can be used down to 127°C (260° F) if the precipitator is equipped with very good heat
insulation of at least 10 cm (4 in) of mineral wool above the top of the stiffeners for the
shell. All stiffeners and flanges in ducts must be insulated.
Heating of the shell is probably not necessary if good heat insulation is supplied. 1-leating of
precipitators was standaid practice before starting of the operation in Scandinavia, but has
10-70
-------
now been abandoned. No corrosion damage has been observed, to date. ‘l’he present practice
in the U.S.A. and Canada of not covering the stiffeners with sufficient insulation is,
however, damaging, and if it is economically feasible to use a heated chamber instead of
good insulation, the method should he used.
Dust emission data from recovery boiler e [ ecLrostatic precipitators used presently in the
U.S.A. are given in Table 10- 10. The values are average values of dceilc groups. The table
includes information from precipitators on 87 recovery boilers.
TABLE [ 0.10
AVERAGE PARTICULATE EMISSiONS FROM
RECOVERY BOIL ER ELECTROSTATIC PRECIPI-
TATORS iN TI-LE UNITED STATES (49)
Emission decile Average Emission rate
kglt lb/ton
First (lowest) 1.1 2.1
Second 1.7 3.3
Third 2.4 4.8
Fourth 3.4 6.8
Fifth 6.2 12.4
Sixth 8.5 17.0
Seventh 9.2 18.4
Eighth 14.2 28.4
Ninth 23.2 46.3
‘l’cnth (highest) 37.6 75.2
The first clecile represents dust concentration of 0.11 g/m 3 (0.05 gr/dsef), the Iii Lh
0.46 g/m 3 (0.2 gr/dscf), and the tenth 2.3 g/m 3 (I gr/dsef). The collection efficiencies arc
99 percent, 95 percent, and 80-90 percent, respectively.
10.8.3 Liquor Scrubbing of Recovery Boiler Flue Gas
The cyclone evaporator and the venturi evaporator scrubber were originally used for the
recovery of heat and for concentration of the black liquor to a level suitable for firing in the
furnace. The cyclone evaporator was a low energy type of scrubber that collected only the
coarsest size fractions of the dust in the flue gas. The venturi evaporator scrubber has high
energy requirements but is capable of collecting finer dust than the cyclone evaporator. The
dust collecting efficiency was decreased by the high viscosity of the concentrated black
]0-71
-------
liquor used in these types of scrubbers. This liquor cannot be atomized into sufficiently
small drops for efficient dust collection. The capital cost for the rccover ’ boiler department
was considerably decreased by the introduction of the venturi evaporator scrubber, as the
venturi scrubber and its associated enlarged induced fan compared favorably with alternative
types of heat recovery equipment with a low efficiency electrostatic precipitator. The
operating costs were probably never favorable because of high power consumption an(l
rather large losses of Na 2 SO 4 . A combination of high interest rate for the capital and a low
price for electric power can, however, sometimes justify the choice of the venturi evaporator
scrubber, if air pollution control is not an overriding factor. These types of scrubbers were
discussed previously in section 10.6.
To achieve high collection efficiencies for the particulate matter in the flue gas, liquors with
low viscosity must l)e used, such as thin black liquor or water. A scrubber using thin black
liquor discharges its liquor with the dissolved dust to a multiple-effect evaporation plant.
The increased load of Na 2 SO 4 and/or Na 2 CO 3 in the black liquor charged to the
evaporation plant increases the fouling rate. A comparison between a scrubber using thin
black liquor and a precipitator should, therefore, consider the changes needed in the
evaporation plant. Such changes include increasing the heating surfaces to compensate for
the increased fouling rate and to accommodate the increased boiling point rise. Another
factor to consider is the cleaning of the evaporation plant by boiling out with water and its
attendant disturbances in the operation. Oxidized liquor must be used to avoid emission of
H 2 S, and so allowance must be made for the corresponding heat loss from oxidation.
Water can be used in the scrubber; the dust is then collected and discharged as a water
solution. This solution cannot be concentrated very much before crystallization occurs. The
recovery of the chemicals will, therefore, have an influence on the evaporation plant of the
same magnitude as when black liquor is used as the scrubbing liquid. The emission of H 2 S,
however, is avoided if the dust is reasonably free from Na 2 S.
Existing precipitators which were designed for insufficient collection efficiency to meet the
present regulations for emission of particulate matter, or which have degenerated to a lower
collection efficiency because of design deficiencies and/or inadequate maintenance, can be
retrofitted with scrubbers. The dust amount collected in these scrubbers is reasonably small,
and influence on the recovery cycle is less than if the scrubber installation were to collect
the total dust load of the flue gas.
Using flue gas scrubbcrs in many cases provides an attractive solution to increasing the total
collecting efficiency of an existing recovery boiler plant. But the space requirement for an
additional precipitator may make it almost impossible to use an additional precipitator.
Operational problems can very likely occur if long horizontal ducts arc used to convey the
gases from an existing plant to an additional electrostatic precipitator. Using a flue gas
scrubber can then provide a practical solution that also results in a capital cost saving. A
10-72
-------
venturi type of scrubber is used in most such cascs but it has a.relatively low pressure drop
of 1500-2500 Pa (6-10 in. w.g.).
Another solution was practiced in a number of mills in cold climates. A low energy type flue
gas scrubber can be used for recovery of heat from the recovery boiler flue gas to produce
hot water for the bleaching plant (and possibly also for the brown stock washing). A
scrubber of this type is placed after the electrostatic precipitator to collect some of the
remaining dust from the flue gas. Combinations of electrostatic precipitators of 95 percent,
or even lower, collecting efficiency with a scrubber, recovering heat for producing hot
water, have given a combined collection efficiency of 99.5 percent and have operated
successfully for several years (41, 50). This is a very economical combination, as the cost of
the scrubbers is justified by both heat recovery for heating water and saving in the cost of a
smaller precipitator. Mills in warm or hot climates where hot water is usually available in
abundance cannot use this approach effectively. The decrease in the generation of back
pressure power is an important factor in an economic evaluation, because it can reduce the
value of the saving in fuel by up to 50 percent.
Use of scrubbers in existing mills to increase the total collection efficiency of the existing
precipitators should be carefully evaluated. The particle size distribution after the
precipitators is important to the collecting efficiency of the scrubber. The collection
efficiency declines rapidly with decreasing particle size. The same precipitator design can
have great variations in amounts of fine dust at the exit at operating conditions which seem
very similar when based on the data for the electric current and voltage in the precipitator
fields. Therefore, new equipment should be guaranteed by the supplier covering both the
efficiencies of the precipitator and the scrubber.
The efficiency of the scrubbing depends mainly on the contact surface area and relative
velocity between the water drops and the dust particles. The atomization of the liquid and
the relative velocity between the gas and the liquor drops can be achicv d in either or both
of two ways, acceleration of the gas and applying the liquor through high pressure
atomization nozzles. Pressures up to 10.3 MPa (1,500 psig) are sometimes used for the
atomization of water. Similar drop size is achieved by using steam or compressed air with a
pressure of about 0.79 MPa (100 psig). Atomizing by accelerating the gas to high velocities
can avoid clogging of the nozzles lithe liquid is recirculated. Atomization of the liquid by
using gas velocity consumes much more power than atomization by high pressure nozzles.
One particularly interesting design is the use of co-current water sprays. The impact from
the water drops will reduce the pressure drop in the venturi throat, and designs are available
in which the impact from the water drops compensates for the pressure drop of the gas in
the scrubber (45).
10-73
-------
10.9 Economy of Recovcr ’ Boiler Operation
Recovery of [ lie chemicals and heat from the black liquor dry solids is of vital importance to
the economics of pulp production. The great variations in climatic conditions and in 1 )tdp
yields for different Species of wood make it ‘cr ’ difficulL to give a general picture. The
prices for electric power also vary considerably and, consequently, affect the economic
feasibility of back pressure power generation. Power generation favors feed waler heating
with steam from the turbine extractions. The correspondingly higher feed water
temperature might make changes necessary in the arrangement of the economizer and air
heater and in the most economic exit gas temperature.
The changes in the design and operation of recovery boilers, as a result of the last few years’
accumulation of research data, might make them more economical. Diagrams for the capital
cost for recovery boilers, for electrostatic precipitators, and for the complete recovery boiler
department arc given in Figure 10-29. The price level is for January 1974. The price for the
boiler is given for steam conditions of 4.2 MPa (600 psig), 400° C (750° F) for low
back-pressure power generation and 8.4 MPa (1,200 psig) and 480° C (900° F) for
reasonably high back-pressure power generation.
Two flue gas temperatures, 160° C (320° F) and 250° C (480° F), were stated for the prices
for electrostatic precipitators. The lower price for the precipitator at 160° C (320° F) must
be compared with an increase in price for an air heater-economizer. The handling
characteristics of the dust arc, however, much better at the lower temperature.
Rather high exit gas temperatures were used in the U.S.A. and Canada as compared to
Scandinavia. They reflect the low fuel prices in North America. Figure 10-30 shows the heat
loss per °F per year in the exit gas for varying firing rates. The value of the heat loss per year
for the difference in exit flue gas temperature between American and Scandinavian practice
is shown for varying fuel prices in cost per million BTU of steam.
10-74
-------
Price Level Jan 1974
“Average Black Liquor Dry Solids”
A. Recovery Boiler, delivery & erection, 400°F exit gas
Low odor with economizer.
B Electrostatic Precipitator, collecting efficiency, 99 5%
C Complete R B Department, including building ventilation,
instrumentation, electric power supply, feed water
treatment, black & green liquor systems connected to
mill, steam lines to back pressure turbine, stack, &
dissolving vent stack condensor
2 3
Dry Solids, 106 Ib/d
5
Steam
psig/°F
1200/900
600/750
Flue Gas
400°F
320° F
FIGURE 10-29
CAPITAL COST FOR RECOVERY BOILER DEPARTMENT
Capital Cost
106 $
/
/
/
/
/
/
/
/
C
I
20
15 -
10 -
5
0
F
/
/
/
F
A
0
4
10-75
-------
“Average Black Liquor Solids”
Black liquor concentration 62%, excess airS 02 at exit.
400-320°F represents exit flue gas temperatures
accord ln9 to the American & Swedish design practice
respective’y.
S 40/lO 6 Btu
I
S/Year
(400-320)°F
80
x io
70
60
50
40
30
20
10
Dry Solids. 106 lb/day
FIGURE 10-30
FLUE GAS ENERGY LOSSES
Si .00/106 Btu
S 60/lO 6 Btu
F
/
F
I
I
I
I
I
/
/
I
I
/
/
/
I
I
I
I
I
I
I
I
I
I
Btu/°F, Year
4000
x10 6
3000
2000
1000
I
/
I
/
I
I
I
I
I
I
I
I
I
‘I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
I
0
Btu/ F.yr
/
/
I
/
/
/
/
$ 20/iO 6 Btu
I
I
I
I
I
I
I
F
I
I
I
I
I
I
I
I
I
,
I
I
I
I
I
I
F
I
I
I
I
I
/
I
I
I
I
/
I
I
0
0 1 2 3 4 5 6
0
10-76
-------
10. 10 References
1. Rydholm, S. A., Pulping Processes. New York, lntcrscience Publishers, 1965, p. 777.
2. ‘Passincn, K. Chemical Composition of Spent Liquors. p. 183. Cullichscn, J. Ileat
Values of Pulping Spent Liquors. p. 211. In: Proceedings of the Symposium on
Recovery of Pulping Chemicals. Helsinki, Finland, May L3-17, 1968. Finnish Pulp and
Paper Research institute and EKONO Oy, Helsinki, Finland, 1969. p 000.
3. Vegeby, A., Scandinavian Practices in the Design and Operation of Recovery Boilers.
Tappi, 49:103A-109A, July [ 966.
4. Alhojarvi, J., Summary Report on the Properties of Spent Liquors. in: Proceedings of
the Symposium on Recovery of Pulping Chemicals. Helsinki, Finland, May 13-17,
1968. Finnish Pulp and Paper Research Institute, EKONO, Helsinki, Finland, 1969,
p. [ 67.
5. Vencmark, E., Svensk Papperstidning (Stockholm), 59(18):629-640, 1956. (Swedish).
6. Vegeby, A., Unpublished investigation for Institute for Vattenoch Luftvoadsforskning,
Stockholm, Sweden.
7. Safe Firing of Auxiliary Fuel in Black Liquor Recovery Boilers. Black Liquor Recovery
Boiler Advisory Committee. April 1967.
8. Emergency Shutdown Procedure approved for Black Liquor Recovery Boilers, Black
Liquor Recovery Boiler Advisory Committee. April 17, 1968.
9. Bauer, F. W., and Dorland, R. M., Canadian Journal of Technology, 32:91, 1954.
10. SilIcn, L. C., and Andersson, T., Solid-Gas Equilibria of importance in Burning
Conventional Ca or Mg Sulfite IVaste Liquor. Svensk Pappcrstidning (Stockholm),
55:662, 1962. (Swedish).
11. Rosen, E., Calculations for the Gasification of Spent Cooling Liquors. Royal Institute
of Technology. Stockholm, Sweden. 1962.
12. Davidsson, S., and SicIling, 0., Corrosion of Carbon Steel in Black Liquor Recovery
Boilers. Royal Institute of Technology. Stockholm. 1968 (in Swedish).
13. Stelling, 0., and Vegcby, A., Corrosion on Tubes in Black Liquor Recovery Boilers.
Pulp and Paper Magazine of Canada 70(10):T236, August 1969.
10-77
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14. Lang, C. J., DeHaas, C. C., Gommi, J. V., and Nelson, W., Recovery Furnace Operating
Parameter Effects on S02 Emissions. Tappi, 56:1 l5 Juiic 1973.
15. Timmerman, J.. Physio-ChemicnL Constants of Primary Systems in Concentrated
Systems, Vol. 3. New York, lnterscicncc Publishers, 1960.
16. Wilson, A. W., “Big Sky” Mill Veather Montana Pollution Battle. Pulp and Paper,
45(9):77.8L, August 1971.
17. Rydholm, ibid, p. 610.
18. Arincrgron, G. D., Haglund, A., and Rydhoim, S. A. Reported by K. Passincn, Ref. 2
above.
19. Hultin, S. 0., In: Proceedings of Symposium on Recovery of Pulping Chemicals:
Helsinki, Finland, May 13-17, 1968. Finnish Pulp and Paper Research Institute,
EKONO, Helsinki, Finland, 1969. p. 167.
20. Lindholm, I., arid Stockman, L., Heat Evolution During Blade Liquor Oxidation.
Svcnsk Papperstidning (Stockholm), 65(19):755, [ 962.
21. jones, K. I-I., Thomas, J. F., and Brink, D. L., Control of Malodors from Pulp Mills by
Pyrolysis. Journal of Air Pollution Control Association, 19:501-504, July 1969.
22. Brink, D. L., Thomas, J. F., and Jones, K. H., Malodorous Products from the
Combustion of Kraft Black Liquor: III. Rationale for Controlling Odors. ‘ [ ‘appi,
53:837.843, May 1970.
23. Brossct, C., Chaimbers Institute of Technology, Sweden. Personal communication.
24. Hisey, W. 0., Abatement of Sulphate Pulp Mill Odor and Effluent Nuisances. Tappi,
34:1.6, January 1951.
25. Harding, C. I., and Galcano, S. F., Using IVeak Black Liquor for Sulfur Dioxide
Removal and Recovery. Tappi, 51 :48A.51A, October [ 968.
26. Galcano, S. F., and 1-larding, C. 1., Sulfur Dioxide Removal and Recovery from Pulp
Mill Power Plants. Journal of Air Pollution Control Association, 17:536.539, August
1967.
27. Maksimov, V. F., Bushmelov, V. A., Torf, A. 1., and Lcsokhin, V. B., Testing the
Turbulent Flow Venturi Apparatus, Bumazhnaya Prornyshlennost, (Moscow), 40: 14.
15, May 1965.
10-78
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28. Blusser, R. 0., Cooper, U. B. It., Duncan, L., Tucker, T. NV., and Megy, J. A., Factors
Affecting Gaseous Sulfur Emissions in the Kraft Recovery Furnace Complex. Paper
Trade journal, 153(2] ):58.59, May 26, 1969.
29. Murray, F. E., and Rayner, I-I. B., Emission of hydrogen Sulfide From Black Liquor
Daring Direct Contact Evaporation. Tappi, 48:588.593, October 1965.
30. VaIther, J. E., and Amberg, H. R., The Role of the Direct Contact Evaporator in
Con trolling Kraft Recovery Furnace Emi stons. Pulp and Paper Magazine of Canada,
72:65.67, October 1971.
31. Sarkanen, K. V., Hrutliord, B. F., .Johanson, L. N., and Gardner, H. S., Kraft Odor.
Tappi, 53:766.783, May 1970.
32. l3losscr, R. 0., Cooper, H. B. II., Duncan, L., Tuckcr,T. XV., and Megy, .J. A., National
Council of the Paper Industry for Air and Stream Improvement. New York. Technical
Bulletin. December 31, 1969.
33. I-lendrickson, E. R., Roberson, J. E., and Koogler, J. B., Control of Atmospheric
Emissions in the Wood Pulping Industry, Volumes I, II, III. Final R(iport, Contract No.
CPA22-69-18, U.S. Department of 1-Icaith, Education, arid Welfare, National Air
Pollution Control Administration. Raleigh, North Carolina, March 15, 1970.
34. Martin, F., Secondary Oxidation Overcomes Odor from Kraft Recovery. Pulp and
Paper, 43:125-127, .Junc 1969.
35. Clement, .J. L., and Elliot, j. S., Kraft Recovery Boiler Design for Odor Control. Pulp
and Paper Magazine of Canada, 70(3):47.52, February 7, 1969.
36. l-lochmuth, F. \V., An Odor Control System for Chemical Recovery Units. Pulp and
Paper Magazine of Canada, 70(8):57.66, April 18, 1969.
37. Air Pollution Problems of the Swedish Forest lndustnes. SLateiis Naturvardsvcrk
(Sweden). Publication 1969:3. 1969.
38. \‘egcby, A. ibid.
39. Arhippainen, B., and Jungerstam, B., Operatmg Erpertence of Black Liquor Evapora-
tion to High Solids Content. Tappi, 52:1095-1099, june 1969.
40. Oloman, C., Murray, F. E., and Risk, J. B., Selective Absorption of 1-lydrogen Sulfide
from Stack Gas. Paper Trade journal, 153(7):92.94, February 17, 1969.
10-79
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41. Vcgcby, A., Canadian Pulp and Paper Association, Montreal, Canada. Technical Paper
T242. January 23-26, 1968.
42. Murray, J. S.. Tappi Environmental Conference, San Francisco, May 15, 1973.
43. Bhatia, S. P., de Souza, T. L..C., and Prahaco, S., Removal of Sulfur Cornpoundsfroin
Kraft Recovery Slack with Alkaline Suspension of Activated Carbo:z. Tappi,
56:164-167, December 1973.
44. Teller, A. J., and Ambert, H. R., Considerations in the Design for TRS and Particulate
Recovery from the Effluents of Kraft Recovery Furnaces. Preprint TAPPI Environ-
mental Conference, May, 1975.
45. Jafs, D., Recovery of Heat and Chemicals from Flue Gas Using the iVarkaus Venturi
System. Papper Och Tra (Helsinki), 48(6):337-342, June 1966. (In English)
46. White, 1-1 ., industrial Electrostatic Precipitation. Reading, Addison-Wesley, 1963.
47. Oglesby, S. Jr., A Manual of Electrostatic Precipitation Technology. Final Report.
Contract No. CPA 22-69-73, United States Department of Health, Education, and
Welfare, National Air Pollution Control Administration, Raleigh, North Carolina,
August 1970.
48. Berg, B. R., Development of New Horizontal-Flow, Plate-Type Precipitator for Blast
Furnace Gas Clea,ung. Iron and Steel Engineer, 36:93-100, October 1959.
49. Atmospheric Emissions from the Pulp and Paper Manufacturing Industry. EPA-450/E-
73-002. September 1973. (Also published as NCASI Technical Bulletin No. 69,
February, 1974).
50. Sodcrstrom, J., in: Minutes from the Swedish Steam Users Associations Recovery
Boiler Conference. 1971.
10-80
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APPENDIX 10-1. AIR AND FLUE GAS QUANTITIES AT COMBUSTION OF BLACK LIQUOR DRY SOLIDS
Assume- Reduction of smelt is 100 R% NOTE Neither vapor from the water in the black liquor or air from soot-
No formation of Na 2 S 0 or Na 2 S 2 0 3 blowing has been included No air leakages have been considered
Dust losses after precipitator negligible
Dry Air is parts 02 and I — j4y “N 2 “(mcludes C0 2 , argon, etc.)
Humidity ratio of air is w lb/lb dry air
Volume ratio of moist air/dry air m I + w
Black liquor D.S composition is
carbon,C lOOc%
hydrogen, II tOO h%
sodium, Na 100 n%
sul(ur,S lOOs%
o .ygen, 0 (as difference) 100 o%
inert oxides 100 i%
c + Ii + n + + o + i = I
Theoretical complete combustion without ecccss air
Air, moles/lb D.S
Air
Product Quantity
moles/lb l 1 S
02
Flue gas, moles/lb D S
CO 2 1120 N 2 ‘l’otal
CO 2
— c r a
a
477 ma
a
477 (m - I) a
3 77 a
477 mj
1120
Ii
1 b
05b
05X477mb
—
[ 1 +05X477(m- 1)]b
05X 3771i
05h+05X477,nb
Na 2 S
R 7
—
—
—
—
—
—
Na 2 SO 4
(I- R) 7 d
2d
2X477md
—
2X 477(m- l)d
2X 377X d
2(477m- l)d
Na 2 CO 3
15e
L5X477me
—
15X477(m-l)e
15X377Xe
l5(477m 1)c
Corr.
for 02
At theoretical complete combustion without excess air
Air consumption, A = 4 77 m (a + 05b + 2d + I 5e - I) moles/lb D S
Dry flue gas flow, Fdry = 477 a + 3 77 (0 5b + 2d + I 5e - I) moles/lb D S.
Totalfluegasflow,Ft 0 t477m(a+03b+ 2d + I 5e- 1)+05b- 2d- I 5e+ fmoles/IbDS
Water vapor, F = 4 ?7 (m - I) (a +05b + 2d + I 5c - I) + b moles/lb D S
- IOOa
/°‘ -‘J2rnaX4l7a+377( 05b+2d+I5e_Q’°
The flue gas flow F 1 , at z% by volume o’ygen in dry g.is is FL = Fdry + . _z]+ Fv
f -f -477m1 — -477(m- l)f -377f -(477m- 1)1
32
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APPENDIX 10-2. HEAT VALUES VS. OXYGEN DEMAND FOR COMPLETE
COMBUSTION
Softwood
Hardwood
Carbo-
Variable Units Lignin
Lignin
hydrates
Analysts A B C
Carbon,C 64 60 46
Hydrogcn ,,H % 6 6 6
Oxygen, 0 % 30 34 48
Bomb heat value (X) Btu/lb 11340 10620 7560
Oxygen demand for complete 10 3 lbmole/lb 58.8 54.3 38.2
combustion
Analysis adjusted to “normal” A 1 B 1 C 1
content of inorganics
Carbon. C % 47.8 44.8 34.6
Hydrogen, H % 4.4 4.4 4.4
Oxygen, 0 % 25.6 28.6 38.8
Sodium, Na % 18.2 18.2 18.2
Sulfur, S % 4.0 4.0 4.0
Bomb heat value (0.734 X) Btullb 8300 7800 5550
“Efficient heat value in reducing Btu/lb 7340 6840 5890
atmosphere” (reduction for
sulfide, and heat of evaporation
for vapor from combustion of
hydrogen)
“Resulting heat value of black Btu/lb 6390 5890 3640
liquor” (above value adjusted
for heat of evaporation of
water in black liquor)
Oxygen demand for complete lO 3 lbmole/Lb 44.2 40.8 29.1
combustion
Note Sec Figure 10.23 regarding linear correlation.
10-82
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CHAPTER ii
LiME BURNING AND LIME DUSrI HANDLING
The causticizing of liquor, commonly called green liquor, from the smelt dissolving tank by
ddiLion of lime or calcium oxide (CaO) results in the generation of a lime mud or calcium
carbonate (CaCO 3 ) sludge. The lime and sludge are then washed and calcinecl at elevated
temperatures in either a rotary kiln or a fluidized bed calciner to recover calcium oxide. This
oxide can then be reused for reclaiming additional white liquor, the chemical solLition for
digesting pulp. The normal auxiliary fuels used as heat sources for lime mud burning arc
natural gasand residual fuel oil. The two major potential air polluLants from lime mud
burning are the gaseous emissions and the particulate emissions of entrained lime dust from
the burning zone. The gaseous emissions arc 1-12 S from the lime mud and, possibly, organic
sulfur compounds from the scrubbing watcr.
11.1 Rotary Lime Kilns
11.1.1 Design Features
The rotary kiln is the most commonly employed device for lime mud reburning in kraft
pulp mills. The device is an open.cnded inclined cylinder that is rotated so that lime mud
added at the upper end gradually passes to the lower end and drops out into a bin as dry
lime. Fuel and air flow countercurrently to the lime from the lower end of the kiln. The kiln
exhaust gases normally pass through a mechanical cyclone collector for lime dust recovery
and finally through a liquid scrubber for particulate control (1).
Rotary lime kilns employed in kraft pulp mills can range from about 2.4 to 4.0 m (8 to 13
ft) in diameter and from about 30 to 120 m (100 to 400 It) in length. They are designed to
burn 36 to 360 t (40 to 400 tons) of lime (as dry CaO) per day (2). The lime kilns are
normally inclined at a slope of about ten degrees from the horizontal plane and can be
supported by two- to four-supports, depending on their length. The lime kilns must be
designed with a number of auxiliary components, including a lime mud feed system, hot lime
conveying system, air inlet and preheating system, gas exhaust system, kiln rotation system,
and instrumentation systems (3). Major kiln design variables include kiln length, kiln
diameter, rotation speed and angle of incline, which influence solids retention time,
gas-solids contact area, and temperaLure.
11.1.2 Operating Parameters
Lime mud at 55 to 65 percent solids and with sodium content of 0.1 to 2.5 percent by
weight (as Na 2 0) enters at the upper end of the lime kiln and passes through successive
11.1
-------
stages of water evaporation, mud preheating, and lime calcination. Temperatures in the kiln
vary from 150 to 2600 C (300 to 500° F), at the upper or wet end, to 1200 to 1300° C
(2200 to 2400° F) at the hottest part of the calcination zone near the lower or dry end.
Energy requirements for the lime kiln operation arc for water evaporation, preheating and
calcining the lime mud, and power to rotate the lime kiln, drive the air fans and flue gas
fans, pump the scrubber liquid, and convey the lime mud and rcburned lime. The major
types of fuels burned in lime kilns arc natural gas and residual fuel oil; turpentine and coal
may also be used. Fuel requirements for lime kilns and fluidized bed cakiners are listed in
Table 11.1.
Two of the major design variables affecting particulate emissions from lime kilns are kiln
length and diameter. These variables can affect the amount of particles swept from the kiln
exhaust gases by governing the gas velocity and thc gas-solids contact area.
TABLE 11.1
ENERGY REQUIREMENTS FOR LIME MUD CALCINING
SYSTEMS (3) (4)
Rotary Kiln Fluid Bed Calciner
kJ/t (BTU/ton) kJ/t (BTU/ton)
2.3.4.7 X 106* (24 X 106)* 2.1.2.5 X 106 (1.8.2.0 X 10
9.3.17.4 X (8.15 X 106)** 8.0-9.2 X 106 (7.8 X 106)**
*per metric (t) or short ton (ton) of pulp.
**per metric (t) or short ton (ton) of lime, as CaO.
H 2 S emissions from the lime kiln are affected by the Na 2 S content of the lime mud
(particularly the aqueous phase) and by the presence of Na 2 S in the scrubber wash water.
Thc use of digester and evaporator condensate as lime kiln scrubber water can result in the
stripping of organic sulfur compounds into the exit flue gas. The presence of sufficient
excess air in the kiln can rcduce the concentration of H 2 S in the cxhaust by providing an
oxiding atmosphere sufficient for H 2 S conversion to SO 2 (5).
11.2 Fluidizcd Bed Calciners
Fluidized bed calciners are alternatives to rotary lime kilns for the calcination of lime
mud to lime. The lime mud is first washed to reduce soluble sodium compounds to a sod iuin
content of 0.1 to 0.5 percent by weight (as Na 2 0) and then dried on a vacuum filter. The
dried lime mud at 55 to 65 percent solids is then suspended in the flue gas from the
11-2
-------
fluidized beds at a temperature of about 1500 C (300° F) to evaporate the water. The solids
are then passed through a two-stage cyclone system to recover the dried lime mud solids and
then fed into a bed of fluidized lime pellets formed by calcination. The bed is kept in
suspension by the action of an air fan located below the cooling chamber from which
reburned lime is removed. Natural gas or fuel oil is injected into the suspended bed and
burned to provide the heat necessary for the calcination reactions to take place at a
temperature of about 825 to 875° C (1500 to 1600° F). The entrained particles and the
combustion gas products then pass out from the calciner to entrain the wet mud and pass
through the two-stage cyclone system and a venturi scrubber for particulate removal (5).
Fluidized bed calciners arc employed at several kraft pulp mills and have lime burning rates
ranging from 23 to 136 t per day as CaO (25-150 ton/day). Fuel requirements for fluidizcd
bed calcincrs are generally lower than for lime kilns because of the small combustion
chambers used which have smaller radiation heat losses. Electricity requirements for
fluidized bed calciners, however, are generally greater than for rotary kilns because of the
energy required for suspending the bed and operating the venturi scrubber. Major operating
variables affecting fluidized bed reactor operation arc the mud drying temperature, the
calciriation zone temperature, the excess air level, the bed fluidization level, and the sodium
content of the lime mud.
Major operating variables affecting particulate emissions from the calciner unit are the air
sweep velocity and the lime feed rate. Variables affecting gaseous emissions from fluidized
bed calciners are the same as those affecting reduced sulfur emissions from rotary lime kilns.
11.3 Particulate Emission Control
The major means of controlling particulate emissions from lime kiln and fluidized bed
calciner exhaust gases are liquid scrubbing, using either an impingement or venturi-type
scrubber, and, recently, electrostatic precipitation. The scrubbing devices are usually placed
following a mechanical cyclone collector used either for removal of the larger lime dust
particles, as with lime kilns, or for predrying the lime mud for fluidized bed calciners.
Particulate inlet loadings to scrubbing devices from lime kilns can range from 7 to 35 g/m 3
(3 to 15 gr/cu ft) at standard conditions of 21.1° C, 1.0 atmosphere, dry gas (70° F,
29.92 in. Hg, dry gas). The dust losses constitute about 1 to 5 percent of the total dry solids
load to •the kiln. Particle size measurements for the above mass concentrations are not
reported with these data, but the lime particles generally comprise the larger sizes and
sodium particles the smaller ones (4). Comparable data are not available for fluidized bed
calciners, but it is necessary to use a two-stage cyclone mud drying system to avoid
overloading f he venturi scrubber.
11-3
-------
11.3.1 Scrubbing Systems
The major types of scrubbers cmploved on lime kilns, to date, are the impingement and
venturi types, with e clonic scrubbers also employed, but to less extent. Impingement type
scrubbers were extensively employed in the past for particulate scrubbing on lime kilns and
have the advantages of relatively low pressure drop and scrubber shower rate, with resultant
reduced operating costs. The devices are limited, however, in their maximum scrubber slurry
water solids concentrations due to possible scrubber plugging. In addition, they normally
have lower particulate removal efficiencies because of less efficient gas.liquid contact.
Impingement type scrubbers have higher capital costs than venturi scrubbers on similar
installations basically due to their larger size and greater complexity.
\‘enturi scrubbers are commonly used on lime kilns at the newer kraft pulp mill installations
primarily because of higher particulate removal efficiencies than achievable by the older
impingement type scrubbers. Venturi scrubbing systems can operate with slurry water solids
concentrations of up to 30 percent by weight without excessive plugging. A summary of
operating characteristics for kraft lime-kiln scrubbcrs is presented in Table 1 1 .2 (4).
TABLE 11.2
OPERATING CHARACTERISTICS FOR PARTICULATE LIQUID
SCR UBBERS EMPLOYED ON KRAFT LIME KILNS (4)
Scrubber Type
Parameter Impingement Venturi
Shower rate ratio, 1/rn 3 0.54.2.0 1.73-3.21
(gal/b 3 ft 3 ) (4.15) (13.24)
Slurry solids, % by wt. 1.2 10-30
Pressuredrop,mmHg 9-13 19-28
(in H 2 O) (5.7) (10-15)
Power required,* kW per t/day 0.041-0.049 0.082-0.099
(hp per ton/day) (0.05-0.06) (0.10-0.12)
Power required,** kW per t/day 0.13-0.16 0.27-0.34
(hp per ton/day) (0.16-0.20) (0.33-0.42)
*pcr mass of pulp.
**per mass of lime.
11-4
-------
11.3.2 Performance Characteristics
A number of studies were conducted to determine particulate collection efficiencies of lime
kiln scrubbers. Stuart and Bailey (6) report that venturi scrubbers were able to achieve
96-97 percent particulate removal from lime kiln exhaust gases at pressure drops of 1.7 to
2.8 kPa (7 to 10 in water); while Landry and Longwell (7) report that venturi scrubbers can
achieve particulate removal efficiencies of 98-99 percent at pressure drops of 2.4 to 3.7 kPa
(10 to 15 in water). A series of studies were conducted on a joint basis by the National
Council for Air and Stream Improvement and the U.S. Environmental Protection Agency to
establish the particulate collection efficiencies of 66 existing lime kiln scrubbers. Venturi
scrubbers were able to produce consistently higher particulate collection efficiencies than
the impingement scrubbers, as shown in Table 11.3 (2).
TABLE 11-3
PARTICULATE COLLECTION EFFICIENCIES FOR LIQUID SCRUBBERS ON
KRAFT PULP MILL LiME KILNS (2)
Impingement Scrubbers Venturi Scrubbers
Parameter Average Range Average Range
Inlet concentration,* g/m 3 27.38 8.00-33.96 18.60 5.85.31.83
(gr/cu It) (11.94) (3.50.14.81) 8.11 (2.55-13.88)
Outlet concentration,* g/m 3 1.78 0.99-3.56 0.73 0.27-2.29
(gr/cu ft) (0.78) (0.43-1.56) (0.32) (0.12-1.00)
Removal efficiency, % by wt. 92.2 86.8-96.9 94.8 85 .5-99.1
Emission rate,** kg/t 1.78 1.14-2.09 1.01 0.33-2.60
(or lb/ton) (3.55) (2.28.4.18) (2.02) (0.66-5.19)
*Concentrations are reported at standard conditions of 21.10 C and 760 mm Hg (70° F and 29.92 in Hg),
dry gas.
**Emi ion rates arc based on an air-dried ton of pulp basis (i.e., 10% moisture, by weight).
Information developed during the study indicates that high pressure drop venturi scrubbers
can achieve significantly lower particulate levels than reported in Table 11-3 (2). Particulate
concentrations at standard conditions of between 0.02 and 0.11 g/m 3 (0.01 to
0.05 gr/cu ft), corresponding to emission rates of 0.01 to 0.05 kg per air dried metric ton of
pulp (0.02 to 0.1 lb/ton), were measured. Very little information exists regarding particulate
emission control following fluidized bed calciners. Erdman (8) reports on a high pressure
drop venturi scrubber following a two-stage cyclonic mud drying system. The pressure drop
through the venturi is 5.4 kPa (22 in water). Although the dust carryover from the calciner
11-5
-------
section is 12 percent, the scrubber emits a particulate concentration of 0.16 g/m 3
(0.07 gr/cu ft), which corresponds to an emission rate of 0.24 kg per metric ton of pulp
(0.49 lb/ton).
11.4 Gaseous Emission Control
Lime mud calcining in rotary kilns or fluidized bed reactors can emit H 2 S, organic sulfur,
SO 2 , and nitrogen oxides to the atmosphere. The gaseous emissions result either from
materials entering the calcining unit system or from materials entering the kiln. The major
process operating variables affecting gaseous emissions include excess air level, operating
temperature, and solid and gas.phase retention times.
Major input material properties affecting gaseous emissions include the respective Na 2 S
contents of the input lime mud and scrubber water, organic sulfur levels in the inlet
scrubber water, and the moisture content of the lime mud. The major design variable
affecting gaseous emissions from the calcining system are the length and, to a lesser extent,
the diameter for rotary kilns, and the diameter and height for fluidized bed calciners.
A summary of gaseous emissions from rotary lime kilns and fluidized bed calciners is
presented in Table 11-4.
TABLE 11.4
GASEOUS EMISSIONS FROM KRAFT PULP MILL LIME KJLNS (2)
Gaseous Concentration Emission Rate
Constituent Average Range Average Range
ppm, by volume kg sulfur per t pulp
(lb sulfur per ton pulp)
H 2 S 108 0.500 0.24 (0.48) 0.1.88 (0.3.76)
CH 3 SH 14 0-90 0.03 (0.07) 0-0.17 (0.0.34)
CH 3 SCH 3 27 0-245 0.02 (0.05) 0-0.22 (0.0.43)
CH 3 SSCH 3 5 0-11 0.01 (0.03) 0-0.10 (0-0.20)
TRS — — 0.31 (0.63) 0-2.37 (0.4.73)
SO 2 34 0-140 0.14 (0.28) 0-1.11 (0-2.20)
11-6
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11.4.1 LimeMud
Thc most important gaseous emissions from lime reburning systems are malodorous reduced
sulfur compounds. Hydrogen sulfide can be volatilized from the Na 2 S present in the lime
mud by contact with CO 2 from the flue gas. Above a threshold Na 2 S concentration of 0.2
percent by weight, the generation of H 2 S is directly proportional to the residual Na 2 S
content of the lime mud. This linear relationship is similar to that for direct contact
evaporation (9). The amount of H 2 S released can be controlled by reducing the residual
Na 2 S level by more efficient lime mud washing. it is not normally feasible, however, to
reduce the residual sodium content in the lime mud to less than 0.1 percent by weight
because of possible mud ring formation.
Prakash and Murray (5) report that H 2 S emissions from the lime mud occur primarily from
Na 2 S dissolved in the aqueous portion and not from the solid portion of the lime mud. The
H 2 S emissions can be reduced by drying the mud to a solids concentration of 70 percent by
weight or more before burning.
1] .4.2 Scrubbing Water
The scrubber water can be a source of both H 2 S and organic sulfur compounds emissions
from kraft mill calcining units equipped with scrubbers. The presence of Na 2 S in the
scrubber water can result in the release of 1-12 S by contact with CO 2 if the liquid ph is
sufficiently low. The emission rate of H 2 S and organic sulfur compounds increases with the
inlet Na 2 S and organic sulfur concentrations, with rising liquid- and gas-phase temperatures,
and with an increasing degree of gas-liquid contact, as represented by the scrubber pressure
drop. The potential for organic sulfur release is particularly great if untreated digester or
evaporator condensates are used as lime kiln scrubber makeup water.
Caron (11) reports that using lime mud wash water instead of fresh water for lime kiln
scrubbing results in stripping of 0.10 to 0.22 kg sulfur per metric ton of pulp (0.2 to 0.4 lb
sulfur/ton) as compared to an absorption of only 0.035 kg sulfur per metric ton of pulp
(0.07 lb sulfur/ton) with fresh water. Normally, fresh water should be employed as the
scrubbing medium to avoid the stripping of odorous gases. If condensate waters are
employed, steam stripping should be employed prior to the scrubbing operation.
One U.S. mill has significantly reduced TRS emissions from a lime kiln venturi scrubber by
adding sodium hydroxide to the scrubber water to raise the pH. The scrubber water is
recycled to the causticizing system (10).
11.4.3 Combustion Variables
The major combustion variables that can affect reduced sulfur emissions from lime mud
calcining operations are the excess air level, the temperature profile, and the mud retention
11-7
-------
time in the kiln. Caron (11) reports that the TRS emissions from the combustion zone are
minimized at excess oxYgen levels of four percent by volume or greater. Though no definite
patterns have been established, the kilns that have cooler wet-end temperatures tend to have
relatively higher reduced sulfur emissions because the sulfur compounds can be volatilized
without burning. Sufficient retention time must be provided at temperatures above 7600 C
(1400° F) to oxidize the reduced sulfur compounds.
Walthcr and Amberg (12) report that shorter lime kilns tend to have lower reduced sulfur
emissions than longer lime kilns, though no definite correlation could be established. The
probable reason is that short lime kilns must operate at higher average temperatures
throughout than the long kilns to achieve an equivalent degree of calcination. The result is a
more complete oxidation of reduced sulfur compounds. An additional factor is that the
evolution of Na 2 S at low temperatures in oxidizing atmospheres promotes H 2 S formation;
its evolution at higher temperatures promotes SO 2 formation (13).
Limited data indicate that reduced sulfur emissions from fluidized bed caleiiiers are
minimal. This may be due to the relatively long retention time at uniformly high
temperature which provides for efficient oxidation of the sulfur compounds (14). One test
shows an emission rate of less than 0.01 kg sulfur per metric ton of pulp (0.02 lb/ton). Flash
drying of the mud tends to minimize H 2 S formation in the fluidized bed units.
The burning of digester and evaporator noneondensable gases in the lime kilns brings an
additional source of sulfur compounds to the units. The conversion of these materials to
SO 2 is essentially complete because they are added with the primary air at the hot end of
the lime kiln and so have sufficient retention time for complete combustion to take place
(15). The addition of green liquor dregs with the lime mud to the cold end of the lime kiln
can substantially increase the reduced sulfur emissions, because these materials are normally
contaminated with Na 2 S from the green liquor. There is also insufficient retention time at
high enough temperatures for complete oxidation to take place.
11.4.4 Sulfur and Nitrogen Oxides
The concentrations of sulfur oxides in lime-kiln exhaust gases are normally minimized
because the CaO can aet as an efficient adsorption and reaction medium to form CaSO 3 and
CaSO 4 . Long kiln length, with sufficient oxygen and high calcination efficiencies, promote
efficient SO 2 removal. To date, no adverse effects on lime kiln operating efficiency were
traced to the sulfur released by the burning of either residual fuel oil or noneonderisable
gases. In a limited series of tests, it was not possible to measure the presence of SO 2 in the
exhaust gases of a fluidized bed ealcincr, probably because the calcining and flash drying
provided a two-stage removal system.
11-8
-------
Galeano and Leopold (16) report that the lime kiln is the only major process source where
significant quantities of nitrogen oxides can be measured. The primary reasons for the
presence of oxides of nitrogen are that there is sufficient excess air at a temperature of 1200
to 1300° C (2200 to 24000 F) to promote the reactions. The amounts of nitrogen oxides
formed in the fluidized bed calciners are probably significantly less than from rotary kilns
because of the lower operating tcmperatures of 825 to 875° C (1500 to 1600° F). To date,
no specific tests have been conducted to determine the amount of oxides of nitrogen in
fluidized bed ealciner exhaust gases.
11.5 Oxygen Addition
Modecular oxygen can be added to the combustion air of a lime kiln to control 1125
generation from the lime mud in the combustion zone. The oxygen must be added together
with the primary air to the firing zone at the dry end of the lime kiln. This practice will
promote effective mixing with the combustion gases and will provide for complete
oxidation of any l l2 S released from the mud. Precautions should be taken to assurc that
overheating of the kiln does not occur in localized areas. Such overheating could damage
refractory materials, interfere with kiln operation, or result in increased emissions of oxides
of nitrogen.
There is very limited field experience, to date, with the addition of oxygen to lime kilns for
reducing 1-125 emissions. Singman (17) reports that oxygen addition to lime kilns can
substantially increase the lime mud throughput rates for previously overloaded lime kiLns
without excessive lime losses. An addition of 0.454 kg (1 Ib) of oxygen results in a net
decrease in lime makeup rate of 1.8 kg (4 Ib) as CaO and, consequently, a considerable
savings in operating costs for causticizing. Decreases in H 2 S emissions may result for
relatively short lime kilns, particularly where higher temperatures are maintained at the wet
end. Additional process variables that would affect 1-125 emissions in addition to kiln length
include kiln diameter, mud washing efficiency and inlet sulfide Level, mud firing rate and
solids concentration, and gas velocities at different locations in the kiln.
11.6 Process Economics
The primary economic factor to consider for effective air pollution control of lime-ealcining
systems is the installation of devices for particulate control. The respective capital and
operating costs for impingement and venturi scrubbing devices are presented in Table tI- S
(4).
Gaseous emission control does not normally require substantial capital investment unless
flash drying of lime mud must be instituted. Maintenance of sufficient excess air, proper
washing of lime mud, and the use of fresh water normally are sufficient to minimize gaseous
11-9
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TABLE 11-5
CAPITAL AND OPERATING COSTS FOR. LIME KILN
PARTICULATE SCRUBBERS (4)
Scrubber Type
Cost Item Impingement Venturi
Capital cost,* S/daily t pulp 30-33 22-27
(S/daily ton pulp) (27-30) (20-25)
S/daily t lime 99-1 ]0 71-93
(S/daily ton lime) (90-1 00) (65-85)
Annual operating cost,** S/t pulp 3-7 7-1 1
(S/ton pulp) (3-6) (6-10)
$/tlime 11-22 22-38
(S/ton lime) (10-20) (20-35)
*Based on 1966 data.
**Ba d on 0.9 cents/kWh.
emissions. Addition of NaOH to the particulate scrubber makeup water to minimize H 2 S
emissions by increasing liquid pH levels will increase operating costs.
11.7 Lime Dust Handling
A minor source of fugitive particulate emissions from the causticizing system of a kraft pulp
mill consists of lime dust relcases from storage tanks and bins, and conveying and transfer
facilities. Activities where lime dust is loaded or unloaded, dumped, or transferred arc
particular problems because of the dryness of the material that is handled. The lime dust is a
localized emission source within the immediate area of the causticizing plant.
The three major approaches to control fugitive lime dust emission are to:
1. Confine the potential emission sources to prevent air leakage,
2. Wet the dust to prevent its becoming airborne by wind or by transfer operations,
a tid
3. Usc special air pollution conLrol equipment.
I I-tO
-------
The first method is effective in limiting thc potential sources or fugitive cmissions by
effective housekeeping. It also facilitates the subsequent installation of particulatc air
pollution control equipment. Wetting the dust is effective in controlling fugitive dust
emissions, but it can make the lime difficult to handle if overdone. It generally is not
recommended as an effective Lechniquc for controlling lime dust emissions.
The third method for controlling dust emissions from lime storage and transfer facilities
involves the use of particulate control techniques, such as centrifugal separation, liquid
scrubbing, and fabric filtration. Keeping the dust us dry as possible facilities its recovery;
therefore, liquid scrubbing is undesirable. Centrifugal collectors are not advantageous in that
they require high pressure drops and tend to have low collection efficiencies.
The only known installation for particulate lime dust recovery from storage facilities
employs a fabric filter baghouse for a 907 metric ton per day (1000 ton/day) kraft pulp
mill. The air vents from the lime storage tanks are vented into a central duct and passed to a
baghouse wiLh a design flow rate of 4,400 m 3 /h (2,600 elm) at a maximum temperature of
2900 C (550° F). The filter bags have a total surface area of 121 m 2 (1,300 It 2 ) with a
cleaning cycle of once each 20 minutes. Thc filter bags are made of a siliconized glass cloth
with a design ratio of air to cloth filter area of 22.8 to 27.4 rn 3 /h/m 2 (1.25 to 1.50
cfm/ft 2 ).
Total capital cost for the system was $5,900 in 1966. The fans have a total capacity of
13.4 kW (18 hp) with a resultant dircct annual operating cost of $1,080 per year. The
system can recover 225 to 450 kg/day of lime (500 to 1,000 lb/day), which is equivalent to
an annual savings of $1,800 to $3,600, based on a lime price of $22/t ($20/ton) (18, 19).
11.8 References
1. Libby, E. C. (ed.)., Pulp and Paper Sctence and Technology, Volume 1, Pulp. New
York. McGraw-Hill Book Company, 1962. p. 211-227.
2. Atmosphertc Emisstons from the Pulp and Paper Manufacturing Industries. Cooperative
NCASI-USEPA Study Project, Publication No. EPA-450/1-73-002, United States
Environmental Protection Agency, Research Triangle Park, North Carolina, September
1973.
3. Kramm, D. J., Selection and Use of the Rotary Lime Kiln and Its Auxiliaries—fl. Paper
Trade Joi rnal, I 56(35):25-31, August 2], 1972.
11-11
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4. Taidor, C. E., Lune ((tins and Their Operation. Proceedings of the International
Conference on Atmospheric Emissions from Sulfate Pulping. Hendrickson, E. R. (ed.).
DeLand, Florida. E. 0. Painter Printing Company, April 28, l966. p. 244-251.
5. Prakash, C. B., and Murray, F. E., Studies on H 2 S Emission during Calcining. Pulp and
Paper Magazine of Canada, 74:99-102, May 1973.
6. Stuart, Fl. H., and Bailey, R. E., Performance Study of a Lime Kiln and Scrubber
Installation. Tappi, 48: 104A- 1 08A, May 1965.
7. Landry, J. E., and Longwcll, D. H., Advances in Air Pollution Control in the Pulp and
Paper Industry. Tappi, 48:66A-70A, June ]965.
8. Erdman, A., Fluidized Bed Burning of ((raft Mill Lime Mud. Paper Trade Journal,
I 54(36):40-42, September 7, 1970.
9. Van Donkelaar, A. J., Air Quality Control in a Bleached Kraft Mill. Pulp and Paper
Magazine of Canada, 69( [ 8):69-73, September 20, 1968.
10. I-low a Mead Kraft Mill Operates without Air Environment Problems. Paper Trade
Journal, 158:26-29, April 8, 1974.
11. Caron, A. L., Suggested Procedures for the Conduct of Lime Kiln Studies to Define
Minimum Emissions of Reduced Sulfur Through Control of Kiln and Scrubber
Operating Variables. Special Report No. 71-01, National Council of the Paper Industry
for Air and Stream Improvement, Corvallis, Oregon, January 1971.
12. Walther, J. E., and Ambcrg, H. R., Odor Control in the Kraft Industry. Chemical
Engineering Progress, 66:73-80, March 1970.
13. Collins, T. T., The Oxidation of Sulfate Black Liquor. Paper Trade Journal,
130(3):37-40, January 19, 1950.
14. Personal Communication with Dr. Hal B. H. Cooper, Texas A&M University, College
StaLion, Texas, June 1974.
15. Blosscr, R. 0., and Cooper, H. B. H., Current Practices in Thermal Oxidation of
Noncondensable Gases in the Kraft Industry. Atmospheric Pollution Technical Bulletin
No. 34, National Council of the Paper Industry for Air and Stream improvement, New
York, New York, November 1967.
11-12
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16. Calcano, S. F., and Leopold, K. M., A Survey of Emissions of Nitrogen Oxides in the
Pulp Mill. Tappi, 56(3):74-76, i 1arch 1973.
17. Personal Communication with Mr. Thomas L. Singman, Union Carbide Corporation,
Tarrytown, New York, August [ 973.
18. Personal communication with Mr. James B. Ellis, Fibrcboard Corporation, Antioch,
California, September 1973.
19. Personal communication with Mr. Andrew F. Reese, Fibreboard Corporation, Antioch,
California, March 1970.
11-13
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CHAPTER 12
SMELT DISSOLViNG TANK
Thc smelt dissolving tank is a large vessel located below the recovery furnace. A molten
mixture, primarily of sodium sulfidc and sodium carbonate (smelt), is continuously
removed from the floor of the recovery furnace. The smelt is mixed with water in the smelt
tank to produce green liquor. The smelt tank is an open, agitated vessel, covered by a hood
from which large volumes of steam arc emitted when the molten smelt and water mix. The
smelt tank, along with the recovery boiler and lime kiln, is one of the main particulate
matter sources in the Kraft pulp mill. In addition, the smelt Lank can be a source of IRS
emissions.
12.1 Smelt Dissolving Tank Particulate Matter Emissions
Particulate matter consisting of both dissolved and undissolvcd NaOH, Na 2 C0 3 , and Na 2 S is
emitted from the smelL Lank with the rising flow of gases. Table [ -6 indicates that typical
smelt dissolving tank particulate emissions arc 0.01-0.5 kg per metric ton of pulp
(0.02.1.0 lb/ton) following control devices. The majority of the smelt tanks that arc
controlled use simple mist eliminator pads to filter the particulate matter from the escaping
vent gases (1).
The mist eliminator pads consist of fine wire mesh screens, approximately 30 cm (I IL)
thick. Droplets condense from the gas on the wire mesh and arc washed back into the
dissolving tank by water sprays. As may be seen from Table 12.1, typical pad efficiency for
particulate matter removal is about 70-90%.
A higher collection efficiency can be achieved by following the mist eliminator with a spray
or packed tower scrubber. Alternatively, some mills have low pressure drop venturi
scrubbers 15-20 cm (6-8 in) of water, cyclone spray scrubbers or packed towers for
particulatc control without mist eliminator pads. One such packed tower gave 98%
collection efficiency (Table 12.1).
It is possible to combine the vent gases from the smelt tank with the main fILIe gases from
the recovery boiler prior to a recovery boiler particulate collection device. One expected
difficulty with this approach would be the effect of the watcr vapor content of the smelL
tank vent gases on recover)’ boiler particulate mattcr collection efficiency in an electrostatic
precipitator. A second potential problem would be the likelihood of H 2 S formation when
the Na 2 S entrained in the smelt tank vent gases come into contact with the recovery boiler
CO 2 .
12.1
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TABLE 12-1
SM ELT DISSOLVING TAN K PARTICULATE. MATTER
CONTROL DEVICES (1)
Collection Control
Control Device Efficiency Emission Rate
Percent kg/t (lb/ton)
Pad entrainment 71.8 0.026 (0.052)
Separator 77.2 0.075 (0.15)
77.8 0.32 (0.63)
90.2 1.2 (2.3)
93.4 0.6 (1.2)
70.8 0.79 (1.58)
Pad plus shower scrubber 96.2 0.21 (0.41)
Pad plus packed scrubber 91.9 0.60 (1.20)
Packed scrubber 98.4 0.025 (0.05)
Another combined treatment method for smelt tank vent and recovery boiler gases would
be an electrostatic precipitator and scrubber combination. As discussed in Section 10.8. I, a
tail end scrubber can offer significant benefits from a heat recovery standpoint. In cases
where “snowing” from the precipitator occurs, a tail end scrubber can also offer significant
particulate matter collection advantages. Introduction of the smelt dissolving tank vent gases
after the precipitator and prior to the scrubber is possible, with additional heat recovery
benefits.
Particulate matter control, as such, has not yet been required for smelt tanks in Scandinavia,
although some mills use indirect condensers to recover heat from the smelt tank vent gases.
The condensate is returned to the smelt tank and the warm water produced is used f6r
washing. The feasibility of designing for heat recovery in combination with particulate
control i&cnhanced by increasing fuel costs.
12.2 Smelt Dissolving Tank TRS Emissions
The presence of some reduced sulfur compounds (Na 2 S) in the smelt proper, and
occasionally some reduced sulfur gases from the recovery furnace, can cause TRS emissions
from the smelt tank vent. The amount of such gases is highly variable, reported to range
from the equivalent of 0 to 1.85 kg H 2 5 per metric ton of pulp (0.3.7 lb/ton) (1). Variables
that effect the TRS emission rate are the sulfide content of the particulate matter in the
vent gases, the turbulence in the dissolving tank, the type of solution used in a scrubber, if
present, and the pH of the scrubber liquor (2). The effect of some of these variables is
12-2
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illustrated in Tablc 12-2. This table, prepared from a NCAS1 special study in 1970-1971 (2),
indicates that thc most effective TR.S control was wet scrubbing with fresh water.
TABLE 12.2
TRS EMISSIONS FROM SMELT DISSOLVING TANKS (2)
TRS
Mill kg/t (lb/ton) Control Device Scrubbing Solution
II 0.005 0.01 None
III 0.06 0.12 Packed Tower Weak Wash and Contami-
nated
0.005 0.01 Packed Tower Fresh Water
0.005 0.01 Spray Fresh Water
IV 0.02 0.04 Showers Fresh Water
0.02 0.04 Showers Fresh Water
0.04 0.08 Demister Fresh Water
0.055 0.11 None Fresh Water
V 0.005 0.01 Demister Fresh Water
0.01 0.02 None
0.0005 0.001 Dcmister Fresh Water
0.0005 0.001 None
VI 0.005 0.01 Demister Fresh Water
0.005 0.01 None
0.0005 0.001 Demister Fresh Water
VII 0.01 0.02 Showers Fresh Water
0.015 0.3 Showers Fresh Water
VIII 0.005 0.01 Demister Contaminated Condensate
IX 0.0005 0.001 None
X 0.01 0.02 Dcmister Fresh Water
0.005 0.01 Demister Fresh Water
0.0005 0.01 Demister Fresh Water
XI 0.0005 0.00 1 Packed Tower Weak Wash
XII 0.005 0.01 Demister Weak Wash and Contami-
nated Condensate
0.005 0.0] Demister Weak Wash and Con tami-
nated Condensate
XVII 0.005 0.01 Showers Fresh Water
12-3
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Weak wash from the lime mud clarifier, lime mud washing filtrate and evaporator
condensates arc used as well as fresh watcr for scrubber liquor. increased organic sulfides
were foLind when evaporater condensates were used. A small amount of hydrogen sulfide
was liberated from lime mud clarifier supernatant in some systems, probably because of
acidification of the scrubbing liquor (3).
Recently, addition of caustic to smelt tank scrubbers has been suggested for improved TRS
control (4). in one mill, caustic will be added to clean evaporator condensates to be used as
smelt tank scrubber liquor. The scrubbing liquor will then be used for lime mud washing (5).
12.3 References
1. Atmospheric Emissions from the Pulp and [ ‘aper Manufacturing Industry. EPA
450/1.73.002. September 1973. (Also published as NCASI Technical Bulletin No. 69,
February 1974.)
2. Factor Affecting Emission of Odorous Reduced Sulfur Compounds fro,n Miscellaneous
Kraft !‘rocess Sources. NCAS1 Technical Bulletin No. 60, March 1972.
3. Blosser, R. 0. Miscellaneous Sources and Trends in Kraft Emission Control: Tappi,
55:1189-91, 1972.
4. Anon. How a Mead Kraft Mill Operates iVithout Air Environment Problems. Paper
Trade Journal, 158:26-29, April 8, 1974.
5. Testimony of W. A. Wrase, S. D. Warren Company before the Air Pollution Control
Commission, Department of Natural Resources, State of Michigan. i lay L7, 1974.
12-4
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CHAPTER [ 3
EMISSIONS OF OXIDES OF NITROGEN, HYDROCARBONS,
AND WATER VAPOR
Additional pollutants from the kraft pulp mill include oxides of nitrogen and hydrocarbons.
Oxides of nitrogen can be emitted from combustion sources, such as the recovery furnace,
the lime kiln of the chemical recovery system, and the power boilers. Hydrocarbons and
other organic nonsulfur compounds can be emitted in varying quantities from the digester,
washers, evaporators, and direct contact evaporators. Both types of pollutants may have
potential significance to photochemical air pollution whcn acting together. Water vapor is
also emitted in varying quantities from all kraft pulp mill sources. Condensation of water
vapor into a visible plume may present some hazard if the plume restricts visibility
adversely, such as across a highway, airfield, or harbor.
13.1 Nitrogen Oxides
Nitrogen oxides arc formed by the reaction of atmospheric nitrogen and oxygen at elevated
temperature when fuels arc burned. Nitrogen oxides also can form from the oxidation of
nitrogen which is present as a trace constituent in fuels. The major oxides of nitrogen
formed during combustion processes arc nitric oxide (NO) and nitrogen dioxide (NO 2 ).
Their chemistry of formation is as follows:
N 2 +02 - 2N0
2N0 +02 - 2NO 2
The nitrogen oxides present in exhaust gases from combustion processes normally arc 90 to
95 percent by volume NO, and 5 to 10 percent by volume NO 2 .
The primary variables affecting the rate and degree of formation of nitrogen oxides during
combustion processes are the flame temperature and the oxygen content of the gas in the
flame zone. The degree of nitrogen oxide formation tends to increase exponentially with
temperature above about 13000 C (24000 F), particularly when the oxygen concentration in
the combustion zone is two percent by volume or greater (I). The rate of NO formation
increases sharply with temperature, as shown in Table 13-i (2).
The major variables that can influence the nitrogen oxide formation during thermal
oxidation processes include the flame temperature, oxygen content in the flame zone, fuci
nitrogen content, and combustion unit configuration. Flame temperature is increased by
13-1
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TABLE 13-1
EFFECT OF FLAME TEMPERATURE ON NITRIC OXIDE
EQUILIBRIUM CONCENTRATION AND REACTION TIME (2)
C as*
Temperature NO Concentration Reaction Time
°C (°F) ppm, by volume sec
2430 (4400) 19,000 0.004
1930 (3500) 14,000 0.090
1760 (3200) 4,000 0.7
1430 (2600) 500 21.0
1090 (2000) 10 162.0
‘ Reactant gases. 77% N 2 , 15% 02, and 8% inerts.
increased fuel heating value, but decreases with increasing fuel moisture content. The heat
release rate and combustion volume available for heat release also influence flame
temperature. Both black liquor and lime mud tend to have moisture contents of 30 to 40
percent by weight, which act to inhibit increases in flame temperature because the water
acts as a heat sink. The fuel nitrogen content of both of these fuels is relatively low, 0.1 to
0.5 percent by weight for black liquor and negligible for lime mud.
The fuel nitrogen content is the primary variable limiting NO emissions at temperature
below 1100° C (2000° F); while flame temperature becomes dominant above 1300° C
(2400° F) (3).
Galeano and Leopold (4) report that nitrogen oxide concentrations from kraft pulp ipill
sources are relatively low when compared to power boilers and are higher from lime kilns
than from recovery furnaces, as listed in Table 13-2.
Nitrogen oxide emissions in both cases can be minimized by operating combustion units at
minimum excess air, minimum flame temperatures, and maximum fuel moisture contents.
13.1.1 Kraft Recovery Furnaces
The configuration of the combustion unit is an important factor in determining potential
nitrogen oxide emissions from kraft pulp mill combustion sources. Kraft recovery furnaces
have relatively large volume rectangular combustion chambers and relatively low heat release
rates of 170 to 340 Mi/h per m 3 (4,500 to 9,000 BTU/hr per ft 3 ), which act to inhibit high
flame temperatures. The fuel has a relatively low heating value of 14 to 16 Mi per kg dry
solids (6000-7000 BTU/lb). Black liquor has a high moisture content of 30 to 40 percent by
13.2
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TABLE 13-2
NITROGEN OXiDE EMISSiONS FROM KRAFT PULP MILL PROCESS
SOURCES (4)
NO Concentration Emission Rate
Process Unit Temperature Average Range Average Range
°C (°F) ppm, by vol. kg NO /t (lb/ton)
Recovery furnace 1010-] 230 (1850-2250) 32 0-53 3(6) 0-5 (0-10)
Lime kiln 1260-1390 (2300-2530) 200 113-260 18(36) 10-25 (20-50)
weight that inhibits flame temperature rise. Flame temperatures arc about 980 to 1260° C
(1800 to 2300° F) for normal operation in a kraft recovery furnace.
In addition, the endothcrmic reduction of Na 2 SO 4 to Na 2 S in the smelt bed of the recovery
furnace also acts as a heat sink to inhibit excessive flame temperatures. Another built-in
control method is that the air flow is split between primary and secondary zones, and
sometimes a tertiary zone. A reducing atmosphere above the smelt bed also acts to remove
oxygen from the combustion zone to inhibit the formation of nitrogen oxides. Introducing
tangential air into the furnace can act to spread out the flame front and, in that way, also
inhibit the increase in gas temperature. As a result of inhibiting both flame temperature and
oxygen level in the combustion zone, the nitrogen oxide levels normally range from 0 to 50
ppm by volume.
[ 3.1.2 Lime Kilns
Lime kilns are used to burn a mud containing CaCO 3 to recover CaO by addition of natural
gas or fuel oil. The fuel and air are added countercurrcntly with the mud at opposite ends of
a long cylindrical rotary kiln, resulting in a maximum temperature of 1260 to 1430° C
(2300 to 2600° F) at the fuel addition end of the kiln and a temperature of 200 to 315° C
(390 to 600° F) at the lime mud addition end of the kiln. The lime mud contains 30 to 40
percent water by weight. Most of this water has been evaporated by the time the mud
reaches the hot end of the kiln and does not exert as great a suppressing effect on flame
temperature as does the black liquor in a recovery furnace. The firing zone is narrower and
is concentrated with less lateral turbulence in the lime kiln, producing higher flame
temperatures. As a result, higher nitrogen oxide concentrations exist in flue gases from lime
kilns than from recovery furnaces in kraft pulp mills.
13.1.3 Power Boilers
Steam is required in pulp mills for process and space heating, driving equipment, and
generating electricity. Although significant quantities of steam are generated in recovery
13-3
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furnaces, conventional industrial boilers supply much of the steam requircd by a pulp mill.
Power boilers are the largest sources of nitrogen oxides in pulp mills. Table 13-3 lists some
values of nitrogen oxide concentrations from kraft and NSSC pulp mill power boilers.
Nitrogen oxide emission factors for auxiliary power boilers were presented in Table 1-8.
TABLE 13-3
NITROGEN OXIDE EMISSIONS FROM POWER BOILERS (4)
NO
Unit Average Range Temperature Excess Air
0 0 o
ppm, by vol. C ( F) ,
Pulverized coal, front end, 375 310-445 1425-1480 (2600-2700)
86,000 kg/h (190,000 lb/hr)
Pulverized coal & bark, 205 150-280 1200-] 425 (2200-2600)
32,000 kg/h (70,000 lb/hr)
Gas fired, 45,000 kg/h 436 325-535 33
(100,000 lb/hr)
Gas fired, 100,000 kg/h 190 161-232 870-980 (1600-1800) 46
(220,000 lb/hr)
Bark fired, ]22,000 kg/h 123 101-145 [ 040-1140 (1900-2080) 89
(270,000 lb/hr)
13.2 Water Vapor
Water vapor is emitted in varying quantities from all kraft pulp mill sources. Water vapor can
be considered as an air pollutant under ccrLain circumstances, such as when it acts to reduce
visibility in highly humid or cold atmospheres or acts to modify climate or rainfall. Water
vapor emissions also represent a potential loss of heat energy that mighL otherwise be
recovered by condensation to reduce overall plant energy consumption.
Crabtree (5) reports that the potential loss of water vapor to the atmosphere from kraft
pulp mill operations is about 5 to 8 t of water per metric ton of air dried ptilp produced (5
to 8 ton/ton). The major potential sources of water vapor released to the atmosphere arc the
recovery furnace and the paper machines, as shown in Table [ 3-4 (5).
rI he loss of water vapor from kraft pulp mill sources is affected by process operating factors,
water reuse and recycling practices, and the relative efficiencies of heat recovery systems.
The use of efficient heat recovery systems for digester blow gases and multiple-effect
evaporators tends to reduce the amount of water vapor released to the atmosphere. The use
I 3-4
-------
TABLE 13-4
WATER VAPOR EMISSIONS TO THE ATMOSPHERE FROM KRAFT PULP
MILL SOURCES (5)
Source Moisture Content Water Emission Rate
%, by vol. kg H 2 0/t (lb H 2 0/ton)
Batch digester 30.99 100-250 (200.500)
Multiple evaporator 50-99 500.1000 (1000.2000)
Recovery furnace 25.35 2200-2500 (4400-5000)
Smelt tank 35-45 150-250 (300.500)
Lime kiln 25-35 350-750 (700.1500)
Paper machines 5.15 1700-2500 (3400-5000)
Total 5000-7250 (10000-14500)
of indirect contact surface condensers also minimizes the amount of water used in the
cooling processes. The usc of cooling towers for heat dissipation tends to provide an
additional source of water vapor emissions to the atmosphere if water is recycled. Fuel
savings in reduced steam usage can range from $0.55 to $3.30 per metric ton of pulp ($0.50
to $3.00/ton) by effective use of heat transfer equipment and water reuse practices (6).
Evaporation of process fuels to high solids concentrations prior to burning results in less
water vapor emission to thci atmosphere. Evaporation of black liquor to 65 to 70 percent
solids by forced circulation evaporation can improve the heat economy of the kraft recovery
system. Concentration of lime mud to 68 to 70 percent solids prior to firing in the lime kiln
reduces the water vapor release rate from the firing zone. The kiln exhaust gas from a
scrubber must be cooled to below 65° C (1500 F) to achieve an actual reduction in water
vapor emission from the lime kiln that is sufficient to compensate for possible water
evaporation in the liquid scrubber.
The two major techniques employed, to date, for controlling local water vapor levels in
ambient air are condensation to prevent its release to the atmosphere and dispersion
resulting from high stack gas velocity or tall stacks. Dispersion of moisture plumes by means
of elevated discharges may be particularly necessary for kraft pulp mills located near
highways, airports, harbors or populated areas to alleviate potential fog. The presence of
large quantities of particulate matter in ambient air can act as condensation nuclei to
accelerate the formation of fog and inhibit its dispersal where an additional source of water
vapor already exists.
Shumas and Hansen (7) describe a system where flue gases from the recovery furnaces and
bark-fired power boilers at a 1,135 metric ton per day (1250 ton/day) kraft pulp mill arc
piped 0.8 km (0.5 mile) to a stack 62 m high (200 ft) and discharged 258 m (850 ft) above
13-5
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thc valley floor. The purpose of the stack is to discharge the moisture plume above the level
of the normal winter inversions in the narrow valley so as to alleviate any potential fog
formation problems, such as interference with aircraft landings and takeoffs. Plume rise
above the top of the stack ranges from 90 to 210 in (300 Lo 700 ft), depending on stack gas
flow rate and temperature and the season of the ‘ear.
Flue gases from the respective combustion units are passed through a venturi scrubber with
a flow rate of 510 m 3 /h (2250 gpm) and arc cooled to 50° C (125° F). The gas then flows
through a wood stove cylindrical duct 4.9 m (16 ft) in diameter at a flow rate of 620 m 3 /s
(1.3 million efm) with a fan discharge gauge pressure ranging from 6.3 to 31.8cm of water
(2.5 to 12.5 in water).
The capital cost, including all engineering fees, for the system is about S6,500,000 to
$7,000,000. The system is powered by two parallel turbine drive booster fans of 2050 kW
(2,750 hp) each. At an electric power cost of $67/kW/vear (S50/hp/ycar), Lhe operating
costs are computed as $275,000 per year for the fans.
A system recently was placed in operation at a kraft pulp mill to alleviate moisture plume
problems from three parallel paper machine drier vents near a major interstate highway (8).
The exhaust gases total 230,000 m 3 /hr (135,000 efm) with a temperature of 46° C (115° F)
and a moisture content of JO to 15 percent by volume. The gases are collected and passed
through a central fan of about 1 12 to 150 kW (150 to 200 hp), then through a cylindrical
metal stack 25 m high (80 ft) with an exit diameter of approximately 2 m (6 It), resulting in
an exit gas velocity of 22.30 rn/s (75-100 Ips). l’hc capital cost for the system is about
$80,000; the annual direct operating cost approximately $10,000, based on $67/kW/year
($50/hp/year).
13.3 Organic Compounds
Organic compounds other than those containing sulfur are also emitted in varying quantities
from several types of processing units in kraft pulp mills. For a discussion of nonsulfurous
organic compounds, sec section [ .1.3.
13.4 References
I. Smith, W. S., and Grubcr, C. W., Atmospheric Emissions from Coal Combustion: An
Inventory Guide. U.S. Public Health Service Publication No. 999-AP-24, Cincinnati,
Ohio, 1966.
2. Ermenc, E. D., Controlling Nitric Oxides Emissions. Chemical Engineering,
77(l2):103- 105, June I, [ 970.
13-6
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3. Bartok, \V., Crawford, A. R., and Skipp, A., Control of Nitrogen Oxide Emissions from
Stationary Combustion Sources. Combustion, 42(4):37.40, October 1970.
4. Galeano, S. F., and Leopold, K. M., A Survey of Emissions of Nitrogen Oxides in the
Pulp Mill. Tappi, 56:74.75, March 1973.
5. Crabtrce, V. F., Abatement Procedures Presently in Use or Feasible: Other Operational
Sources. In: Proceedings of the International Conference on Atmospheric Emissions
from Sulfate Pulping, April 28, [ 966, Sanibcl Island, Florida, Hendricksen, E. R. (ed.).
l)eland, Florida, E. 0. PainLer Printing Co., 1966. p. 252.264.
6. Tomlinson, G. I-I., Science of Wood Pulping. Tappi, 44:133A-]42A, January 1961.
7. Shumas, F. J., and Hansen, C. A., A Unique Solution to Punching through the
Inversion Layer. (Presented at 1973 Tappi Environmental Conference. San Francisco.
May 15, 1973).
8. Personal communication with Mr. David E. Mansfield, Western Kraft Corporation,
1973.
13.7
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CUAP’l’ER. 14
AIR POLLUTION CONTROL IN SULFITE PULP MILLS
Even though many technological similarities exist bctwcen Lhc sulfite and the sulfate or
kraft process, the air pollutants generated in each of these processes are quite different. The
sulfite process usually operates with acidic SO 2 solutions and, therefore, SO 2 is the main
gaseous air pollutant. In certain special cases of alkaline sulfite liquor buriiing iti recovery
boilers where chemical reduction may take place, H 2 S may also be emitted. Organic reduced
sulfur compounds arc not produced. Since the odor threshold is roughly one thousand times
higher for SO 2 than for reduced sulfur compounds, odors generated in the sulfite pulp mill
will generally be much less than in the kraft process. The sulfite process will also emit
particulate matter from spent liquor burning.
Even within the sulfite industry itself, there are many differences in SO 2 and particulate
emissions because of differences in cooking liquor bases, acidities, and recovery methods.
14.1 Sulfite Pulping Processes
14.I.l General
Sulfite pulping is practiced with numerous modifications. Cooking can be (lone with
different bases, namely calcium (Ca), magnesium (Mg), ammonium (NI - I 4 ), and sodium (Na).
Cooking can be done:
1. At different pH levels, from acidic to alkaline, subject to the solubility constraints
imposed by the base used;
2. In several stages with changes in cooking liquor between;
3. With a high yield of pulp and with additional mechanical defibration (the neutral
sulfite semi-chemical processes).
The main sulfite pulping processes, pll range, principal cooking liquor composition, and
bases are presented in Table 14-I. This table also lists the allowable pH range for the
different bases.
The potential for release of SO 2 into the atmosphere increases with decreasing pH.
14-1
-------
TABLE 14-1
MAIN SULFITE PULPING PROCESSES
Cooking Method
Acid bisulfite
Bisulfite
Ncutral sulfite
Alkaline sulfite
Two-stage
‘I’wo -stage
Three-stage
pH Range
I -2
2-6
6-9
10-12
6-8 and 1-4
2 and 5-8
4 and 2 and 8
Cooking Liquor
M(HSO 3 ) 2 , SO 2
M(HSO 3 )2
MSO 3 , MCO 3
MSO 3 , M(OH) 2 , J lS
MSO 3 , M(HSO 3 )2, SO 2
M(HSO 3 ) 2 , SO 2 , MSO 3
M(HSO 3 ) 2 , SO 2 , MSO
Base Alternatives
M = Ca, Mg, (NH 4 ) 2 , Na 2
M = Mg, (NH 4 ) 2 , Na 2
isi (NI- 1 4 ) 2 ,Na 2
M Na 2
M = Na 2
M = Na 2
M = Na 2
The cost of chemicals and the possibility of chemical recovery provide more information on
the emission potentials of the different sulfite processes. Based on different chemical prices,
combustion products and recoveries of chemicals, calcium-based processes have the highest
air pollution potential, ammonium-based processes next, and magnesium- and sodium-based
processes are the least polluting, see Table 14-2.
TABLE 14-2
SULFITE PROCESS CHEMICALS, PRICE, COMBUSTION AND RECOVERY
1/7
2/3
3/4
*Spent Sulfite Liquor.
SO 2
SO 2
SO 2 , N 2 , NO
SO 2 ,H 2 S
None
MgO+SO 2
SO 2
Na 2 S, NaCO 3
Many sulfite mills continue to dispose of their spent sulfite liquor (SSL) to receiving waters.
These arc primarily calcium-based mills, but some magnesium-based mills and a number of
NSSC mills using ammonium or sodium practice this disposal technique. These mills are
usually old and though they will have less SO 2 emission than similar mills with SSL
recovery, their water pollution contribution through SSL dumping is so great that major
Relative
Price
Base Per Mole
Ca
Mg
NH 4
Na
pH
Range
1-2
1-6
1-9
1-12
Gas
Products
SSL* Combustion
Dust or Smelt
CaSO 4 , CaO, CaCO 3
MgO
None
Na 2 S, Na 2 CO 3
Chemicals
Feasibly
Recovered
14-2
-------
mill modifications, or in some cases, closures, may be required Lo comply with water
pollution regulations.
14.1 .2 Atmospheric Emissions from Different Sulfite Processes
Typical SO 2 and dust emissions for different Scandinavian sulfite processes are given in
Table 14-3. Examples of the typical SO 2 and particulate emissions from United States
sulfite pti 1 p mills arc presented in Table 14-4.
TABLE 14-3
SCANDINAVIAN SULFITE PULP MILL EMISSIONS (1)
Mill Number
Parameter 1 2 3 4 5 6 7 8 9
Base Ca Ca Mg Mg Na Na Na NH 4 NI -I 4
Cooking pi-l 1-2 1-2 1-2 3-5 3-5 3-5 3-5 3-5 3-5
Cooking yield, % 50 50 55 55 50 75 75 75 75
SSLcoLlection,% 0 70 85 85 90 0 85 0 80
SSL evaporation No Yes Yes Yes Yes No Yes No Yes
SSL combustion No Yes Yes Yes Yes No Yes No Yes
Chemicals recovery No No Yes Yes Yes No Yes No No
Sulfur losses,
kgSO 2 It
Washing loss 210 56 20 40 IS 100 16 100 16
Spills 0 20 10 10 10 0 10 0 10
Condensates 0 26 15 hO 10 0 8 0 8
Gaseous loss [ 0 10 10 2 2 2 2 2 2
Flue gases 0 83 15 25 20 0 14 0 74
Flue gas dust 0 25 0 0 0 0 0 0 0
Sulfur make-up, 220 220 70 87 57 100 50 100 100
kgSO 2 /t
Total SO 2 emission, 10 93 25 27 22 2 16 2 76
kg SO 2 /t
Flue gas dust, 0 90 70 100 20 0 10 0 0
kgdust/t
Dust collection, % 0 80 99 99 97 0 97 0 0
Dust emission, 0 18 1 1 1 0 0 0 0
kg dust/t
14-3
-------
TABLE 14-4
TYPiCAL POTENTIAL AMERiCAN SULFITE PULP MILL EMISSIONS (2)
Mill Number
Parameter 1 2 3 4
Base Ca Mg Mg NH 4
Type acid sulfite acid sulfite bisulfite acid sulfite
Cooking yield, % 45 43 50 50
SSL evaporation No Yes Yes Yes
SSL combustion No Yes Yes Yes
Chemicals recovery No Yes Yes No
SO 2 emissions, kg S0 2 /t
Digester blow 75 8.5 8.5 1.5
Washers and screens 2 2 2
Evaporators 2.5 2.5 5
Boiler 9.5 10 200
Acid towers 2.5 — — 0.25
Particulate matter emissions, kg/t 0 1 ] 1.5
*Packcd tower absorber with 91% SO 2 removal.
*#Venturi absorber after recovery boiler.
In the following sections, the air pollutants generated in the basic sulfite process operations,
from digestion to cooking acid preparation, arc discussed, and the effects of different bases
and acidities arc treated separately for each basic operation.
The general treatment for SO 2 emissions is scrubbing with alkaline solutions. The recovery
of SO 2 is a function of pu and gas film resistance, and usually exceeds 90 percent.
Different dust separation methods can be used for collecting particulate matter. The three
most widely used arc cyclones, scrubbers and electrostatic precipitators.
14.2 Digester Gases
Sulfite cooking is performed both in batch digesters and in continuous digcsters, the former
being more prevalent. The continuous digester causes no air pollution problem; the batch
digester may release substantial amounts of SO 2 depending on the blow system and cooking
acid ph.
14-4
-------
14.2. 1 Batch Digesters
Acid bisulfite cooking in batch digesters creates high SO 2 pressures in the digester; bisulfite
and neutral sulfite cooking have correspondingly lower pressures. Digester relief and blow
cause 502 emissions.
502 and cooking liquor are relieved from the digester to the acid preparation accumulator
system in several stcps, top relief, side relief, high-pressure blowdown, and, finally,
low-pressure blowdown. All the SO 2 relief is recovered in the acid preparation system (section
14.7). For bisulfite and neutral sulfite cooking, the 502 pressure in the digester is lower and
a less elaborate relief system will return the 502 to acid preparation. The digester relief will,
therefore, usually not constitute an air pollution problem for any base or for any pH range.
The digester can be emptied in two ways. The usual North American practice is to relieve
the digester down to a certain blow pressure, (.14..28) MPa (20-40 psig), and them blow the
contents into the blow pit. This procedure is called hot blow (Figure 14-I).
The European practice is to relieve the digester down to atmospheric pressure and then flush
Lhe contents into the blow pit by pumping spent liquor into the digester. This procedure is
called cold blow.
FloL blow will release considerable amounts of 502 when the spent liquor is flashed,
especially for acid bisulfite cooking. The 502 emission from the blow pit can amount to
about 30 kg SO 2 per metric ton of pulp (60 lb SO 2 /ton) (2, 3). The blow gas is roughly 95
percent water vapor, 3 percent SO 2 , and 2 percent CO 2 (2).
One feasible treatment for the SO 2 emission is to scrub the blow gas with an alkaline
solution of the base and return the solution to acid preparation. SO 2 recovery efficiencies
are reported to be as high as 97 percent for this type of treaLmeuL (3). This method works
well with Na and NH 4 bases. Magnesium and calcium require cumbersome slurry scrubber
systems.
Another method is to scrub the blow gases with cold water to recover heat and SO 2 ; the
502 is then steam stripped from the water solution. This method is reported as being
insufficient for 502 recovery (3).
Hot blow also may be performed in ordinary blow tanks with appropriate heat recovery
systems. This is usually practiced with bisulfite or neutral sulfite cooking, either of which
has much less potential for SO 2 release.
l3low pit emissions for bisulfite or neutral sulfite cooking are about 10 kg per metric ton of
pulp (20 lb S0 2 /ton) (2). Cold blow will show considerably smaller SO 2 release than hot
14-5
-------
D
0
I—s
10
1r
0
m
C l )
-I
P1
C l )
H
z
0
r
0
-u
-l
I a )
H
PRINCIPAL FLOW DIAGRAM FOR SPENT LIQUOR COLLECTION IN BLOW PITS, 2 STAGE
DISPLACEMENT WASH
FIGURE 14-1
-------
blow. With acid bisullitc cooking, the SO 2 emission from the blow pit with cold blow will
be 2- 10 kg SO 2 iier metric ton of pulp (4-20 lb SO 2 /ton).
Again, tim most practicable treatment is scrubbing with alkaline solution of the base.
Calcium and magnesium base are not generally suitable for this application.
With bisulfite or neutral sulfite cooking, the cold blow pit 502 emission is less than 2 kg
SO 2 per metric ton of pulp (4 lb/ton), making scrubbing impractical.
14.2.2 Continuous Digestcrs
The continuous digester is essentially a completely closed system cooking with magnesium,
ammonium, or sodium base in the pl-I range 4-6. As with a kraft continuous digester, a
continuous diffusion washing stage is intcgratcd with the digester. These features make the
continuous sulfite digestion process a very clean one, with negligible emissions of SO 2 . The
502 containing gases from the presteaming vessel relief and from the flash evaporators are
returned to the cooking acid preparation system.
14.3 Washer Gases
There are many different sulfite pulp washing systems. For example, pulp may be washed
by displacement in the digester or in the blow pit. Additional washing in rotary drum filters
or in continuous diffusers is usually necessary (Figure 14.2) because of current water
pollution regulations for SSL collection efficiency and the high cost of the sodium- and
magnesium-based chemicals. ‘I’ypically, the SO 2 emissions from washers and screens are 8 kg
SO 2 per metric ton of pulp (16 lb S0 2 /ton) for bisulfite cooking and even less for hisulfite
and neutral sLillite cooking. Because of the large gas flows involved in vacuum rotary drum
washing, 502 recovery through scrubbing is not practicable. For washing methods with
smaller vent gas flows, scrubbing may be used.
14.4 Evaporator Gases
The evaporation of SSL results in the release of SO 2 from the liquor. The more acidic the
cooking liquor, the more 502 is liberated during SSL evaporation. The evaporation systems
in use are the same as those used for kraft black liquor, and the multiple-effect vacuum
evaporation plant is the most commonly used for both. Each evaporation effect is vented to
the vacuum system, and the vacuum system vent becomes the main emission point for
evaporator gases. The hotwell is the main emission point for evaporator condensates. These
condensates contain SO 2 and can contribute to the evaporation plant SO 2 emission (Figure
14-3).
14-7
-------
VENT
PRINCIPAL FLOW DIAGRAM FOR SPENT LIQUOR COLLECTION IN ROTARY WASHER PLANT
(3 2-ZONE FILTERS)
PULP
H
II BLOW
TANK
F LTRATE
TANKS
FIGURE 14-2
-------
EFFE
CT NO.1
NO.2 NO.3 NO.4 NO.5
(BODY
AandB)
(BODY
AandB)
SUR FACE
CONDENSER
PRINCIPAL FLOW DIAGRAM FOR SPENT LIQUOR EVAPORATION IN MULTIPLE EFFECT VACUUM PLANT,
5 EFFECT, 7 BODIES (SPARE BODIES NOT SHOWN)
FIGURE 14-3
-------
For bisulfite nd neutral sulfite cooking, the evaporator gases contain small amounts of
SO 2 usually less than I kg SO 2 per metric ton of pulp (2 lb S0 2 /ton). With acid hisulfite
cooking, the evaporator gases contain 20-30 kg SO 2 per metric ton of pulp (40-60 lb
S0 2 /ton) (I). This significant SO 2 emission can be treated and recovered by several
methods. The feasibility of each method depends on what other results can be achieved
simultaneously by that particular method.
A widely 1 )racticed method of eliminating evaporator SO 2 emissions is to return the
evaporator gases to the acid preparation plant to recover the SO 2 .
Another method is to scrub the gases with an alkaline solution of the base and then return
the solution to acid preparation. This procedure, however, requires a greater investment and
is difficult with calcium and, to a lesser degree, with a magnesium base.
If the sulfite mill has a weak liquor fermentation plant, either for CFI 3 CH 2 OU or yeast, the
SSL must be neutralized prior to its use in fermentation. This can be done by stripping it in
a separate evaporation stage or by adding base. After fermentation, the SSL can be
evaporated with insignificant SO 2 emissions.
The weak SSL can be steam stripped in the evaporation plant, as mentioned before. The
SO 2 stripped off is returned to acid preparation, and subsequent evaporation of the stripped
SSL will yield very little SO 2 .
The weak SSL can be neutralized by adding base, CaO, MgO, ammonia (N H 3 ), or NaOll, to
the liquor before evaporation. This practically eliminates gaseous SO 2 emissions from the
evaporation plant. SSL neutralization, however, is mainly proposed to reduce the
evaporation condensate BOD by hydrolyzing acetic acid (Cl-I 3 COOH) so that the resulting
nonvolatile acetate will proceed with the liquor to combustion. Full-scale neutralization
trials began in Sweden in 1972, but, thus far, continuous operation over long periods of
time has been difficult to maintain. Acid bisulfite cooking with magnesium base requires the
addition of about 35 kg MgO per metric ton of pulp (70 lb MgO/ton) to the SSL before
evaporation to raise the pI- I to 6.5 (4).
14.5 Combustion Gases
All sulfite process spent liquors can be burned. Calcmm, magnesium and ammonium-base
SSL can be burned about any supporting combustion fuel if cvaporaLcd to 55 percent dry
solids, atomized thoroughly, and burned with preheated air in a separate combustion
chamber ahead of the main furnace. During combustion of any SSL, emissions of SO 2 and
particulates of base oxides, carbonates, and sulfates can occur, depending on the particular
prevailing equilibrium conditions.
14-10
-------
Sodium-based SSL is subject to either oxidizing or reducing combustion. in the latter case
combustion is in a black liquor recovery boiler and 112 S emissions arc possible. The majority
of particulate matter emissions will be Na 2 SO 4 . Most of Lhc sodium and sulfur is recovered
in Lhc smelt.
14.5.1 Calcium SSL Combustion Products
Collecting calcium SSL with 70 pcrccnt efficiency and burning it produces about 80 kg SO 2
and 90 kg dust per metric ton of pulp (160 lb and 180 lb/ton) (Table 14-3). The dust is
about 60 percent CaSO 4 and 40 percent CaO, which includes some CaCO 3 and traces of
other salts.
rl hc flue gases arc usually passed through a cyclone dust separation system with
approximately 80 percent collection efficiency. The dust is stored in a pile or dumped into
receiving water, because no economic usc has been found for it so far (5). Better removal
can be achieved by electrostatic prccipitators, but they arc too expensive for recovering a
virtually worthless dust.
An SSL collection efficiency of 70 percent is very low, and water pollution abatement
requirements may demand a smaller washing loss. Dumping calcium sulfite ash into the
water also may not be permitted.
SO 2 can be removed from the flue gases by scrubbing them with CaQ or CaCO 3 slurry. But
even if all the CaO of the dust from a 100 percent flue gas dust separation could be used for
this, the SO 2 emission would still lie around 40 kg/t (80 lb SO 2 /ton). A byproduct of this
control method is about 250 kg per metric ton of pulp of a 50 percent sludge (500 lb of
sludge/ton), for which no use is known.
More effective SO 2 removal requires additional lime and pro(ILices more sludge. At this
point, changing the base might be a more feasible alternative.
14.5.2 Magnesium SSL CombListion Products
Collecting magnesium SSL with 85 percent efficiency and burning it produces around 15 kg
S0 2 /t plus 70 kg magnesium oxide (MgO)/t for acid bisulfitc cooking and 25 kg 50 2 /t plus
100 kg MgO/t for bisulfite cooking (30 plus 140 lb/ton and 50 plus 200 lb/ton respectively).
These figures assume MgO recirculation (Table 14-3).
These emissions, however, arc effectively treated in a chemicals recovery system that
separates the MgO dust by cyclones and scrubbers, hydrolyzes the MgO into magnesium
hydroxide slurry, and scrLibs the SO 2 from the flue gases with this slurry. The net result is a
14-1 1
-------
small dust emission and moderate 502 emission, which can be further decreased by
additional scrubbing stages as discussed in section 14.9.1.
L4.5.3 Ammonium SSL Combustion Products
Collecting ammonium SSL with 85 percent efficiency and burning it produces around
130 kg S0 2 /t (260 lb S0 2 /ton) for bisulfite cooking. During combustion, the NI-I 3 is
converted to water and nitrogen (N 2 ), and consequently the base is lost. erubbing 502
from the flue gas with fresh ammonia solution has been successfully demonstrated in at least
two American ammonium-based mills (6, 7). The nature of any air pollution problems
reLated to the nitrogen compounds emitted from ammonium-based pulping is discussed in
section 14.10.
14.5.4 Sodium SSL Combustion Products
Sodium SSL burning may yield various combustion products depending primarily on the
method of burning. Usually the SSL is burned in a reductive recovery boiler similar to a
black liquor recovery boiler. Most of the inorganic dry solids arc then transformed into a
smelt of primarily Na 2 S and Na 2 CO 3 . Part of the inorganic dry solids arc entrained in the
flue gases as Na 2 SO 4 . It, however, is efficiently recovered by electrostatic precipitators and
returned to the furnace to be reduced to Na 2 S. The dust emission, therefore, stays low and
usually it is about I kg Na 2 SO 4 /t (2 lb Na 2 SO4 /ton). The SO 2 emission may vary widely
and is a function of the SSL sulfidity (5).
The treatment of the emissions from sodium SSL combustion depends on the particular
chemicals recovery system chosen. At present there arc 10 major sodium SSL chemical
recovery systems with additional ‘customizcd” variations (IPC, Mead, Stora, Sivola, Western
Precipitation, AST, SCA-Billerud, Tampclla, ITT-Rayonier, and cross-recovery with a kraft
mill). Both the CO 2 and SO 2 in the flue gases may be used, and depending on the system,
502 emissions will be about 6-20 kg S0 2 /t (12.40 lb 50 2 /ton).
Combustion under reducing conditions similar to the conditions in a black liquor recovery
boiler, especially to those in a pyrolysis reactor (AST and SCA Bilicrud), will also produce
1 I2S. (l-l 2 S generation and abatement in a black liquor recovery boiler were discussed in
section 10.2.) 1-125 from a pyrolysis reactor is usually converted to sulfur in a Claus reactor
or oxidized to SO 2 in a separate furnace. H 2 5 emissions are negligible.
14.6 Acid Preparation Gases
Preparation of the cooking acid occurs in two steps, preparation of the raw acid and
eventual fortification to cooking acid strength. With chemical recovery, raw acid preparation
is simultaneously the recovery process. Acid preparation gases pose a minor 502 emission
problem.
14-12
-------
14.6.1 Calcium Cooking Acid Preparation
With calcium base, SO 2 derived from burning of sulfur or pyrite is passed eountereurrently
to water in towers packed with limestone (Figure 14-4). The vent to the atmosphere will
emit some 502, but this emission is rather minor and usually amounts to less than I kg
SO 2 It (2 lb S0 2 /ton). Typical values are 0.2 kg S0 2 /t (0.4 lb S0 2 /ton) with a 20° C
(68° F) water temperature (1, 8). Higher water temperatures might increase the emission of
SO 2 by decreasing absorption efficiency.
The raw acid is fortified to cooking acid strength with digester relief and blowdown; the
vent from the acid fortification system is connected to the acid tower (Figure 14-5).
14.6.2 Magnesium Cooking Acid Preparation
Magnesium raw acid is prepared in the chemicals recovery system (Figure 14-6). The raw
acid may be used directly for bisulfite cooking or fortified for acid bisulfite cooking (Figure
14-5). In both eases, the only emissions will be with the flue gases from the chemicals
recovery system. About 15 to 25 kg S0 2 1t (30 to 50 lb S0 2 /ton) will be emitted. This
amount can be decreased by additional recovery stages.
LIMESTONE
WATER
RAW ACID
VENTS FROM
EVAPS & ACID
FORTIFICATION
FIGURE 14-4
FLOW SHEET FOR CALCIUM BASE RAW ACID PREPARATION
VENT
WATER
WAS H
TOWER
14-13
-------
FIGURE 14-5
Vent To Acid Tower
Or Recovery System
Low - Pressure Blowdown
High- Pressure Blowdown
Cooking Acid
Side Relief Acid
Raw Acid
ACID BISULFITE FORTIFICATION SYSTEM FLOW SHEET
-------
WATER
VENTS FROM
EVAPS.
FLUE GAS
1% SO 2
MAKE-UP
SO 2
RAW ACID 10
FILTERS S
EVENTUAL
FORTI Fl CATION
FIGURE 14-6
RETURN
Mg(OH) 2
SLURR’(
FLOW SHEET FOR MAGNESIUM BASE RAW ACID PREPARATION
-------
14.6.3 Ammoniiim Cooking Acid Preparation
Ammonium cooking acid preparation systems are similar to those described above.
Sulfur dioxide, either from sulfur burning or from flue gases, is absorbed in an NH 3 solution.
Depending on pH, there is a potential for NH 3 emissions. Performance runs with spray
tower absorbers using liquid circulation have shown that when the pH dropped below 6,
almost 10 percent of the SO 2 feed was lost through the tower vent, while NH 3 losses were
negligible. When the circulation pH rose above 7.6, NH 3 losses were substantial, but SO 2
losses went down to 0.3-0.6 percent. A pl-l of 7.1 kept both SO 2 and NI-I 3 losses below I
percent, but caused a large plume from the absorption Lower vent (1). Fine pH control is
required to maintain SO 2 and NI-I 3 emissions at acceptable minimums.
14.6.4 Sodium Cooking Acid Preparation
There are as many systems for sodium cooking acid preparation as there are for sodium
recovery. These recovery systems are covered in section 14.7.
14.7 SSL Recovery Boilers
14.7. L Design Parameters
Spent sulfite liquor is, as a rule, incinerated to recover both heat and cooking chemicals in
reusable form. Whenever chemicals are recovered, their recovery is considered the primary
purpose of the incineration process, especially for magnesium- and sodium-based processes
where the base is recovered. In the calcium-based process, recovery of the base is not
feasible, and in the ammonium-based processes, only SO 2 can be recovered.
Since sodium bisulfitc recovery was previously discussed under kraft processing, magnesia
(MgO) recovery is discussed in this section.
There are a variety of SSL combustion systems. To help in evaluating diffcrent systems, a
short survey of combustion fundamentals and requirements will be presented.
Burning intensity, which is the heat input rate per unit of furnace volume, depends mainly
on temperature. The theoretical combustion temperature is a function of SSL heat value,
combustion air temperature, and fuel-air ratio. The actual combustion temperature depends,
in addition, on the heat and mass exchange with surrounding zones.
Stable ignition requires transport of energy to the ignition zone. This is normally
accomplished by recirculation of hot furnace gases or hot air. The rccirculation can be
external or internal.
14-16
-------
Complete combustion is essential for the recovery of chemicals and heat. A necessary
condition for complete combustion is sufficient residence time in the high temperature zone
of the furnace. This requires sufficient furnace volume and controlled flow pattern.
Since magnesium sulfate (MgSO 4 ) cannot be economically converted into magnesium
bisulfite, a low sulfate content of the combustion product is important relative to recovery.
The formation of sulfate depends on factors, such as combustion excess air, combustion
temperature according to the chemical equilibrium, the residence time of the ash in
intermediate temperature zones, and different catalysts. Not all sulfate formation
mechanisms are fully understood.
Additional design parameters are set by requirements of boiler availability and
controllability. Therefore, the most important factors become heat surface fouling and
partial load combustion control.
14.7.2 SSL Recovery Boiler Systems
Burning of SSL is usually carried out in power-type boilers with some modifications.
Ash-free auxiliary fuels can normally be burned in the same furnace or boiler without great
disadvantage, if the auxiliary fuel input is not too high when compared to the SSL input.
Possible disadvantages of using auxiliary fuel are dilution of the gases to the SO 2 recovery
system and an increased percentage of sulfate due to temperature conditions and catalytic
conversion. SSL boilers commonly have a low furnace heat release rate, which requires a low
furnace exit temperature, widely spaced heat exchange surfaces, and effective heat surface
cleaning equipment. Similar features exist in successful kraft recovery boiler designs.
SSL burning in a fluidized bed reactor was practiced, to some extent, for all soluble bases.
This system is normally supplied as an integrated recovery unit and not in combination with
a power boiler. Environmentally, it does not differ very much from a conventional recovery
unit. The actual burning of SSL can be carried out in a variety of different burning systems.
The most important systems are:
1. Small precombustion chambers, called Loddby furnaces;
2. Large furnaces, called Lurgi-Lenzing-Steinmuller (LLS);
3. Small furnaces of Babcock & Wilcox (B&W) for liquor burning;
4. Furnaces, designed by Tampella, that incorporate features of both LLS and
Loddby furnaces; and
5. Fluidized bed reactors by Copeland and Dorr-Oliver.
14-17
-------
Loddby furnaces arc small precombustion chambers [ hat are usually horizontal and arc
attached to a main boiler furnace. The refractory walls of the cylindrical chambcr heat the
tangentially-introduced combustion air to a high tcmpcraturc. External rccirculation
stal)ilizes the flame in the muffle. The burnout of the flame occurs in the main fLirnace. The
flame temperature at the muffle exit is approximately 175° C (3500 F) below the
theoretical combustion temperature. The dust from a Loddby furnace usually has a smaller
mean particle size than dust from most other combustion systems. Loddbv furnaces have
good controllability and, if a few of these chambers arc used jointly, they are suitable for
partial load operations. in some Scandinavian mills, SSL is considered a good pressure
control fuel for recovcry.power boilers.
Lurgi-Lcnzing-Steinmuller (LLS) furnaces arc large units that are integrated with a steam
boiler main furnace, but are separated by a verticle membrane tLibe wall that has a screen
passage in the lowest part. Excess cooking of the furnace is prevented by covering the tube
walls with a thin refractory layer. The LLS furnace exit temperature is about 305° C
(550° F) below the theoretical combustion temperature. Effective ignition and mixing is
accomplished by spraying the SSL countercurrently to the air flow.
Babcock & Wilcox furnaces for liquor burning are uncooled or slightly cooled. They can be
similar to the LLS furnace. Unlike other systems, B&W uses self.stabilizing liquor burners
that have their own air registers. Auxiliary fuel is burned in the same or in different burners.
The controllability of a B&W system is good.
Tampella furnaces are combinations of LLS furnaces and Loddby furnaces. Only Ca-based
liquor has been fired thus far.
Fluidized bed reactors, by Copeland and Dorr.Oliver, are used to some extent for SSL
treatment. They provide an environment for combustion that gives a high intensity of
reaction at comparatively low tern peratures.
14.7.3 Operation Parameters
Because the recovery boiler is normally followed by chemical recovery, boiler operation has
a greater direct influence on the chemical cycle than on the gas and solids emission from the
plant.
Excess air and the use of auxiliary fuel, for an existing installation, arc the most important
operating parameters. The firing rate is also important, although large variations are not
normally expected.
For calcium-base SSL without recovery, the excess air has little or no influence on the
amount of SO 2 or SO 3 emitted from the boiler. High excess air can somewhat improve the
14-18
-------
formation of sulfate and, thus, reduce the amount of SO 2 and SO 3 . Even a complete
sulfatizing of the base will bind only about 65 percent of the sulfur. A normal value is about
35 percent. Using too much excess air can make heat recovery uneconomical.
The dust content of the flue gases from calcium-base SSL firing is often approximately
11.5 g/rn 3 (5 gr/cu it) when continuous sootblowing is practiced. With intermittent
sootblowing, the concentration can be as low as 8-9.2 g/m 3 (3.5.4 gr/cu it); during
sootblowing, it is correspondingly higher.
l’hc use of auxiliary fuel dilutes the flue gases to lower concentrations, but increases the
costs of dust separation and eventual SO 2 removal. The usc of wood refuse as fuel in the
same boiler improves the binding of SO 2 and SO 3 to the ash.
Variation of excess air to Mg-base SSL liquor firing changes the sulfate content of the ash.
Low excess air should be used to avoid sulfate formation so that all sulfur is converted Lo
SO 2 in the flue gas for more efficient recovery, in practice, a part of the Mg, usually less
than 10 percent, leaves the boiler as sulfaLe with a correspondingly slight decrease in the
concentration of SO 2 .
Only natural gas and oil can be used as auxiliary fuels. Even a high sulfur heating oil dilutes
the flue gas SO 2 concentration, iiicl no appreciable extra SO 2 absorption is expected from
oil use. An excessive use of auxiliary fuel may increase the sulfatizing of Lhe base or the
formation of weakly soluble corn binations.
Soot formation must he avoided, especially when burning auxilliary fuels. The soot will
discolor the MgO ash and may also discolor Lhe cooking acid made from the ash. Ultimately,
the soot may discolor the pulp.
The operating parameters seem to have a great influence on the results of firing
ammonia-based SSL. I)ilferent problems were reported, such as plLiggage of boilers, boiler
corrosion, and the emission of smoke, often referred to as blue haze. According to one mill
plant, by installing and using glass fiber-bed demisters on an SO 2 absorption system (6), the
blue haze is eliminated.
In tests by EKONO, the pl- I of the ash correlates strongly with the excess air of liquor
burning. This correlation is probably caused by formation of SO 3 during the combustion.
By combining ammonia-base SSL and bark-firing, an alkaline ash reaction with a low excess
air is possible. Where sulfur recovery is not feasible, an alkaline ash is the only way to
prevent SO 3 formation and subsequent direct discharges of it into the atmosphere.
14-19
-------
14.8 SSL Recovery Systems
The combustion gases for bases that are feasibly recoverable are usually treated in a
chemical recovery system. The efficiency of the recovery process is also an indication of the
amount emissions. Because of high costs, SO 2 recovery has been limited to 50-80 percent
and dust separation to 95-99 percent, both depending on the base and the recovery system.
These recovery efficiencies, however, can be increased with higher investment. l)ust
emissions below I kg/t and SO 2 emissions below 10 kg S0 2 /t (2 lb and 20 lb/ton,
respectively) have bccn adopted as emission regulations for new sulfite mills (9, [ 0). These
regulations will effectively phase out calcium-based sulfite mills in the future and will
require more efficient recovery systems for the other bases.
All sodium recovery systems that. displace H 2 S from dissolved sulfide solutions with flue gas
CO 2 or sodium bicarbonate (NaHCO 3 ) will have I I2 S as an intermediate gas in the system.
The l-l 2 S is usually converted to sulfur in either a Claus reactor or oxidized to SO 2 in a
furnace with almost 100 percent efficiency in both cases. Under normal circumstances, H 2 S
is not emitted into the atmosphere.
The recovery of soluble base chemicals and SO 2 is carried out using a few well known
chemical reactions. Nevertheless, the number of systems offered is large. Also some planis
use their own methods, which have not yet been described in the literature.
14.8.1 Magnesium Base
The basic reaction in all systems is an absorption of flue gas SO 2 into Mg(OH) 2 to produce a
bisulfite solution.
In the B&W process (see Figure 14.7), the recovered MgO together with the make-up MgO is
slaked with steam addition. The hydroxide is circulated in absorption venturis in
countercurrent flow. The hydroxide reacts with recirculated Mg(HSO 3 ) 2 to form MgSO 3 .
In the recirculating acid, the sulfite reacts with SO 2 to form new Mg(l-lSO 3 )2. The extent of
SO 2 absorption in each stage is governed by acid concentration, SO 2 concentration in the
gas, temperature, recirculation rate, and gas velocity. The use of venturi absorbers has made
much higher monosulfite concentration possible than produced with packed towers.
Combustion gases discharged to the recovery system stack can be expected to contain less
than 250 ppm SO 2 by volume.
The system as practiced in Veitsiluoto, Finland, has no dust separation and no separate
slaking system. The initial circulation solution is formed in the first tower, somewhat
simplifying the system. The design of the dList absorption system depends on close
operational tolerances, but seems to work satisfactorily.
14-20
-------
BABCOCK & WILCOX PROCESS FOR MAGNESIA BASE SULFITE CHEMICALS RECOVERY
FIGURE 14-7
-------
Systems with only one absorption unit have been offered by several manufacturers, among
them Ahlstrom, Svcnska Flaktfabriken, and Tampella. The reactors used are more complex.
but can be considered as multistage reactors. Marbles have been mtrodLiced to increase the
reaction surface, without significant effect.
In most cases, some heat is recovered from the gases in a cooling venturi before the
absorption units.
The recovery plant stack is practically the only emission point. Excessive emissions of SO 2
usually occur only when the pH control of the absorption units (toes not work properly.
High MgO emissions are normally the result of separator malfunction at plant overload.
14.8.2 Sodium Base
All recovery methods produce a solution of Na 2 SO 3 or Na 2 C0 3 , or both, with the least
possible concentration of undesired sulfur combinations. To obtain this result, the sulfur
combinations are removed by carbonization with CO 2 . The carbonate or bicarbonate is then
converted to sulfite by sulfitizing with SO 2 . In some processes the sulfides and btsiilfides are
removed by crystallization of the carbonate. A few modern recovery systems are:
I. The STORA process,
2. The SCA-Billerud process,
3. The Tampella process,
4. The Sivola-Lurgi process, and
5. The Institute of Paper Chemistry method.
In the STORA process, the clarified green liquor from the recovery boiler, containing Na 2 S
and Na 2 CO 3 , is carbonated with recirculated CO 2 . Na 1 S is converted to H 2 S. The H 2 S is
stripped from the liquor by CO 2 as the carrier gas. The 1125 reacts in a Claus reactor with
502 to form elemental sulfur (5). The CO 2 is recirculated to the process. In this way the
green liquor is converted to NaHCO 3 . The NaHCO 3 is then reacted with NaHSO 3 to
produce Na 2 SO 3 . Part of the sulfite solution absorbs SO 2 in an absorption tower and is
returned as bisulfite.
Depending on process requirements, the cooking liquor is made from suitable proportions of
the sulfite and bisulfite solutions. When a higher pH is needed as for semichemieal pulping,
some bicarbonate is bypassed to the sulfite solution. The principal process components of
the STORA process arc illustrated in Figure 14-8.
14-22
-------
STRIPPING OF H2S AND
S02 ABSORPTION
FROM FLUE GAS
PRODUCTION OF No 2 SO 3
ABSORPTION AND
STRIPPING OF SO2
FIGURE 14-8
PRODUCTION OF ELE-
MENTAL SULFUR
(CLAUS REACTOR)
WATER
FLUE GAS’
GREEN
LIQUOR
DREGS
ELEMENTAL
SULFUR
PRODUCTION OF NaH SO 3 FORTIFICATION
STORA PROCESS FOR SODIUM BASE SULFITE CHEMICALS RECOVERY
-------
In the SCA-Billerud process, evaporated spent liquor is sprayed into a stream of hot
combustion gases containing a small excess of air, but still insufficient for the combustion of
the organic substances in the liquor. A pyrolysis of the liquors Lakes place, resulting in a
combustible gas and a powder. The gas contains almost all the sulfur as H 2 S, as well as other
combustible components, including H 2 , CO, and some hydrocarbons. The powder contains
all the sodium, mostly as carbonate with a minor part as sulfate, as well as carbon. The
process is illustrated in Figure 14-9.
The gas aiid powder mixture is cooled in a reactor boiler to a temperature well above the
dew point of the gas, which produces high pressure steam. Separation of the powder from
the gas occurs in both dry and wet separators. The gas is further cooled by condensing water
vapor in a scrubbing tower, which enriches the gas before burning. The H 2 S in the pyrolysis
gases is converted to SO 2 in a boiler. The SO 2 in the flue gases is then absorbed by a soda
lye solution before being exhausted.
The powder separated from the gas is mixed with water, and the soluble salts are leached
out to produce the Na 2 CO 3 solution used in the absorption just described. The remaining
carbon can be burned separately.
To produce cooking liquor, the product from the SO 2 absorption is fortified with makeup
sodium and sulfur in the acid making plant.
Like the STORA process, the Tampella process (Figure 14.10) is very flexible and can be
applied to various cooking methods. The essential part of the Tampella process consists of
three subprocesses: prccarbonation, bicarbonation-crystallization evaporation, and prepara-
tion of bicarbonate by carbonation. The starting point is combustion of the black liquor in a
reducing atmosphere. In the Tampella recovery process, a smelt consisting mostly of Na 2 S
and Na 2 CO 3 is formed in a conventional way.
The smelt is dissolved in a strong green liquor, which, depending on the need, can be
converted to H 2 S, a high sulfidity smelt solution, a low sulfidity smelt solution, an Na 2 CO 3
solution, or monohydrate crystals free from sulfide and thiosulfate. The H 2 S can be burned
to SO 2 for the preparation of Na 2 SO 3 solution, or converted, by means of the Claus
Process, to elementary S for the production of polysulfide liquor (1]).
‘t’o prepare polysulfidc liquor, a high sulfidity solution (88.92 percent) can be used by
combining evaporation and mechanical separation of carbonate crystals. Low sulfidity
carbonate solution can be used for preparing NSSC liquor or kraft white liquor. Crystal
Na 2 CO 3 can be employed for the neutralization of sulfite and black liquor and also for
preparing low-yield Na 2 SO 3 cooking liquor.
14.24
-------
PRODUCTION OF
DUST SEPARATION
H 2 S (DRY) (WET) STEAM
OIL
SPENT
LIQUOR
ABSORPTION OF
SO
CARBON TO EVAPORATION
PLANT
SULFUR
MAKE—UP
LEACHING OF
CHEMICALS
FIGURE 14-9
ACID TO
ACID
PLANT
FORTIFICATION
SCA PROCESS FOR SODIUM BASE SULFITE CHEMICALS RECOVERY
-------
PRODUCTION OF
NaHS
STRIPPING OF
Hz S
PRODUCTION OF
NaHCOi
FIGURE 14-10
SOt ABSORPTION FROM
FLUE GAS
I
WATER
FLUE GAS
GREEN
LIQUOR
DREGS
II
ii
0
CONTAMINATED CONDENSATL.
BLEED-OFF TO EVAPORATION
SULFUR
MAKE— UP
FORTI F ICATION
TAMPELLA PROCESS FOR SODIUM BASE SULFITE CHEMICALS RECOVERY
-------
In the Sivola-Lurgi process, green liquor, containing mainly Na 2 S and Na 2 C0 3 , reacts first
in the precarbonation tower with the flue gases from the recovery boiler and then with CO 2
in the main carbonation stage. The purpose is to convert the active sodium of the green
liquor to NaHCO 3 and to liberate the sulfide as H 2 S in thc carbonation stage. The
bicarbonate obtained is Lhcn split thermally in a decomposcr, producing Na 2 CO 3 and CO 2
for carbonation.
The major portion of the carbonate reacts with NaHSO 3 in a reactor and produces
Na 2 SO 3 and CO 2 . The reactor and the decomposer are built one above the other and
connected with a clock bottom. Portions of the Na 2 SO 3 obtained are used for cooking
liquor preparation and for sulfitation. In sulfitation step Na 2 SO 3 reacts with SO 2 to
produce bisulfite. The SO 2 is obtained by burning elementary sulfur and H 2 S in the same
furnace.
The method of the Institute of Paper Chemistry is probably the simplest. On the other
hand, Lhc products of the reaction can be controlled only within certain limits. The green
liquor is piped after clarification to the sulfitation tower, where it is treated with SO 2 from
a sulfur furnacc. The formation of thiosulfate cannot be prevented, but it can be kept at a
reasonable level, 7-15 percent iii the cooking liquor. The Na 2 SO 4 content can be kept below
5 percent. As a final product, a mixLurc of Na 2 CO 3 and Na 2 SO 3 is obtained, which, after
makeup, can be reused in the NSSC process.
14.8.3 Ammonium Base
Since this base cannot be recovercd, recovery is limited to SO 2 . SO 2 recovery can take place
in the same type of equipment as for the magnesium base. For the operation of a multistage
venturi absorber system, pH control is essential for proper functioning (see section 14.7.3).
14.9 Problem of Nitrogen Compounds from Ammonium-Based Pulping
Ammonium base SSL contains about 3 to 10 percent N 2 by weight on a dry solids basis
(ii). Nitrogen in spent sulfite liquor is present primarily in the form of ammonium salts,
which can result in the emission of ammonium sulfite and ammonium sulfate containing
particles from the recovery furnace. The gaseous forms of N 2 that can be emitted from the
recovery furnace include diatomic nitrogen gas (N 2 ), NI-I 3 , and nitrogen oxides. Amine
compounds can theoretically be emitted from the ammonium-base sulfite recovery furnace,
particularly in strongly reducing atmospheres.
Variables affecting the generation and relative amounts of the respective individual nitrogen
compounds which can be formed include the nitrogen content of the fuel, the flame
temperature, the flame zone retention time, and the configuration of the combustion unit.
]4-27
-------
The flame temperature is influenced by the heating value of the fuel, the moisture content,
the heat release rate, and the combustion volume of the furnace.
Blakeslee and Burbach (12) report that nitrogen oxide emissions tend to increase with the
nitrogen content for the coal- and oil-fired power boilers. Little information is available re-
garding Lhe effect of the fuel nitrogen content of ammonium base SSL; however, there are
indications that at the lower flame temperatures, particularly below 12000 C (2200° F), the
primary gaseous products are N 2 and NH 3 - The nitrogen present in the fuel is not present in
the diatomic elemental form, but as NH 4 ion.
Flame temperature is an important variable affecting the relative amount of nitrogen oxide
emissions formcd during the combustion of SSL. Palrnrose and Hull (13) report that the
flame temperature during ammonium base SSL combustion is about 980 to 1315° C (1800
to 2400° F). The flame temperature is observed to increase as the solids content of the
spent liquor increases, because less water is present from the fuel being evaporated to act as
a heat sink. Heat release rates for ammonium base sulfite recovery furnaces generally are
about 1.9 to 8.8 GJ/h per cubic meter (50,000 to 230,000 BTU/hr per cubic foot) (14).
Heating values are generally between 16 and 21 M i per kg of dry solids (7,000 to 9,000
BTUIIb). Loddby-type furnaces normally have relatively large combustion volumes so that
excessively high temperatures are not found when burning ammonium base SSL.
Studies by Bartok, Crawford, and Skopp (15) indicate the effect of fuel nitrogen content on
nitrogen oxide emissions. Tests run on fuel oil-fired power boilers indicate that the amount
of No formed is directly proportional to the nitrogen content of the fuel over a range of 0.2
to 1.0 percent nitrogen by weight. The nitrogen content of the fuel appears to be the
predominant factor affecting NO emissions at temperatures below 1300° C (2370° F).
Flame temperature appears to be the predominant factor affecting NO formation above
1100° C (2000° F), particularly at 02 concentrations above 2.0 percent by volume in the
flue gas.
Because of the lower nitrogen content of fuel oil (0.2 to 1.0 percent by weight) compared
to ammonium base SSL (3.0 to 10.0 percent by weight), the results of the Bartok et aL study
are not directly applicable to recovery furnaces. They do tend to indicate the general trend
of increasing NO emissions in direct proportion to fuel nitrogen content. This relationship is
applicable particularly at the lower flame temperatures observed in ammonium base sulfite
recovery furnaces because of the large combustion volumes and the high moisture contents
of the fuels.
A very limited amount of measurements have been made regarding NO emissions for
ammonium base sulfite recovery furnaces. Results of one series of tests indicate that total
nitrogen oxide concentrations ranged from 200 to 500 ppm by volume as NO 2 with peak
values observed of up to 1000 ppm (16)- A summary of results is presented in Table 14-5.
14-28
-------
TABLE 14-5
NITROGEN OXIDES* EMISSIONS FROM AN AMMONIUM BASE SULFITE
RECOVERY FURNACE (16)
Concentration Emission Rate**
Condition ppm by volume kg NO It (lb/ton) kg NO /10 6 kJ (lb NOx/10 6 BTU)
Average 350 8.3 (16.6) 0.36 (0.84)
Minimum 200 4.7 (9.4) 0.22 (0.51)
Maximum 500 11.8 (23.6) 0.52 (1.21)
*Nitrogen oxides are computed as equivalent nitrogen dioxides.
on flow rate 625 rn 3 /h per daily metric ton of capacity (24,300 ft 3 /h per short ton) and a liquor
heating value of 22.7 MJ/kg (9750 BTU/lb).
Concentrations as reported arc substantially greater than the 25 to 75 ppm of NO for kraft
and magnesium base sulfite recovery furnaces, indicating the possible role of nitrogen
content of the fuel in causing NOx emissions.
Nitrogen oxide emissions can be controlled by one or any combination of the following
methods:
1. Evaporating to the minimum solids content (about 50 percent) necessary to
support combustion and to provide efficient heat recovery, with the remaining
water acting as a heat sink during combustion;
2. Operating furnaces at relatively low firing rates to minimize the rate of heat
release;
3. Modifying the type of combustion unit in which the SSL is burned, as by shifting
to the use of fluidized beds; and
4. Operating the furnace at minimum excess air consistent with efficient combustion
in an oxidizing atmosphere.
It is also possible to reduce the nitrogen content of the fuel by reducing the amount of
NH 3 added to the chemical makeup. The effect would be to reduce both the NH 3 and
nitrogen oxide emissions from the combustion process. The NH 3 produced could be
removed in the acidic scrubbing solution during passage through the series of liquid
scrubbers. The danger in reducing the NH 3 makeup is that it could alter the pulping
conditions adversely and reduce the pH of the liquid scrubbing solution, resulting in
increased emissions of SO 2 .
14-29
-------
14.10 References
1. EKONO Oy, Helsinki, Finland, files.
2. Hendrickson, E. R., Roberson, J. E., and Koogler, J. B., Final Report, CPA 22-69-18,
U.S. Department of Health, Education, and Welfare, National Air Pollution Control
Administration, Raleigh, North Carolina, March 15, 1970.
3. Johnson, W. D., and Gansler, H., I-low Rayonier Cuts Blow Pit Emissions with
Chemical Scrubbing System. Pulp and Paper, 45(13):54-56, December 1971.
4. Axelsson, 0., and Wahlund, L. C., Occurrence of Volatile BOD Compounds in the
Sulfite Process with Spent Liquor Neutralization and Condensate Reuse. Svensk
Pappcrstidning, 75(8): 287-296, April 30, 1972 (Stockholm).
5. Simmons, T., Svcnsk Papperstidning, 67(7):286-293, April 15, 1964 (Stockholm).
6. Guerrier, J. J., Cooperative Effort Solves Small Mill’s Air Problem. Pulp and Paper,
48:62-64, March 1974.
7. Copeland, C. C., and Wheeler, C. M., A Progress Report on the Copeland Recovery
System at the Franconia Paper Corporation, Lincoln, New Hampshire. (Presented at
the TAPPI New Hampshire Section Spring Meeting, Lincoln, New Hampshire, April 23,
1970.)
8. Air Pollutants Abatement Problems of the Forest Industry. Statens Naturvarsdverk
(Sweden). Publication 1969: 3, July 1969.
9. Alaska Administrative Code, Title 18, Chapter 50.060 (1 and 2) (1973).
10. Oregon Administrative Regulations, Chapter 340, Section 25-360 (2) (1973).
11. Clement, J. L., and Sage, W. L., Ammonium Base Liquor Burning and Sulfur Dioxide
Recovery. Tappi, 52:1449-1456, August 1969.
12. Blakcslec, C. E., and Burbach, H. E., Controlling NO Emissions from Steam
Generators. Journal of Air Pollution Control Association, 23:37-42, January 1973.
13. Palmrose, G. V., and I-lull, J. H., Pilot Plant Recovery of f-feat and Sulfur from Spent
Ammonia-Base Sulfite Pulping Liquor. Tappi, 35:193-198, May 1952.
14-30
-------
14. Saiha, E., The Tampella Process. in: Proceedings of the Symposium on Recovery of
Pulping Chemicals. Helsinki, Finland, May 13.17, 1968. Finnish Pulp and Paper
Research Institute and EKONO Oy, Helsinki, Finland, 1969.
15. Bartok, W., Crawford, A. R., and Skoop, A., Control of Nitrogen Oxide Emissions
from Stationary Combustion Sources. Combustion, 42(4):37-40, October 1970.
16. Waddington, E. B., ITT.Rayonier, Inc. Shelton, Washington. September 1973. Personal
communication.
14-31
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CHAPTER 15
OTHER PROCESS SOURCES
15.1 Bleach Plant Gases
Bleach plant gases differ significantly from other gaseous emissions from a kraft mill.
Depending on the bleaching system and the bleach chemical preparation, the gaseous
effluents may include Cl 2 , chlorine dioxide (dO 2 ), and SO 2 . All of these gases have about
the same odor threshold of about 0.1.1 ppm. The bleach plant belongs to the minor air
pollution sources in a kraft mill; its influence is usually restricted to the mill area itself or to
the immediate surroundings. The noticeable effects of the bleach plant gaseous emissions are
mainly odor and corrosion of nearby objects.
15.1.1 Bleach Plant Cl 2 Emissions
II a bleach plant has a chlorination stage, there are Cl 2 emissions from the Cl 2 bleach tower
vent and from the hood vent of the succeeding washing stage. If the washer is of a vacuum
rotary drum type, the total Cl 2 emission is about 0.5 kg Cl 2 per metric ton of pulp (1 lb
Cl 2 /ton) and the total vent flow is about 900 m 3 /t (29,000 ft 3 /ton) (1).
For a pressure filter or a continuous diffuser, the vent gas flow is much smaller and more
concentrated in Cl 2 . The Cl 2 emission is much smaller, too. One feasible way of eliminating
the Cl 2 emission is to scrub the Cl 2 -containing gases with an NaOH solution. This, in turn,
can be used as a part of the hypochlorite solution needed in the hypochiorite bleaching
stage. By eliminating Cl 2 emissions, Cl 2 is saved. Profits are based on Cl 2 savings and on
scrubber investment and operational costs.
15.1.2 Bleach Plant dO 2 Emissions
if the bleach plant has two C10 2 stages, Cl0 2 is emitted from both the washer hood vents
after the bleach towers and the CIO 2 manufacturing process itself.
If the washers are of a vacuum rotary drum type, the total dO 2 emission is about 0.3 kg
CIO 2 /t (0.6 lb CIO 2 /ton) and the corresponding vent flow is approximately 800 m 3 ft
(26,000 cu ft/ton) (1). Pressure washers or continuous diffusers will give smaller flows and
emissions.
The washer hood vent gases are too diluted for CIO 2 recovery. Therefore, CIO 2 destruction
through scrubbing with an alkaline hydrogen peroxide (H 2 02) solution is the only feasible
treatment. The reaction is:
15-1
-------
2 Cl0 2 + 2 NaOH + H 2 0 2 = 2 NaCIO 2 + 02 + 1420
The sodium chlorite (NaCIO 2 ) solution must be eliminated with the mill wastewaters, since
oxidation to dO 2 with Cl 2 is hardly feasible. The H 2 0 2 is expensive, and yearly scrubbing
costs may amount to almost 30 percent of the scrubber investment.
In every CEO 2 manufacturing process, the last step consists of an absorption stage where the
dO 2 gas is run countercurrently to a stream of cool water to produce a dO 2 solution. The
C10 2 concentration is usually between 5 and 10 gIl. The C10 2 loss or emission from this
absorption tower is directly proportional to the water temperature (Figure 15-1). At a water
temperature of 200 C (68° F), the Cl0 2 emission is about 0.2 kg C 10 2 /t (0.4 Lb C 10 2 /ton)
and the tower vent flow is about 60 m 3 /t (1900 ft 3 /ton) (L).
In warm climates, where the fresh water temperature is approximately 20° C (68° F), it may
be feasible to recover more Cl0 2 in the absorption tower simply by cooling the water.
Cooling the water from 20° C (68° F) to 4° C (39° F) will decrease the dO 2 emission from
the absorption tower by 50 percent (Figure 15-1). In a colder climate, as in Canada or
Scandinavia, cooling is not feasible. Here, alkaline H 2 0 2 scrubbing must be used.
15.1.3 Bleach Plant SO 2 Emissions
A bleach plant may need SO 2 in either of two situations:
1. If it has a Mathieson process for dO 2 manufacture, or
2. If it has a final SO 2 treatment stage (to destroy Cl 2 and hypochiorite rests and to
adjust the pH of the pulp).
This treatment requires an absorption tower for manufacturing SO 2 water, and the tower
will emit some SO 2 . Gas flows are probably about 25 m 3 /t (800 ft 3 /ton) and the SO 2
emission is approximately 1.5 kg SO 2 /t (3 lb/ton) (1).
In warm climates, SO 2 recovery through cooling of the fresh water may be possible.
Reducing the fresh water temperature from 20° C (68° F) to 4° C (39° F) will decrease the
SO 2 emission by 75 percent (Figure 15.2).
In colder climates, fresh water cooling is not feasible. Tower vent gases may be scrubbed
with an alkaline solution in combination with Cl 2 emissions from the bleach plant, but then
tIie Cl 2 is no longer recovered as hypochlorite (section 15.1.1).
15-2
-------
20
o
z
0
I— ___ ___ ___ ___ ___
z
w
C-,
z
0
0 ___ ___ ___ ___ ___ ___
C ”
0
0
G b 2 PARTIAL PRESSURE, mm Hg
FIGURE 15-1
EQUILIBRIUM SOLUBILITY OF CHLORINE DIOXIDE
IN WATER (2)
15.2 Wastewater Treatment
15.2.1 Process Variables
Release of malodorous gases from liquid effluent streams during transport or wastewater
treatment operations has been frequently overlooked as an air pollution source. The major
sources of odorous gases in liquid effluent streams are from the digester and evaporator
condensate waters, with additional contributions from black liquor spills. Odorous gases can
be released to the atmosphere from free liquid surfaces during open channel flow, from
manholes and sewer vents (particularly if there is a change in liquid elevation), from
pumping stations, from liquid recycling, and from wastcwatcr treatment operations (4).
8
16
14
I2
10
8
6
4
2
0
/ v E 7 _
k ’_____
/1
0 20 40 60 80 bOO 120
140 160
15-3
-------
120
100
)80
w
I-
40
LU
a.
20
L i i
I—
0
0.1 0.2 0.4 0.6 0.8 1.0 ao 4.0 6.0 8.010
SO CONCENTRATION, %BY VOLUME
FIGURE 15-2
EQUILIBRIUM SOLUBILITY OF SULFUR DIOXIDE IN WATER (3)
The major physical and chemical processes that can result in evolution of odorous gases
from liquid effluent streams include:
1. Volatilizing of dissolved gases by increasing liquid temperature;
2. Increasing the degree of liquid turbulence by agitation, pumping, or other means;
and
3. Releasing acidic gases, such as H 2 S and CH 3 SH, by a reduction in liquid pH.
Specific processes that can result in increased evolution of odorous gases from wastewater
streams include:
1. Contact with hot flue gas streams, as during liquid scrubbing;
2. Liquid pH changes resulting from contact with acidic flue gas streams containing
C0 2 , acidic liquids such as chlorination effluent waters, or spent acid streams
from tall oil acidulation or C10 2 manufacture; and
3. Agitation of liquid streams by mechanical aeration, diffuser aeration with gases,
or other means.
15-4
-------
Process variables affecting odorous gas release include inlet concentrations of the particular
constituents, liquid temperature, pH changes, degree of agitation, physical configurations,
02 content of the liquid, and biological activity.
The major chemical constituents that produce odors from pulp and paper mill effluent
waters are organic sulfur and organic nonsulfur compounds. A major potential problem can
exist with kraft pulp mills, where organic sulfur compounds, such as CH 3 SH and CH 3 SCH 3
formed during the cooking operations, can be released into digester condensate waters. They
can also be released to a lesser extent from black liquor into evaporator condensate waters
or from black liquor spills from storage tanks or overloaded multiple-effect evaporator
systems.
Organic nonsulfur compounds, such as terpenes and other wood extractive materials, can
also be evolved from effluent waters in either kraft or sulfite pulp mills. Terpene compounds
can be evolved from digester condensate waters in kraft pulp mills in the form of droplets that
can absorb sulfur gases from subsequent transport downwind from treatment processes.
Overloaded or incompletely mixed biological waste treatment facilities with sludge
accumulations or sections with inadequate dissolved oxygen (DO) levels can be the source of
odorous gases, such as organic acids formed by anaerobic fermentation. Agitation of liquids,
where these conditions exist, can result in the liberation of substantial quantities of
malodorous gases to the atmosphere. If these gases are associated with particulate matter,
such as bacterial cells or liquid droplets, they can be transported for long distances
downwind.
15.2.2 Treatment Methods
The two major approaches to minimize malodorous gas generated from liquid effluent
streams are by inplant treatment and by modification of the wastewater treatment
facilities. The major inplant treatment methods to reduce odorous gas release to the liquid
effluents are by air or steam stripping of digester and evaporator condensate waters. (These
methods are discussed in Chapter 5 of the manual for kraft pulp mills and Chapter 14 for
sulfite pulp mills.) An additional inpiant treatment method involves effective housekeeping
to minimize black liquor spills, as well as other spills, by accidents or carelessness, systems
for segregation and containment of spilled liquids, and enclosure of flowing liquid effluent
streams.
One step that can be taken to prevent release of odorous gases from effluent waters is to en-
close tanks, as well as the biological aeration process. This step is feasible for activated sludge
aeration processes, that it would normally involve a prohibitive expense if applied to aerated
stabilization basins, because of the extensive surface areas involved. A modification to the
activated sludge prQcess was recently developed that employs molecular oxygen (02) as the
oxygen supply and uses four enclosed aeration tanks in series (5). The system results in a
15-5
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small 0 2 -rich gas stream that is easily incinerated in a lime kiln or other combustion unit.
The process also has potential as a pretreatment system, which is upstream of a conventional
aerated stabilization basin, to biologically oxidize or stabilize rapidly reactive materials, such
as terpenes or organic sulfur compounds.
Addition of specific oxidizing gases to the liquid stream will reduce odor levels in pulp and
paper mill effluent waters. Gases that have possible application for deodorization of liquid
effluent waters include Cl 2 , C10 2 , 02, and ozone (03). Alferova, Panova, and Titova (6)
describe a two-stage system, developed in the Soviet Union, for treatment of digester and
evaporator condensate waters by aeration and chlorination in series. The combined
condensate liquid is first aerated in a multiple tray tower with air and then contacted with
acidic chlorination bleachery effluent for further oxidation of the odorous compounds
present. Chlorine dioxide has been applied for effluent treatment odor control in the
petroleum industry, and can have application in the pulp and paper industry (7).
The 02 can also be used for odor control by augmenting dissolved oxygen levels in existing
secondary wastewater treatment facilities. Murray and Rayncr (8) report that CH 3 SH is
easily oxidized in the presence of 02 to CFI 3 SSCH 3 . The CH 3 SSCH 3 can then be oxidized
to sulfones, sulfonic acids, and other relatively innocuous products, but long retention times
arc required. Any CH 3 SCH 3 is not easily oxidized in the presence of 02 except at very high
temperatures not normally found in wastewater treatment.
Two recent studies describe the effects of 02 addition to liquid effluent streams for
augmenting dissolved oxygen levels in aerated stabilization basins and receiving waters. A
sidestream oxygenation system installed at the Brewton, Alabama, kraft pulp mill, where
4.5 t (5 tons) 0 2 /day were added to the liquid effluent, resulted in an increase in dissolved
oxygen levels and a decrease in odor from the aerated stabilization basin (9). Arnberg (10)
describes similar results when sidestream oxygenation is employed for augmenting dissolved
oxygen levels in receiving waters.
15.3 Odor Problems from Diffuse Sources
A complete inventory of the odorous emissions from a kraft pulp mill has, so far, comprised
measurements in stacks and other point sources only. In addition, however, odorous sulfur
compounds are emitted from sources of a diffuse nature, such as leaking process equipment
and settling basins; these must also be taken into account. Expansion of odor elimination
systems for large point sources will increase the relative importance of the diffuse sources. A
method for quantitative assessment of these emissions was developed and is described as
follows (11):
1. The method uses simultaneous sampling and flow measurement in a
suitable number of sections on the lee side of the diffuse source. The
15-6
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total emission is calculated by summation of the contributions from
individual sections.
2. The instrumentation consists of a fixed analysis unit (a gas chromatograph with a
flamephotometric detector) and portable units that arc all capable of
simultaneous gas sampling and flow measurement.
The method was tested at a kraft pulp mill in a survey of the diffuse emissions from the
settling basin, as well as from the digestery, recovery furnace, and limewashing building. The
results of the measurcments are given in Table 15-1.
TABLE 15-1
ODOROUS GASES FROM DIFFUSE KRAFT PULP MILL SOURCES (11)
Emitted Compound-Emission Rate
Diffuse Emission Source H 2 S CH 3 SH CH 3 SCH 3 CH 3 SSCH 3
g/h (lb/hr)
Settling Basin Measurement I 1,900 600 60 400
(4.2) (1.3) (0.13) (0.88)
Settling Basin Measurement II 1,700 600 70 400
(3.8) (1.3) (0.15) (0.88)
Settling Basin Measurement III 1,800 600 80 600
(9.0) (1.3) (0.18) (1.3)
Digestery I (fir) 2 10 60 40
(0.004) (0.022) (0.13) (0.088)
Digestery II (birch) 3 10 40 20
(0.007) (0.022) (0.088) (0.044)
Recovery furnace I 4 10 10 30
(0.009) (0.022) (0.022) (0.066)
Recovery furnace II 10 40 20 60
(0.022) (0.088) (0.044) (0.13)
Limewash building 1
(0.002)
15-7
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15.4 References
1. EKONO Oy, Helsinki, Finland, files.
2. HaIler, I. F., and Northgraves, W. W. Tappi, 38: 199, April 1955.
3. The Finnish Paper Engineers’ Association (SPY). The Pulping of Wood. Frenckdllin
Kirjapaino Oy, Helsinki, 1968 (Finnish).
4. SabLeski, J. J., Odor Control Lfl Kraft Pulp Mills: A Summary of the State of the Art.
U.S. Department of Health, Education, and Welfare, Public Health Service, National
Center for Air PoUution Control, Cincinnati, Ohio, May 10, 1967.
5. Grader, R. J., South, W. D., and Djordjevic, B., The Activated Sludge Process Using
High Purity Oxygen for Treating Kraft Mill Wastewaters. Tappi, 56:103-107, April
1973.
6. Alferova, L. A., Panova, V. A., and Titova, G. A., Deodorization of Effluents from the
Manufacture of Kraft Pulp. Bumazhnaya Promyshlennost (Moscow, USSR), 38, (6),
5.8, June 1963.
7. Samsel, J. J., and Hawkins, E. A., Waste Water Treatment at Texaco’s Puget Sound
Refinery. Proceedings of the American Petroleum Institute, Division of Refining,
40(III):302.308, 1960.
8. Murray, F. E., and Rayner, H. B., The Oxidation of Dimethyl Sulfide with Molecular
Oxygen. Pulp and Paper Magazine, 69(9):64.67, May 3, 1968.
9. 0 and 03 - RxforPollution. Chemical Engineering, 77(4):46-48, February 23, 1970.
10. Amberg, H. R., Wise, P. W., and Aspitarte, T. R., Aeration of Streams with Air and
Molecular Oxygen. Tappi, 52:1866-1871, October 1969.
11. Institutet for Vattenoch Luftvbrdsforshning Butte tin, 2(2) :6.7, 1973 (Stockholm).
15-8
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CHAPTER 16
16.1 Supply Patterns
POWER BOILERS
The pulp and paper industry is a major energy consumer in the United States with a total
process energy consumption requiremcnt in 1970 of about 1.6 X 101 8 J per year
(1.5 X lOl 5 BTU per year). (Sec Chapter 1.) Based on 1972 figures compiled by the
American Paper Institute, about 33 percent of the industry’s total energy requirement is
supplied by combustion of its waste pulping liquors and an additional 7 percent by
combustion of wastewood and bark. The remaining 60 percent of the total energy
requirement must be supplied by purchase of auxiliary fossil fuels, such as coal, oil, and
natural gas or by purchase of electricity. A breakdown of the total national process energy
supply patterns for the pulp and paper industry is presented in Table 16-1 (1). Considerable
variations in energy supply for mills in different regions of the United States can exist.
TABLE 16.1
OVERALL NATIONAL DISTRiBUTION OF ENERGY SOURCES FOR THE PULP
AND PAPER INDUSTRY (1) (2)
Energy Source
% of TotaP
Overall Energy Consumption
MJ/yr
( BTU/yr)**
Pulping liquors
Waste wood
Bituminous coal
Residual fuel oil
Distillate fuel oil
Natural gas
Purchased electricity
Other sources
Total
33
7
11
20
2
21
1
100
522 X iO
110 X iO
174 X i0 9
317 X iO
32 X i0 9
332 X iO
79 X iO
16 X i0 9
1582 X iO
(495)X 1012
(105)X 1012
(165)X 1012
(300)X 1012
(30)X 1012
(315)X 1012
(75)X 1012
( 15)X 1012
(1500) X 1012
*Based on 1972 data.
**Based on 1970 data.
16-1
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16.2 Combustion Parameters
To obtain usable energy from fuel, fuel is normally burned in air to release heat and then
the heat energy is harnessed by using it to generate steam. The kinetic and thermal energy of
the steam, in turn, can be used in any variety of ways. ‘I’o make the fuel burn satisfactorily,
it must be mixed with air in the proper proportions in a suitable furnace. To convert water
to steam at appropriate pressure, the water normally flows through tubes, which arc warmed
on their outside by the hot gaseous products of combustion. Once steam forms, it is further
heated to the desired degree of superheat by channeling it through another set of similar
tubes, called the superheater. The types and amounts of air pollutants emitted from the
combustion process depend on the characteristics of the fuel burned, the configuration of
the combustion unit, and the operating parameters of fuel and air supply.
16.2.1 Fuel Characteristics
The major parameters affecting air pollutant formation during fuel combustion in power
boilers arc the physical state, heating value and moisture content of the fuel, as well as its
ash, sulfur, and nitrogen contents. Coal and wastewood are the primary types of solid fuels
employed; both require extensive materials handling facilities for the unburned fuel before
combustion and for the noncornbustible ash following combustion. Both distillate and
residual liquid fuel oils are employed in the pulp and paper industry. Heavy residual fuel oil
is more commonly used and has a high air pollution potential. Natural gas is the primary
gaseous fuel employed in the pulp and paper industry; it is easiest to burn and has the
lowest air pollution potential of all the fuels. Each of the different fuels requires a different
type of combustion unit because of differences in burning characteristics.
The heating value of a fuel determines the amount that must be burned to generate a given
amount of usable energy. The heating value varies with the kind of fuel employed, the
location from which it was derived, and its moisture content. Heating values for bituminous
coal and wastewood can vary substantially among those derived from different locations, as
can their ash and sulfur contents. A summary of typical heating values, moisture contents,
sulfur and ash contents of coal, oil, gas, and wood fuels employed in pulp and paper
manufacture are presented in Table 16.2 (3).
Major fuel consumption parameters affecting air pollutant emissions from power boilers in
the pulp and paper industry include sulfur and ash contents of the fuels. The SO 2 emissions
from fuel combustion are directly proportional to the sulfur content. The sulfur content is
normally significant in bituminous coal and residual fuel oil. The major methods for
minimizing air pollution from fuel combustion of SO 2 in power boilers are the substitution
of low sulfur oil, coal, or natural gas for high sulfur fuels and the construction of tall stacks
for dilution of ground level SO 2 concentrations.
16.2
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TABLE 16-2
CHARACTERISTICS OF FUEL BURNED IN POWER BOILERS AT PULP AND
PAPER MILLS IN THE UN1TED STATES (3)
% by Wt. Content of
Heating Value (As Fired) Moisture Ash Sulfur
Fuel Units Average ( Range) Average (Range )
Coal Mi/kg 31 (24-34) 10 (7-15) 8.1 (3.5-15) 1.9 (0.5-10)
Btu/lb 13,500 (10,500-14,700)
Oil MJ/l 41 (34-43) 0.5 (0.1) 0.1 (0.01-0.2) 1.8 (0.1-3.5)
Btu/gal 149,000 (122,000-155,000)
Gas Mi/rn 3 38 (37-40) 0.2 (0-1) Neg. Neg.
Btu/cu ft 1,030 (1,000-1,070)
Wood Mi/kg 11 (9.3-13) 25 (10.60) 2.9 (0.1-20) 0_i (0-0.2)
Btu/lb 4,600 (4,000-5,500)
The particulate emissions from fuel combustion generally are proportional to the ash
content of the fuel and arc significant for coal and wood combustion. The ash material
contained in the fuel can be removed from the boiler as fly ash in the exhaust gases or as
bottom ash in solid form or as liquid slag, depending on the ash fusion temperature.
Normally 80 to 95 percent of the ash material is emitted from thc boiler as fly ash.
Unburned or partially burned fuel aslo can be emitted from the combustion chamber as
particulate matter. Devices normally used for particulate emission control on coal-fired
power boilers in the pulp and paper industry are electrostatic precipitators or mechanical
collectors; mechanical collectors and liquid scrubbcrs arc most commonly used for
wastewood- and bark-fired power boilers.
Recent investigations show that other constituents in fuels may also contribute to air
pollution. Martin and Berkau (4) report that the nitrogen content of fuels can contribute to
the formation of nitrogen oxides, though the reaction between atmospheric N 2 and 02 is
predominant for flame temperatures above 1,300° C (2,400° F). In addition, trace elements,
particularly in bituminous coal and residual fuel oil, may be significant air pollutants, such
as the nonmctals chlorine (Cl), fluorine (Fl), and phosphorus (P), and the heavy metals,
such as beryllium (Be), mercury (Hg), lead (Pb), cadmium (Cd), zinc (Zn), arsenic (As), and
selenium (Se). A summary of air pollutant emission factors from specific fuel combustion
processes for industrial power boilers is presented in Table 16-3 (5).
16-3
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TABLE 16-3
UNCONTROLLED AIR POLLUTANT EMISSION FACTORS FOR FUEL COMBUSTION IN POWER BOILERS FOR THE
PULP AND PAPER INDUSTRY (5)
Particulate Matter Sulfur Oxides Nitrogen Oxides Hydrocarbons Carbon Monoxide
kg Part lb Part kg SO 2 lb SO 2 kg NO 2 lb NO 2 kg CH 4 lb CH 4 kg CO lb CO
Fuel Burned 106 kJ 106 BTU 106 kJ 106 BTU 106 kJ 106 BTU 106 kJ 106 BTU 106 kJ 106 13’l’U
Bituminous Coal
Pulverized General 0.34 (%A) 0.80 (%A) 0.82 (%S) 1.90 (%S) 0.39 0.900 0.0064 0.015 0.021 0.050
Pulverized Wet Bottom 0.28 (%A) 0.65 (%A) 0.82 (%S) 1.90 (%S) 0.64 1.500 0.0064 0.015 0.021 0.050
Pulverized Dry Bottom 0.37 (%A) 0.85 (%A) 0.82 (%S) 1.90 (%S) 0.39 0.900 0.0064 0.015 0.021 0.050
Pulverized Cyclone 0.04(%A) 0.10(%A) 082(%S) 1.90(%S) 1.17 2.740 0.0064 0.015 0.021 0.050
Spreader Stoker 0.28 (%A) 0.65 (%A) 0.82 (%S) 1.90 (%S) 0.32 0.750 0.0064 0.015 0.021 0.050
9” Fuel Oil
Residual-Tangential 0023 0053 046 (%S) 1.06 (%S) 0.115 0.267 0.0005 0.00] 00005 0.001
Residual-Horizontal 0.023 0.053 046(%S) L06(%S) 0.230 0.535 0.0005 0.001 0.0005 0.001
Distillate-Tangential 0.015 0.035 0.41 (%S) 0.96 (%S) 0.123 0.285 0.0005 0.001 0.0005 0.001
Distillate-Horizontal 0.015 0.035 0.41 (%S) 0.96 (%S) 0.246 0.570 0.0005 0.001 0.0005 0.001
Natural Gas
Process Boilers 0.006 0.0 14 0.00048 0.00 1 0.160 0.372 0.0 16 0.038 0.0005 0.00 1
Gas Turbines 0.006 0.014 0.00048 0.001 0.183 0.425 0.016 0.038 0.0005 0.001
Waste Wood
NoReinjection 1.50 3.500 0.16 0.375 0.43 1.010 0.11 0.250 0.11 0.250
50% Reinjection 1.79 4.150 0.16 0.375 0.43 0.010 0.11 0.250 0.11 0.250
100% Rcinjection 2.32 5.400 0.16 0.375 0.43 0.010 0.11 0.250 0.11 0.250
%A = Percent ash in fuel.
%S = Percent sulfur in fuel.
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16.2.2 Furnace Characteristics
Each fuel burned must be fired in a separate type of boiler depending on its physical state
and burning characteristics. Natural gas is normally burned for steam generation in package
type boilers of varying sizes, where the gas-air mixture can be added in a horizontal
configuration along one side. Burning residual fuel oil in power boilers is similar to burning
in gas-fired boilers where the oil-air mixture is normally added in either a horizontal or
corner tangential firing con figuration.
The burning of oil requires a more complex fuel delivery system, normally with pumping,
than gas. Residual oil is also usually heated in inlet piping by steam tracing to reduce its
viscosity and, thcreforc, prevent plugging of the fuel lines. Tangential firing of oil or gas can
result in substantial reductions in nitrogen oxide concentrations as compared to horizontal
firing because of the reduced flame temperatures, greater air-fuel mixing, and resultant
lower excess air requirements (6).
Burning solid fuels in power boilers, such as coal or wastewood, requires even more complex
design than burning oil in power boilers because of the more complex combustion process
and the more complex fuel delivery and ash handling systems. The major types of coal-fired
power boiler employed in the pulp and paper industry are the spreader stoker, the chain
grate stoker (now in limited use), and several varieties of pulverized firing systems, including
dry bottom, wet bottom, and cyclone type units, which can be horizontally, tangentially, or
vertically fired. The type and design of boiler used depends on the burning characteristics
and chemical composition of the coal; these include the ash content, ash softening
temperature, fixed carbon content, volatile carbon content, and heating value. More
thorough presentations of design parameters for coal-fired power boilers are available in
several references (7, 8, 9, 10).
Wastewood- and bark-fired power boilers can burn the wood alone or can be modified to
burn other fuels on an auxiliary basis or in combination. Wastewood and bark are burned in
power boilers on chain grates in a radiant Dutch-oven type boiler or in a horizontal
air-blown suspended firing configuration in a vertical Stirling boiler. Wood handling systems
(including hammermill grinding to a given particle size for suspended firing), bottom and fly
ash handling systems, and underfire and overfire air controls must be provided. Major fuel
characteristics affecting the design of wastewood-fired power boilers include ash and
moisture content, particle size variations, and fixed and volatile carbon content. The
amount of particulate matter swept from the combustion chamber is normally greater from
the horizontal suspension firing units than from the Dutch-oven type units.
16.2.3 Combustion Variables
The most important variable affecting the combustion process is the fuel-to-air ratio, which
can be varied between wide limits even when the other variables remain unchanged.
16-5
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Variations in the fuel-to-air ratio can be expressed as variations in the amount of excess air
or as variations in the oxygen content of the flue gas. The variations influence the
combustion temperature, combustion efficiency, ignition speed, flame length, flame
radiation, and flue gas composition. Poor combustion efficiency increases particulate
emissions, while high combustion temperature, together wiLh an oxygen-rich furnace
atmosphere, produces oxidcs of nitrogen.
Appropriate mixing of fuel and combustion air is essential to achieve fast and complete
combustion. The design of firing and furnace has the major influence on the mixing, but
mixing also can be controlled to some extent by operational variables such as velocity and
direction of fuel, and especially air streams. If mixing is not sufficient, part of the fuel
remains unburned and results in particulate emissions. The formation of nitrogen oxides is
known to be higher in diffusion flames, where mixing takes place in the flame rather than in
premixed flames (11). Nitrogen oxide formation increases with both flame temperature and
02 level in the combustion zone.
16.3 Boiler Types
The major parameters affecting the design of power boilers to minimize potential air
pollutant emissions include:
1. Physical state, chemical composition, and burning characteristics of the fuel,
2. Method of firing fuel and mixing,
3. Combustion chamber volume and configuration, and
4. Other operating parameters, such as fuel firing rates, excess air levels, and air
distribution.
The type, size, and collection efficiency requirements for the gas treatment system are
determined by the gas flow rate and flow conditions from the boiler, the exit gas pollutant
concentrations from the boiler, and specific parameters, such as particle size. Whatever the
fuel, the air pollutant emission characteristics from power boilers are very much influenced
by design and operating parameters.
16.3.1 Gas-Fired Power Boilers
Natural gas-fired power boilers normally are the simplest in design and operation of any of
those used in the pulp and paper industry. ‘They require only piping of the gas to the boiler
and mixing it with air to provide adequate heat for steam generation; they require no
complex fuel or ash material handling systems. To this extent, the design and operation of
16.6
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these boilers are essentially the same as gas-fired boilers in the electrical utiliLy industry, but
the units in the pulp and paper industry are normally smaller and generate steam chiefly for
process use instead of for driving turbines. The only potential air pollutants from natural
gas-fired power boilers arc oxides of nitrogen generated at elevated temperatures during the
corn bustion process.
The two major types of natural gas-fired power boilers used in the pulp and paper industry
are horizontal- and tangential-fired units. These units differ only in the method of air-fuel
introduction to the combustion chamber. The horizontal-fired units employ firing along one
side of the furnace; the tangential units use corner-firing. The nitrogen oxide emissions from
horizontal-fired units are normally greater than from tangential-fired units because of their
characteristically higher flame tern pcraturcs.
The lack of radiating particles in a gas flamc requires a large furnace volume to ensure
sufficient cooling of the flue gas before entering the other parts of a boiler. A certain
amount of particle formation is useful and can be achieved with special burner design or
with two-phase firing for boilers designed primarily for burning natural gas. Natural gas can
also be burned in pulp and paper industry power b ilers that are designed primarily for
other fuels and where the gas can be used either for startup or combination firing. Because
of the relative simplicity of gas firing in combination boilers, the other fuels dominate the
design requirements and the boiler emissions.
The major operating variables for gas-fired power boilers arc the amount of excess air added
and the distribution of primary and secondary air if off-stoichiometric firing is used. Flue
gas reeirculation to the furnace can also be added to improve combustion efficiency and to
redLice excess air requirements. Operating natural gas-fired power boilers at sufficient excess
air allows complete combustion, but more important it also promotes nitrogen oxide
formation (12).
The sulfur and ash content of natural gas are negligible, and the respective SO 2 and
particulate matter emissions are insignificant. Complete combustion assures that minimal
emissions of CO and hydrocarbons will occur. The major air pollutants from natural gas
combustion are nitrogen oxides, which can be minimized by off-stoichiometric firing, use of
minimal excess air, and flue gas recirculation (13).
16.3.2 Oil-Fired Boilers
Oil-fired power boilers are normally similar in design to natural gas-fired units except Lhat
they require more complex fuel handling facilities and longer retention times in the
combustion zone. The two major types of liquid petroleum fuels burned are No. 2 distillate
fuel oil and No. 6 residual fuel oil. Both types of oils require fuel storage, pumping and
injection systems, with steam or electrical tracing of fuel lines needed for residual fuel oil to
16-7
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avoid plugging by the highly viscous liquid. Atomization of the oil injected into the furnace
from the firing guns is necessary to provide a spray of fine droplets, assuring complete
combustion of the fuel. The major potential air pollutants from fuel oil combustion are
nitrogen oxides and sulfur oxides, and, to a somewhat lesser extent, particulate matter.
The two major types of furnace configurations for burning fuel oil in pulp and paper
industry power boilers arc horizontal front-fired units and tangential corner-fired units.
These arc similar in design to those employed for natural gas firing except that longer
retention times in the combustion zone arc normally required to assure complete
combustion of the oil. Atomizing and mixing promote combustion, and thereby reduce the
formation of unburned carbon particles to the practical minimum.
The furnace volume must be sufficiently large enough to prevent the flame from impinging
on the cooled walls; a cooled flame would inhibit the last steps of the combustion reactions
and result in the emission of carbon soot particles. The volume of the furnace varies directly
with the released thermal energy, and the formation of nitrogen oxides, in turn, is
proportional to this encrgyr. Because a smaller combustion chamber causes an increased hcat
release rate per unit volume, an increase in flame temperature results, which is accompanied
by an increase in formation of nitrogen oxides.
Major parameters affecting the operation of oil-fired power boilers arc:
1. Excess air level,
2. Air flow distribution pattern,
3. Fuel inlet temperature as it affects oil viscosity, and
4. Amount of steam added for atomization.
The excess air level dtiring fuel oil firing must be carefully controlled within the maximum
and minimum limits determined by’ consideration of both thermal energy conversion
efficiency and air pollutant emission levels. The lower practical limit for e xcess air is reached
when carbon particles and combustible gases, such as CO and hydrocarbons, are detected in
the flue gas. The tipper practical limit for excess air is defined by the formation of
significant quantities of nitrogen oxides and the presence of sulfur trioxide (SO 3 ). Sulfur
trioxide formation is accelerated by the presence of trace metal catalysts, such as vanadium
(V) in the oil. An additional consideration is that excess air levels above the tipper limits
result in a decrease in the thermal efficiency of the boiler by increasing the flue gas flow rate
and the flue gas outlet temperature.
16-8
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An important variable in maintaining sufficient atomization is the viscosity of the fuel oil,
which is maintained by keeping the oil temperature wiLhin rather narrow limits, increased
viscosity at low oil temperatures dampens the vibration of oil droplets and thus reduces the
atomization effect. Too high a fuel temperature tends to cause coking of the oil on the
hottest parts of the burner and disturbs the atomization process by producing steam and gas
bubbles in the oil before it is atomized (14).
The major potential air pollutants from fuel oil combustion are sulfur oxides, nitrogen
oxides, and particulate matter.
The SO 2 emissions from fuel oil combustion are essentially a function of the sulfur content
of the fuel, which normally ranges from 0.1 to 0.5 perccnt by weight for light distillate oils
and from 0.5 to 3.0 percent, or more, by weight for heavy residual fuel oils, depending on
the source of the crude. The major method for controlling SO 2 emissions from oil
corn bustion in the pulp and paper industry, to date, is the substitution of natural gas or low
sulfur oil for high sulfur oil. Varying quantities of 503, accounting for up to 10 percent of
the total sulfur burned, can also be formed during fuel oil combustion. The SO 3 can be
hydrolyzed in the presence of water to 1-12 SO 4 and cause corrosion on cold metal surfaces.
The amount of SO 3 formed depends on the sulfur content of the fuel, the excess air lcvel,
and the possible presence of trace metal catalysts in the oil; high concentrations of up to 50
or 60 ppm by volume have been observed (15).
Particulate matter emitted from fuel oil combustion consists of inorganic material from the
ash content of the fuel and organic materials resulting from incomplcte combustion. The
organic matter from fuel oil combustion consists primarily of unburned carbon soot
particles resulting from incomplete combustion of the oil droplets, approximately 95
percent by weight of the soot particles are less than 10 pm (4 X 10 in) in diameter. The
carbon content of the particulate matter can be as high as 58 percent by weight, and
because of this high carbon content, these particles are not particularly suitable for
collection by electrostatic precipitation (16).
The emission of particulate matter from oil firing is strongly dependent on the excess air
level during combustion. Any major disturbance in burner operation is almost certain to
produce a considerable increase in the particulate matter emissions. The burning of residual
fuel oils with sulfur contents above 1.5 percent by weight can cause the particulate matter
from the boiler to become saturated with adsorbed H 2 50 4 droplets when the gas stream
temperature drops below the acid dew point. The sticky acidic particle deposits can cause
rapid corrosion of metal surfaces and, therefore, lower the efficiency of the mechanical
cyclone collectors.
Nitrogen oxide controls for oil.fired power boilers are similar to those for gas-fired units; the
objectives are to minimize flame temperatures and excess air levels. Nitrogen oxide
16-9
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formation is controlled by off-stoichiometric firing with split air distribution, minimum
excess air, and possibly, flue gas rceirculation.
16.3.3 Coal-Fired Boilers
Coal-fired power boilers are normally more complex in design than either oil- or gas-fired
units because complex solid fuel handling and ash handling systems are required, the fuel
injection system is more complicated, and longer retention times with greater air-fuel
mixture turbulence levels in the furnace are generally required to assure complete
combustion. The main type of coal burned at present for auxiliary power generation in the
pulp and paper industry is the bituminous grade from the eastern or midwestern United
States. The properties of coals can vary considerably between individual fields. The
respective sulfur contents, ash contents, and heating values all affect potential air pollutant
emissions. The major potential air pollutants from coal combustion in pulp and paper
industry power boilers are sulfur oxides, particulate matter, and nitrogen oxides.
There are several types of coal-fired power boilers employed in the pulp and paper industry.
These differ mainly in the methods by which the fuel is added to the furnace, the
configurations by which air is added to the combustion chamber, and the boiler bottom
configuration. The major coal feed configurations commonly employed in the pulp and
paper industry include pulverized firing, spreader stoker, and chain-grate stoker units. The
type of unit employed depends on the amount and properties of the coal to be burned. The
ash content, ash softening temperature, fixed carbon and volatile carbon contents, moisture
content, and, to a lesser extent, sulfur content all influence the type of unit employed. The
sizes and shapes of the furnaces are largely influenced by the heating value and ratio of fixed
to volatile carbon contents in the coal.
The smaller coal-fired power boilers in the pulp and paper industry are generally either
underfeed or overfeed chain-grate stoker units or suspended-firing spreader stoker units. The
larger coal-fired power boilers arc generally pulverized-firing, employing either vertical,
horizontally opposed, tangential corner, or wall-fired units, with either dry or wet bottoms.
The pulverized units are employed for coals with high ash softening temperatures, while
coals with ash softening temperatures below 1,2000 C (2,200° F) arc generally burned in
cyclone furnaces, though these are not commonly employed in the pulp and paper industry.
The design features of the individual types of coal combustion units are extensively
reviewed in other references (7, 8, 9, 10).
The particulate emissions in the flue gases from coal-fired power boilers tend to increase
with increasing coal ash content and with increasing ash softening temperature. They also
basically show the greatest increase where pulverized units employ dry bottoms because of
suspension of small particles. Particulate emissions are lowest from pulverized cyclone units
because most of the ash is removed as slag from the furnace bottom. The flue gas
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temperature of the furnace section outlet normally must be maintained below Lhc ash
softening temperature of the coal being burned in order to keep particulate matter from
melting arid sticking on tLlbcs in the superheater section. Coal furnaces must generally be
designed to provide sufficient cooling by means of waLer wails or other mclhods to prevent
superheater slagging. The resultant lower flame zone temperature also acts to reduic
nitrogen oxide emissions.
Major air operating variables affecting emissions from coal-fired power boilers include the
excess air level and the relative distribution of primary and secondary air (or unclerfirc or
overfire air). The excess air level is significant in that increased CO and hydrocarbon
emissions are favored by minimum excess air levels. These minimum levels also can increase
the carbon content of the particulate matter leaving the boiler. This increased carbon
content can adversely affect operation of the electrostatic precipitator by increasing particle
resistivity levels. high excess air levels tend to increase the amounts of nitrogen oxides
present. Excessive overfirc air levels on chain.grate or spreader stoker units can also exert a
sweeping action to entrain small particles in the exhaust gases.
Fuel.related operating variables influence particulate emissions from coal.fircd power
boilers. The fineness of pulverized coal particles is influenced by the degree of grinding,
which is a factor that may vary during the lifetime of the grinding elements in the
pulverizer. If not controlled, pulverizer wear affects the combustion efficiency by causing
progressively larger particles that increase the carbon content in the fly ash. In grate firings,
an excessive air flow can penetrate areas in the fuel layer if its deposit on the grate is riot of
uniform Lhickness. Localized particulate entrainment results.
The particulate emissions from coal-fired power boilers depend on the dsh content of the
coal, its ash softening temperature, the method of firing employed, and to a lesser extent on
the excess air for organic constituents. Particulate emissions generally arc greatest from
pulverized dry bottom units, lower from chain grate and spreader stoker units, and lowest
from cyclone furnaces. The degree of coal grinding before firing and the type of firing
employed both influence the particle size distribution in the exhaust gases. There is
insufficient information regarding the particle size distribution from pulp and paper
industry coal-fired units at present. it is known, however, that the amount of particulate
matter discharged as fly ash relative to the amount of slagged bottom ash increases with ash
softening temperature.
SO 2 emissions from coal combustion are directly proportional to the sulfur content of the
coal being burned. Increases in excess air level tend to cause increases in SO 3 emissions from
coal combustion. But the presence of calcium and magnesium oxides in the particulate
matter tends to bind any SO 3 as CaSO 4 or MgSO 4 . Present control methods for sulfur
oxides include fuel substitution and the construction of tall stacks. Nitrogen oxide levels
from coal combustion tend to increase with combustion zone flame temperatures and excess
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air levels in the exhaust gases. Nitrogen oxide levels arc comparable to those from oil-fired
units, except for cyclone furnaces, where the high temperatures required for ash slagging
also cause excessive formation of nitrogen oxides.
16.3.4 Wood-Fired Boilers
Waste wood combustion is often employed in pulp and paper mills because the material is
readily available from wood debarking or associated lumber and plywood manufacturing
operations. The major types of materials that can be burned in waste wood boilers are bark
from mechanical or hydraulic debarking operations and waste wood materials, such as
sawdust, shavings, slabs, and chips from lumber and plywood manufacture. Major factors
affecting waste wood boiler design and use are the quality and amount of material available
and its burning characteristics. Waste wood materials arc often burned in combination with
other fuels such as oil, gas, coal, or clarifier sludge from wastewater treatment. The major
potential air pollutant from waste wood combustion is particulate matter; hydrocarbons and
nitrogen oxides are also emitted.
The starting point in designing a waste wood firing furnace is different from that for other
fuels because bark and wood arc waste fuels to be incinerated. The objective of firing waste
wood is to maximize the use of the released energy from waste wood to minimize the
amount of other fuel to be purchased.
Major design parameters for waste wood boilers are heating value of the fuel, moisture
content, ash content, particle size, waste type, and wood species. The moisture content of
the fuel is an important parameter; an increasing moisture content value results in a
decreasing net heating value due to increased water evaporation and a resultant increase in
the flue gas volume. The moisture content of the fuel depends on the type of debarking and
the storage time. Mechanical debarking (10 to 30 percent water by weight) results in a lower
moisture content than hydraulic debarking (40 to 60 percent water by weight).
The net heating value for most waste woods is about 17 to 21 MJ per kg of dry wood (7,200
to 9,000 BTU/Ib), or 8.4 to 13 Mi/kg (3,600 to 5,400 BTU/lb) on an as-fired basis. The
heating value for waste woods tends to vary with wood species. The presence of extractive
materials, such as terpenes and tall oils, can substantially add to the energy content of the
wood.
The particulate emissions from waste wood combustion vary with ash content and parLicle
size of the material bcing burned. The ash content can vary from below 1 to 20 pcrcenL by
weight on an as-fired basis. The sizes and shapes of the wood particles being burned can
influcnce the design of grating systems, the type of firing employed, and the relative
distributions of underfirc and overfire air in the furnace. Furnace fouling is often found in
bark firing, especially when bark is fired together with other fuels, in some cases, small
16-12
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amounts of minerals, gathered in the fuel during storing in sea water or on ground, can
lower the ash softening temperature so that it is very sticky it the furnace and cannot be
easily removed.
Thw major types of combustion units employed for waste wood burning employ pile
burning or suspension burning. The successive operations of water evaporation, volatile
carbon distillation and oxidation, and fixed carbon oxidation must occur in series. Flat grate
Dutch.oven boilers have been extensively used for waste wood combustion in the past.
These employ both underfire- and overfire.air jets as applied to a stationary grate and do not
normally require extensive pregrinding of the fuel. Thin bed suspension firing is often used
for combination firing of wood and other fuels. It allows higher firing rates and is normally
employed for large units. A disadvantage is that it normally requires prcgrinding of the
wood in a hammcrmil.
Major operating variables affecting the combustion reactions in wood-fired power boilers are
the excess air level, the ratio of overfire to undcrfirc air, the combustion temperature, and
the pile or suspension bed thickness to regulate fuel burning rate. The overfire air jets should
be located and operated to produce a minimum of entrained particulate matter.
The operation of a grate firing should be as even as possible for the best efficiency, so that
its combustion air can be correctly regulated. A grate firing in a separate combustion
chamber can be operated with substoichiometric air flow, if the furnace temperature must
be kept low, or if the furnace atmosphere must be reducing. If neither of these conditions is
required, the grate must have the proper excess air to burn combustible gases before they
enter the main furnace.
The major potential air pollutant from wood.fired power boilers in the pulp and paper
industry is particulate matter, which can result from either inorganic ash in the wood or
from incomplete combustion.
The particulate matter from wood-firing is often large in size, 5 to 10 pm (2 to
4 X iO in) or greater. Its specific gravity is usually low so that the use of mechanical
cyclone collectors is not always possible. The electrical properties of the fly ash are not very
suitable for electrostatic precipitation because of the high carbon content that causes high
particle resistivity. The minerals in the wood can cause abrasion in the collecting equipment
and ducts resulting in rapid metal wear (3). More complete analysis and classification of the
chemical composition and physical size characteristics of particulate matter emitted from
wood-fired power boilers is needed than has been reported to date.
Potential gaseous emissions from wood-fired power boilers are oxides of nitrogen, oxides of
sulfur, and hydrocarbons resulting from volatilization of the wood. The nitrogen oxide
emissions from wood-fired power boilers are generally lower than for fossil fuel firing
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because of the large combustion volumes pcr unit amount of fuel burned, the normally high
excess air levels of 50 percent or more, and the high fuci moisture content that results in
low flame temperatures of 980 to 1,200° C (],800 to 2,2000 F). Emissions of 502
from wood-fired power boilers generally arc low because the sulfur content of wood is
generally less than 0.1 percent by weight. The emissions of terpenes, hydrocarbons, and
other volatile organic constituents by distillation and incomplete combustion vary with
wood species, furnace temperature, and retention time. The extent of these emissions as
potential air pollutants has not been fully described.
16.4 Particulate Emissions
Particulate emissions from power boilers consist of inorganic ash from the fuel and partially
burned or unburned fuel from combustion processes. Both of these components vary
considerably with the type of fuel and firing. The relative amount of organic material
present depends primarily on the excess air level and retention time in the combustion zone.
The ash contents of most fuels used in the pulp and paper industry normally are less than 5
percent by weight, with the exception of coal, which normally varies from 5 to 15 percent
by weight. The fly ash in coal consists mainly of quartz, aluminum oxides, iron oxides, and
alkali oxides (17). Smaller amounts of trace metals and other materials also may be present.
The composition of the ash in coal, wood, or oil depends largely on the source of the fuel.
The amount of unburned carbon in the form of soot and grit is very dependent on the
quality of oil firing; while fly ash from the firing of coal and wood normally contains only a
few percent unburned carbon. This smaller amount again depends on the quality of firing,
on excess air addition, on the ratio of primary to secondary air, and especially on the
behavior of the fuel on the grate.
Particulate emissions from power boilers normally do not contain large quantities of trace
metals or organic hazardous materials (18).
The control of particulate emissions from coal- and wood-fired power boilers is not
normally as complex as other processes within the pulp and paper industry. The problems of
control are more similar to those found in the power generation field, and the approaches to
their solution are the same. Electrostatic precipitation is the most efficient method for
particulate control for coal-fired power boilers; cyclone collectors offer a less efficient, but
also less expensive, way for either coal- or wood-fired boilers. Liquid scrubbers or fabric
filters are effective for particles below 5 pm (2 X io in) and offer an alternative to
electrostatic precipitation, particularly for wood-fired boilers. Typical particulate emissions
from power boilers in the pulp and paper industry are presented in Table 16-4 (3), and
typical particle size characteristics for coal- and wood-fired power boilers are presented in
Table 16-5 (3).
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TABLE 16-4
PARTICULATE EMISSION CHARACTERISTICS FROM SELECTED U.S. POWER BOILERS (3)
Collector Equipment
Number Percent of Fuel Supplied, Pressure
of Btu Basis Drop Particulate Concentration Collection Emission
Boilers Coal Oil Gas B/W* Type** in. of water Inlet Outlet Efficiency Rate
gIm 3 (gr/cu ft) % kg/h (Ib/hr)
18 100 0 0 0 C 2.5 4.28 (1.87) 0.85 (0.37) 80 129 (284)
2 100 0 0 0 5 0.57 (0.25) — 136 (300)
2 100 0 0 0 P 2.2 11.2 (4.89) 0.98 (0.43) 91 180 (397)
16 0 46 0 54 C 2.7 7.9 (3.47) 1.05 (0.46) 87 140 (309)
2 75 0 0 25 C 3.9 0.41 (0.18) — 73 (160)
2 0 0 62 38 C 2.8 5.3 (2.30) 0.39 (0.17) 93 70(153)
2 73 16 0 11 C 2.5 — 2.75 (1.20) — 228 (502)
3 0 25 39 36 C 2.8 4.3 (1.88) 0.71 (0.31) 84 202 (445)
2 0 0 0 100 C 0.2 3.2 (1.40) 0.89 (0.39) 72 44 (96)
*B/W = bark and wood waste.
= Cyclone.
S = Liquid scrubber.
P = Electrostatic precipitator.
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TABLE 16-5
TYPICAL PARTICLE SIZE DISTRIBUTION OF FLY ASH FROM COAL- AND WOOD.l?IRED POWER BOILERS (3)
Percent by Weight of Particles in a Given Size Range
Coal-Fired Power Boilers Bark- 1’ired
Particic Pulverized Cyclonc Spreader Traveling Underfeed Power
Diameter Coal Furnace Stoker Grate Stoker Boilers
pm (in)
0-10(04X]0 4 ) 25 72 ]1 7 12
E0-20(4-8X 10 ) 24 15 12 — 8 10
20-30(8-12X10 4 ) 16 6 9 I I 6 7
30-40(12-16X10 4 ) 14 2 10 — 9 6
40-75(16-30X10 4 ) 13 — 12 12 8 14
75-150 (30-60 X ]0 ) 6 5 17 30 19 16
I50 (60X]0 4 +) 2 29 47 43 35
Total [ 00 100 100 100 100 100
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16.4.1 Electrostatic Precipitators
Electrostatic precipitators are used primarily for coal-fired power boilers in the pulp and
paper industry. Electrostatic precipitation is suitable for particulate emission control where
high collection efficiency is required, but only a small pressure drop can be tolerated. The
capital cost of electrostatic precipitators is relatively high, and electric power is consumed
during their operation. The particulate collection efficiencies are lower for oil.fired or
wood.fired boilers because the carbon content of the fly ash causes considerable increases in
particle resistivity as compared to coal-fired units.
The design of electrostatic precipitators depends on gas flow rate, gas temperature, gas
humidity, inlet dust loading, and the electrical properties of dust, such as particle resistivity.
The resistivity essentially defines the design migration velocity of the particles to the
collection electrodes. The gas flow rate determines the design gas velocity, which defines the
necessary width of the precipitator for operation. The particle size distribution in the inlet
dust has minor influence on the design of the precipitator, unless a major portion of the
dust is less than 1 pm (39 X 1O— in) in diameter. At such small particle sizes, the
decreasing particle collection efficiency must be compensated for by reduced gas velocity.
The precipitator reduces the particle size distribution by selectively removing the larger
particles. On the other hand, when a precipitator failure occurs, the resulting particles
emitted are selectively both larger and heavier. This operating characteristic should be taken
into account when selecting continuous monitoring equipment (19).
The factor which varies the most in the operation of an electrostatic precipitator is the gas
velocity or gas flow rate, which varies with boiler load and, in combination boilers, with
fuel-to-fuel ratio. Smaller variations in the gas velocity are caused by changes in the gas
temperature and in the humidity of the gas. Temperature and humidity are treated as
independent variables in determining the efficiency of the precipitator. The humidity of the
flue gas depends on the moisture and hydrogen content of the fuel. The quality of the fuel
and the firing together cause variations in the unburned carbon content of the dust, which
influences the precipitator efficiency and the inlet dust loading. Electrostatic precipitators
can be applied to the combination firing of coal and bark.
16.4.2 Cyclone Collectors
Mechanical cyclone collectors are used primarily for particulate emission control on
wood-fired power boilers, or as first stage collectors for coal-fired units. In a cyclone
collector, the flue gas with its dust burden is fed into a centrifugal field, where centrifugal
forces separate dust particles from the flue gas. Single cyclones or multicyclone equipment
is used. Multicyclones consist of a group of small cyclones arranged in parallel. The
minimum particle size collected by a single cyclone is about 10 pm (4 X iO in); while
multiple cyclones will collect 5 pm (2 X i t T 4 in) particles. A separated flow with high dust
16-17
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content may further be conducted into a secondary separation with cyclones, and the flue
gas recirculated to the main collector.
A cyclone collector has a lower capital cost than an electrostatic precipitator and may also
be applicable to coal- and oil-fired power boilers. Because of the energy required to maintain
centrifugal movement, the pressure drop across the cyclone is inevitably high. Fouling can
be serious, especially in multicyclone equipment, where gas and dust must pass through
rather small ducts. Therefore, frequent maintenance of equipment may be required.
The design of a cyclone collector depends on gas flow rate, gas temperature, gas humidity,
inlet dust loading, density of particles, and particle size distribution. The abrasive properties
of particles may set special requirements for materials used in construction of cyclones and
associated ductwork; while the temperature and humidity of gas have their effects on
fouling, particularly for bark char or sander dust combustion (20).
The gas flow rate from a power boiler defines the size of a single cyclone or the number of
cyclones required in multicyclone equipment, depending on the inlet velocity required for
effective separation of particles from the gas stream by centrifugal forces. This velocity, in
turn, depends on the density and size distribution of particles in the flue gas stream from
the power boilers. The circulating movement can be introduced to the gas either by
tangential inlet or by radial vanes with axial inlet. The central outlet tube for clean gas may
be located either straight ahead in a horizontal cyclone or at the inlet end of a vertical or
inclined cyclone, causing a 180° turn and additional separation. A secondary circuit, with a
fan to induce suction, may be connected into the separated dust outlet.
A single- or monocyelone-collector is applicable for rather coarse dust and small flow rates.
A small boiler firing woodwaste and bark is a typical application. A multicyclone collector
allows for large flow rates by adding to the number of cyclones. The cyclone size can be
optimized for required separation and pressure drop.
The separation in a cyclone depends strongly on the centripetal acceleration of the gas and,
therefore, on the gas flow rate. This must be kept within acceptable limits to achieve
operation between inadequate separation at low velocity and too high a pressure drop at
high velocity. Flow rate can be controlled by shutting off certain parts in a multicyclone
coLlector; otherwise, a low collection efficiency must be expected at low boiler loads.
The particle size distribution may change along with boiler load, different fuels, and other
changes in firing. Therefore, the collection efficiency of a cyclonic collector also changes.
There is no general rule to control the particle size distribution during firing except to try to
maintain relatively uniform conditions.
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16.4.3 Liquid Scrubbers
Liquid scrubbing is used primarily for particulate collection on wood- and bark-fired power
boilers in the pulp and paper industry. Scrubbing traps thc particulate matter entrained in
the gas stream in liquid for subsequent removal and disposal. Liquid scrubbing has the
advantages of being able to remove gaseous and particulate materials simultaneously, of
removing fine particles below 1 pm (4 )< 10 in) in diameter, of recovering additional
thermal energy by cooling the gas stream, and possibly of improving primary clarifier sludge
settling characteristics. Vertical impingement and venturi scrubbers have been the primary
types of unit employed, to date. Successful installations were reported in Montana by
Effenberger (21), in Texas by Ritchey (22), and in South Carolina by Pearce (23).
Several major variables affect the design of liquid scrubbing systems for particulate emission
control on wood-fired power boilers. Major gas stream variables include the overall gas flow
rate, fuel firing rate, flue gas temperature and moisture content, and the allowable scrubber
pressure drop as determined by fan capacity characteristics. Major particulate matter
characteristics include the total particulate mass concentration and the particle size
distribution, particularly for particles less than 10 pm (4 X i0 in) in diameter.
Liquid-phase design parameters include the makeup and recycle shower rates, liquid
pumping capacity, nozzle configurations and sizes, and allowable slurry solids
concentrations. The physical configuration and gas.liquid contact geometry chosen arc
determined by the type of scrubber purchased or designed. The materials of construction,
such as stainless steel, must be chosen to avoid corrosion, abrasion, and thermal damage.
The operation of a liquid scrubber is influenced primarily by the liquid phase parameters
such as recycle flow rate, makeup water flow rate, nozzle liquid pressure drop, and slurry
solids concentration. Gas phase pressure drop is usually subject to a certain amount of
adjustment, but is also influenced by the gas flow rate as determined by fuel firing rates.
The scrubbing systems described by Effcnberger (21) and Etitchey (22) operate at gas
pressure drops of about 15 to 25 cm (6 to 10 inches) of water and are able to achieve
particulate mass concentrations at standard conditions of 0.02 to 0.05 g/m 3 (0.01 to 0.02
gr/cu ft) or less, corresponding to particulate mass removal efficiencies of 99 percent or
greater. Detailed particle size measurements for these units were not provided on either
inlets or outlets to the respective collectors. Liquid pH must be controlled to avoid
corrosion, particularly if coal or oil is burned in combination with wood. Liquid scrubbcrs
can prove particularly useful to upgrade the particulate collection efficiencies of existing
mechanical cyclone collectors on wood-fired power boilers.
16.4.4 Fabric Filters
Fabric filters have not been extensively used for particulate collection on power boilers in
the pulp and paper industry. They have the advantages of high removal efficiencies for fine
16-19
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particles of less than 1 .im (4 X i0 in) in diameter with pressure drops of 8 to 15 cm (3 to
6 in) of water. There is not enough operating experience with fabric filtration, to date, on
wood-, coal-, or oil-fired power boilers in the pulp and paper industry to present detailed
design and operating parameters.
16.5 References
L. Slinn, R. J., The Paper Industry ’s Energy: A Survey by the American Paper Institute.
Southern Pulp and Paper Manufacturer, 37:39, March 1974.
2. Miller, R. R., One Pulp and Paper Company’s View of the Energy Crisis. Tappi,
57:62-64, February 1974.
3. Hendrickson, E. R., Roberson, J. E., and Kooglcr, J. B., Control of Atmospheric
Emissions in the Wood Pulping Industry, Volume II. Final Report, Contract No. CAP
22-69-1 8, U.S. Department of Health, Education, and Welfare, National Air Pollution
Control Administration, March 15, 1970.
4. Martin, C. B., and Bcrkau, E. E., Combustion Processes and Air Pollution. (Presented
at the National Meeting of the American Institute of Chemical Engineers. Atlantic
City. August 30, 1971.)
5. Compilation of Air Pollutant Emission Factors. U.S. Environmental Protection
Agency, Office of Air Programs, Research Triangle Park, North Carolina. Publication
No. AP-42, February 1972.
6. Cuffc, S. T., and Cerstle, R. XV., Emissions from Coal-Fired Plants: A Comprehensive
Survey. U.S. Public Health Service, Cincinnati, Ohio. Publication No. 999-AP-35. 1967.
7. Fryling, G. R., Combustion Engineering, Revised Edition. New York, Combustion
Engineering, Inc., 1966.
8. Steam: its Generation and Use, 38th Edition, The Babcock & Wilcox Company, New
York, 1972.
9. Caron, A. L., The Control of Particulate and Gaseous Emissions from Coal-Fired
Stationary Contbustion Units. National Council of the Paper Industry for Air and
Stream Improvement. New York. Atmospheric Pollution Technical Bulletin No. 42.
October 1969.
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10. Technology Transfer Process Design Manual for Pollution Control in the Fossil Fuel
Electrical Utility Industry. Prepared by Radian Corporaticin, Austin, Texas, for the
U.S. Environmental Protection Agency, Washington, D.C., 1974. In press.
11. Lowes, T. M., and Heap, M. P., The Emission of Oxides of Nitrogen from Natural Gas
and Pulverized Fuel Flames. International Flame Research Foundation. Ijmuiden, The
Netherlands, Document No. D 09/a/8 2 , 1972.
12. Walsh, R. T., Boilers, Heaters and Steam Generators. In: Air Pollution Engineering
Manual, Danielson, J. A. (ed.). U.S. Public Hcalth Service. Cincinnati, Ohio.
Publication No. 999.AP .40, 1967.
13. McGuire, W. F., Thompson, P. C., and Smith, L. L., Theory and Application of Nitric
Oxide Emission Reduction in Utility Boilers. In: Proceedings of the First Annual
Symposium on Air Pollution Control in the Southwest. Texas A&M University, College
Station, Texas. November 5-7, 1973.
14. Bagdon, K. M., Viscosimetry of Oil Burner Control. Instrumentation Technology,
19:43-46, February 1970.
15. Mineur, J., and Hulden, B., The Sulphur Problems in Oilfired Boilers: A Review.
EKONO OY. Helsinki, Finland. Publication Series No. 116. 1968.
16. Smith, W., Atmospheric Emissions from Fuel Oil Combustion. U.S. Dept. of Health,
Education and Welfare, Public Health Service Publication No. 999-AP.2. November
1962.
17. Tankha, A., Try Fabric Dust Collectors on Small Boilers. Power, 117(5):72-73, August
1973.
18. First Draft Report, Group of Experts on Emission Measurement Techniques for
Particulate Matter from Selected Sources. OECD. Paris. Addendum 1 to
NR/ENV/73.25. 1973.
19. Cooper, H. B. H., The Particulate Problem: Continuous Particulate Monitoring in the
Pulp and Paper Industry. In: Proceedings of the Symposium on Instrumentation for
Continuous Monitoring of Air and Water Quality. Miami University, Pulp and Paper
Foundation. Oxford, Ohio. June 20, 1973.
20. Barron, A., Studies on the Collection of Bark Char Throughout the Industry. Tappi,
53:1441-1448, August 1970.
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21. Effcnbergcr, H. K., Gradle, D. 0., and Tomany, 1. P., Control of Hogged Fuel Boiler
Emissions. Tappi, 56:111-115, February 1973.
22. Ritchey, J. R., Ventun Wet Scrubber for Particulate Control on a Bark Boiler. In:
Proceedings of the First Annual Symposium on Air Pollution Control in the
Southwest. Texas A&M University, College Station, Texas. November 6, 1973.
23. Pearce, A. E., Mechanical Dust Collection with Secondary Wet Scrubbing as Applied to
a Bark Fired Power Boiler. In: New Approaches to Particulate Collection at Bark Fired
Power Boilers. National Council of the Paper Industry for Air and Stream
Improvement, Inc., New York. Atmospheric Pollution Technical Bulletin No. 51,
October 1970.
16.22
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CHAPTER 17
PROCESS MONITORING
17.1 Source Measurements
To determine compliance with existing and proposed pollutant emission standards and to
inventory material losses, the types and amounts of gaseous and particulate materials from
pulp and paper mill flue gases must be measured. Developing methods for direct and
accurate measurement of air pollutant levels for TRS compounds, oxides of sulfur, and
particulate matter is important because of increasingly stringent emission standards in all
levels of government.
Source testing of particulate matter emissions from pulp and paper mill flue gas streams can
be performed by both batch- and continuous-sampling methods. Batch testing provides an
average concentration value for a given time period. Continuous monitoring provides a
record of instantaneous concentration values over a prolonged time interval to determine
compliance with air pollution regulations and to act as a monitor of pçoeess equipment
operation. Continuous monitoring instruments normally necessitate a higher capital cost
than batch sampling equipment, but manpower requirements are normally lower once these
systems are placed in operation and maintained by competent, trained personnel.
The successful operation of continuous monitoring systems for measuring particulate matter
and gaseous emissions requires instrumentation that is accurate, reliable, stable,
reproducible, of sLmple operation and low maintenance requirements, and subject to
minimal interferences. In the design and operation of continuous monitonng
instrumentation, both the sample handling and detector systems must be considered. A
suitable system for reducing and reporting the voluminous amounts of data that can be
generated is also needed.
17.2 Gaseous Monitoring
The major classes of gaseous pollutants emitted from pulp and paper industry sources that
may require continuous or batch monitoring are malodorous sulfur compounds, oxides of
sulfur, oxides of nitrogen, and organic nonsulfur compounds. The major malodorous sulfur
compounds of interest include H 2 5, mercaptans, dialkyl sulfides, and dialkyl disulfides.
These are commonly classified together as TRS compounds.
The major oxides of sulfur include SO 2 and SO 3 from combustion processes where
sulfur-containing fuel is burned. Oxides of nitrogen of interest include NO and NO 2 from
17-1
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combustion processes. Organic nonsulfur compounds include aliphatic, olefinic, and
aromatic hydrocarbons, terpcnes, phenols, and other organic compounds from kraft and
sulfite mill sources.
Gaseous monitoring systems normally have similar, and in some cases, common, sample
conditioning systems. The different gaseous constituents are monitored by gas detection
devices that depend on the constituent monitored.
17.2.] Sample Conditioning Systems—General
The purpose of the sample handling and conditioning system is to remove the sample from
the flue gas and transfer it to the detector for subsequent analysis without changing the
concentration or character of the constituents to be measured. The major elements of the
sample conditioning system are the sample probe, the liquid scrubber, the preselective gas
separator, the transfer tubing, and the prime mover, as shown in Figure 17-1 (1).
17.2.1.1 Sample Probe
The sample probe is located internally within the stack to remove a portion of the moving
gas stream from the duct into the sample conditioning system. A straight stainless steel tube
with a 90 degree bend is sufficient for flue gas streams containing negligible quantities of
particulatc matter, such as digester, washer, evaporator, and black-liquor oxidation tower
gases. The probe should be aligned so that any condensate formed drains away from the
stack towards the scrubber.
The particulate matter present in flue gases from recovery furnaces, lime kilns, and smelt
tanks must be removed by one of two types of filtration systems. The system devised by the
National Council for Air and Stream Improvement employs an open end 25 mm (1 in)
diameter tube packed with glass wool for particle removal that must be cleaned and
repacked periodically to prevent plugging (2).
The system devised by Thoen (3) is designed to prevent plugging by using a porous ceramic
probe equipped with a compressed air blowback feature that is actuated by solenoid valving
and a timer for 30 second intervals once each 10 minutes. The ceramic probe is normally
employed in particulate laden gas streams above their dew points where condensation is not
likely, such as in power boilers and kraft recovery furnaces. Where significant quantities of
SO 3 are present, such as from oil-fired power boilers, the sample probe should be heated to
150° C (300° F) or higher to prevent condensation of sulfuric acid.
17.2.1.2 Conditioning Device
The major purpose of the conditioning device is to prevent condensation of water vapor in
the gas stream. One approach calls for passing the gas stream from the sample probe to the
17-2
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I AMPLIFIER
RECORDER
GAS SAMPLE HANDLING & CONDITIONING SYSTEM FOR EXTERNALLY
LOCATED CONTINUOUS GASEOUS MONITORING SYSTEM
—
C .A
SAMPLE DETECTION
CONDITIONING SYSTEM
FLOW FLOW
CONTROL MEASUREMENT PRIME
MOVER
FIGURE 17-1
-------
detector through a line heated to a temperature above its dew point. Maintaining the sample
lines at above 1100 C (230° F) is normally necessary. inlet sample probes must also be
heated to 150° C (300° F) or higher to effect evaporation, particularly if there are mist
droplets entrained in the gas stream, such as sulfuric acid droplets.
A second approach is to employ some type of condenser in the sample line downstream of
the probe. The condenser serves the dual functions of removal of the water vapor and
cooling of the gas stream before it enters the detector. The condenser can effectively remove
the water if the gas stream is cooled to 30° C (86° F) or lower, but it is normally necessary
to acidify the condensed moisture to a pH of 2.0 or less by adding H 2 SO 4 to minimize
absorption of SO 2 .
Under some circumstances the gas stream should be diluted with air upstream of the
detector to cool the gas stream or to prevent condensation. To control the air flow rate into
the system accurately, the exact degree of dilution must be known. This dilution technique
is not suitable for extrcmely low gas concentrations approaching the sensitivity of the
detector.
17.2.1.3 Prime Mover
The remaining portions of the gas handling and conditioning system are the flow
measurement and control section, the drying section, and the prime mover. A desiccant,
such as silica gel or Drierite, can be used to protect either the flow meter or the vacuum
pump downstream of the detector, but it is not commonly used. Micrometering needle
valves made of stainless steel can be used to control the gas flow rate, since they are able to
resist the corrosive gas conditions. Flow metering is done with a rotometer or orifice flow
meter.
The prime mover can be located either upstream or downstream of the detection cell
depending on the system in use. A positive displacement vacuum pump that is leakproof and
sealed may be put in use upstream of the detection unit. Upstream location of the vacuum
pump is particularly desirable if a considerably larger volume of stack gas is removed from
the duet than is sent to the detector, or if the detector must be operated under positive
pressure. Such an arrangement does involve potential problems of particulate plugging and
moisture condensation, as well as sample dilution by air leakage. Location of the prime
mover downstream of the detection cell allows use of either a mechanical vacuum pump, or
a steam, air, or water aspirator.
17.2.1.4 Sampling Lines
Sampling lines should be of sufficient diameter to provide for minimum pressure drop, but
small enough for minimum retention time. Tubing of 0.635 cm (0.25 in) to 1.27 cm (0.5 in)
17-4
-------
is normally optimum for sample line construction. Wall materials of inert polyethylene or
Teflon should be used to avoid possible losses by physical adsorption or chemical reaction.
Also, electrically heated Teflon tubing is commercially available for sampling lines of up to
60 m (200 ft) in length. Tygon and rubber tubing, plus carbon steel, cast iron, and copper
fittings all react with sulfur compounds and should not be used.
17.2.2 TRS Monitoring Systems
The major malodorous sulfur compounds of interest for continuous monitoring applications
in kraft pulp mills include H 2 S, CH 3 SH, Cl-I 3 SCH 3 , and CH 3 SSCH 3 . Major sources of H 2 S
from the kraft process include the recovery furnace, smelt tank, lime kiln, multiple-effect
evaporator, and tall oil vent gases. Major sources of the organic sulfur compounds include
the digester blow and relief gases, brown stock washer hood and seal tank vents,
multiple-effect evaporator noncondensablc gas and condensate liquid streams, black liquor
oxidation tower exhausts, and the recovery furnace used following direct contact
evaporation. The organic sulfur compounds can often create problems in sample handling
because of their tendency to condense and adhere to tubing walls.
17.2.2.1 Gas Conditioning
Accurately transferring TRS compounds from the flue gas to the detector is complicated by
the presence of large quantities of water vapor and particulate matter in the flue gas at
elevated gas temperatures. Terpenes, SO 2 , and organic sulfur compounds can be interfering
constituents in coulometric detection systems.
It is first necessary to remove sulfur dioxide from combustion sources such as recovery
furnaces, lime kilns, and smelt tanks. A liquid scrubber is located immediately downstream
of the probe which contains a solution of potassium acid phthalate (K11C 8 H 4 0 4 ). The
purpose of the scrubber is to selectively remove sulfur dioxide from the flue gas, condense
excess water vapor, remove additional particulate matter, and cool the gas stream from stack
to ambient temperature conditions. Two different types of potassium acid phthalate
scrubbers can be used, the continuous flow and the batch, nonflow types.
The NCASI continuous flow scrubber employs a two chamber system containing liquid
absorption and overflow chambers in series. A three percent solution of potassium acid
phthalate passes from a storage bottle through a glass wool filter in the bottleneck to
remove particulate matter to prevent plugging (2). The liquid flow rate to the scrubber is
controlled by a limiting flow orifice of capillary tubing at a rate between 0.5 and
1.0 cm 3 /min, which requires replenishment once per week. Gas is drawn at a rate of
25 cm 3 /min through the scrubber of 40 cm 3 capacity. A 1.0 m (3 feet) long dropleg is used
to maintain suction and prevent leakage.
17-5
-------
TRS losses, as 1-125, are between 0.1 and 0.2 ppm by volume at a liquid flow rate of
1.0 cm 3 /min because of the finite solubility of the compounds in the KHC 8 H 4 O 4 solution.
However, the TRS losses increase considerably with increasing liquid flows above this rate.
The scrubbing system developed by Thoen (3) is a batch system employing a saturated
solution of KHC 8 H 4 O 4 which requires replenishment once every 30 days and which has
lower total reduced sulfur losses once the system reaches equilibrium. As the solution is
depleted, SO 2 removal is less efficient, and may cause interferences in the detector.
Condensation at water vapor in the system may also cause malfunctions in the system.
After the SO 2 scrubber, the gas stream passes through a system employing preselective
serubbers where analytical selectivity can be obtained between the respective reduced sulfur
gas constituents. Thoen (4) developed scrubbing solutions which could be used for selective
removal of H 2 S, I-1 2 S plus CH 3 SH, and H 2 5 plus Cl-1 3 51-1 plus CH 3 SCH 3 . Respective
constituent concentrations are monitored by differences in instrument readings. The
solutions used are listed in Table 17-1.
TABLE 17-1
SELECTIVE PRESCRUBBING SOLUTIONS FOR SULFUR GAS SEPARATION (4)
Scrubbing
Gases Removed Solution Concentration
%bywt.
SO 2 KHC 3 H 4 O 4 3
H 2 S÷SO 2 Cd50 4 -H 3 BO 3 1-2
H 2 5+5O 2 +RSH NaOH 10
SO 2 +H 2 5+RSH-4-RSR AgNO 3 0.5
An additional procedure used on certain gas conditioning systems is to pass the flue gases
through a combustion furnace at a temperature of approximately 815° C (1500° F) to
convert the reduced sulfur compounds to 502 upstream of the detector (5). Conversion of
the reduced sulfur compounds to SO 2 makes it feasible to use SO 2 sensitive detection
methods, such as flame photometry and ultraviolet spectrophotometry in addition to
coulometrie titration. An additional advantage of the combustion process is that it converts
potentially interfering organic compounds, such as terpenes and olefinic and aromatic
hydrocarbons to nonreactive CO 2 and water. The combustion step is necessary for kraft mill
sources containing extensive quantities of terpenes, such as digester blow and relief gases,
washer hood and seal Lank vents, and black-liquor oxidation tower exhausts.
17-6
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17.2.2.2 Gas Detection Systems
Available detectors for continuous monitoring of reduced sulfur compound levels in kraft
pulp mill process streams include coulometric titration, clectrochemical membrane sensing,
flame photometry, and ultraviolet spectrophotometry. TRS monitoring is complicated by
interference from SO 2 and by the sensitivity of some detectors to only 1125 and 502. It is
often either necessary or desirable to convert the organic sulfur compounds and 1125 to SO 2
by oxidation upstream of the detector, particularly when using detection methods other
than coulometric titration.
Coulometric titration is useful for measuring concentrations of H 2 S and organic sulfur gases
in kraft pulp mill flue gases; but it is also sensitive to olefinic and aromatic hydrocarbons,
terpenes, acrolein (CH 2 CHCHO), and NO 2 (6).
The technique operates on the principle that the ‘-‘2 , 502, and organic sulfur compounds
present arc oxidized to sulfate ion by the action of a halogen electrolyte. A certain electrical
current is required to maintain a constant concentration of halogen gas generated from the
electrolyte. The current level is proportional to the overall concentration of reactive gases
passing through the detection cell (7). Coulometric titration is sensitive to all compounds
that can be oxidized by the halogen in varying degrees peculiar to each compound and is not
specific to any one compound.
A commercially available instrument employs a solution of 16 percent hydrobromic acid
(HBr) as the electrolyte with sensing, generating, and referehce electrodes all located on a
common shaft (8). The detection cell can be actuated to generate a series of fixed levels of
bromine gas (Br 2 ) by changes in instrument attenuation setting corresponding to
concentration ranges varying from 1 to 800 ppm full scale as 1 125-
Flue gas is drawn into the cell where the presence of reactive gases is detected by the
consumption of bromine gas generated as the result of the chemical oxidation reactions. The
current required for maintaining a constant bromine level in the cell increases as the total
concentration of reactive gases increases. The current is converted into an equivalent voltage
potential in an amplifier and the signal is transmitted to a continuous printout on a 0 to
100 my recorder.
Calibration of jndividual cells for particular gases to be measured is necessary to assure
accurate results. The reason is that the response of the instrument to the gases passing
through the cell depends both on their concentrations and on the valence states of the sulfur
atoms relative to bromine. Each gas then must be calibrated individually to determine its
relative response in the cell. The response to particular gases also varies between individual
detection cells because of differences in hydraulic characteristics, background noise level,
and quality control. A summary of approximate calibration factors for sulfur gases observed
for a Barton coulometric titrator is listed in Table 17-2 (9).
17-7
-------
TABLE 17-2
APPROXIMATE RANGES IN CALIBRATION FACTORS FOR SULFUR GASES
WITH COULOMETRIC TITRATOR (9)
Concentration Factor (ppm by volume/scale unit)
Attenuation H 2 S CH 3 SH CH 3 SCH 3 CH 3 SSCH 3 SO 2
0.1 0.008-0.013 0.014-0.016 0.035-0.040 0.030-0.035 0.03-0.04
0.3 0.030-0.035 0.040-0.045 0.090-0.100 0.09-0.12 0.09-0.12
1.0 0.09-0.12 0.12-0.13 0.31-0.37 0.25-0.30 0.30-0.35
3.0 0.24-0.26 0.35-0.38 1.0-1.2 0.5-1.0 0.9-1.0
10.0 0.7-0.9 1.2-1.4 3.5-4.0 2.0-3.0 2.5-3.5
30.0 1.5-2.5 3.0-3.5 11.0-12.0 5.0-8.0 8.0-10.0
100.0 4.5-9.0 10.0-13.0 35.0-40.0 20.0-25.0 25.0-30.0
Cells should be calibrated on a weekly basis to assure continued accuracy, particularly for
high concentration levels (above 50 ppm by volume as H 2 S) because of cell response drift.
The electrolyte solution must be changed at least once per month to avoid depletion. The
system also appears to require frequent maintenance and suffers from particulate plugging in
the cell. Three possible arrangements for the system are illustrated in Figure 17-2.
Electrochemical membrane cells that are specific for either SO 2 or E l2 S can be used for
monitoring reduced sulfur gas emissions. It is normally necessary to convert the reduced
sulfur compounds to SO 2 in an oxidation furnace upstream of the SO 2 selective membrane
cell. The H 2 S selective membrane cell cannot be used to measure TRS levels in kraft pulp
mill flue gases because of interferences from organic sulfur compounds.
The principle of operation is similar to that of coulometric titration. SO 2 operates as the
working material in a specially constructed electrochemical transducer cell of proprietary
design. The detection cell is a totally enclosed system where the gas sample passes adjacent
to a semipermeable plastic membrane across which the SO 2 or H 2 S diffuses into the
electrolyte solution. The gas produces a change in the electrochemical potential across the
cell that is directly proportional to the concentration in the gas stream over a concentration
range from 0.01 to 5,000 ppm by volume.
The gas handling system for the electrochemical membrane cell is similar to that of the
coulometric titrator as shown in Figure 17-3. The two major differences arc that it is
necessary to locate a leakproof vacuum pump upstream of the detector because the
membrane cells must be operated under positive pressure, and a combustion furnace is
needed to oxidize the reduced sulfur compounds to SO 2 . Parallel detection cells can be
located to record both SO 2 and TRS simultaneously. The gas flow to the detector cell is
approximately 500 cm 3 /min.
17-8
-------
A. NCASI SYSTEM(1):
B. WEYERHAEUSER SYSTEM(2):
FIGURE 17-2
CONTINUOUS
SOURCE MONITORING SYSTEM FOR REDUCED SULFUR
EMISSIONS WITH COULOMETR IC TITRATION
RECORDER
(O-lOOmv)
TRAP
DRYING
TUBE
DETECTION
CELL
VACUUM
PUMP
AIR TIMERS
RECORDER
(O-lOOmv)
SURGE SO 2
BOTTLE SCRUBBER
TUBE
PUMP
DETECTION
CELL
17-9
-------
C. HUMBOLT COUNTY SYSTEM(4):
FIGURE 17-2
CONTINUOUS SOURCE MONITORING SYSTEM FOR REDUCED SULFUR
EMISSIONS WITH COULOMETR IC TITRATION—CONTINUED
FIGURE 17-3
TOTAL REDUCED SULFUR MONITORING WITH AN ELECTROCHEMICAL
MEMBRANE CELL DETECTOR
SpccLrophotometric methods are used to measure SO 2 and must be used with an oxidation
furnace upstream of the detection device to convert the reduced sulfur compounds to SO 2 .
Methods of detecting SO 2 are discussed in scction 17.2.3.
Some reduced sulfur compounds may be monitored by gas chromatographic methods, as
discussed in section 17.2.5.
RECORDER
(0- lOOmv)
VACUI
PUMP
so,
SCRUBBER
DRYING
TUBE
DETECTION
CELL
‘I
KHP
STORAGE
:R
DETECT ION
VACUUM CELL
PUMP
17-10
-------
17.2.3 Sulfur Dioxidc Monitoring Systems
SO 2 can be released to the atmosphere from either kraft or sulfite pulp mill process sources
or from the combustion of sulfur-containing fuels in coal- or oil-fired power boilers. Major
potential sources of SO 2 emissions in the kraft process include the recovery furnace, lime
kiln, and smelt tank. The major potential sources of SO 2 in sulfite pulp mills include the
digester, evaporator, and acid-making stage vent gases, and the recovery furnace.
A small amount of the SO 2 formed can be subsequently oxidized to SO 3 , which is
converted to H 2 SO 4 mist at temperatures sufficiently below the acid dew point. Normally,
the major constituent to be monitored in flue gas streams in SO 2 . The SO 2 concentrations
can be monitored either within a duct or externally following sample collection and
withdrawal.
17.2.3.1 Sample Conditioning
The major constituents making sampling difficult are the large quantities of water vapor and
particulate matter in flue gas streams. The SO 2 can be easily removed from the inlet gas
stream during condensation of water because of its relatively high solubility. The SO 2 can
also react with certain tubing wall materials, resulting in additional loss of SO 2 . Therefore,
design of the sample handling and conditioning systems must be carefully made to avoid
sucli losses.
17.2.3.2 Detection Systems
Detection systems used for continuous monitoring of SO 2 include electrolytic conductivity,
electrochemical transducer conversion, coulometric titration, and ultraviolet spectrometry.
The first three are all located externally from the stack and require prior sample withdrawal,
while ultraviolet spectrophotometry can be performed either internally or externally.
Internal location of the detector eliminates any possible problems caused by air leakage or
moisture condensation that might occur during sample withdrawal. External location is the
only feasible approach for certain detectors and results in fewer problems with particulate
interferences and high gas temperatures.
Miller, Brown, and Abrams (10) describe a continuous monitoring system for measuring
total sulfur oxides (SO 2 and SO 3 ) in the flue gas streams from digester blowpits and
acid-making absorption towers in sulfite pulp mills. The stack gas is withdrawn from the
duct and passed through a water scrubber. The soluble gases dissolve and pass through an
externally located conductivity cell. The relatively simple, inexpensive nonspecific detector
is suitable for measurement of sulfur oxide levels in the highly moisture-laden sources
because only negligible quantities of interfering, conductivity-producing gases and
particulate matter are present in the gas stream.
17-11
-------
The gas is withdrawn from the duct at a rate of 3 1/mm (0.1 cfm) and mixed with deionized
water at a rate of 1 1/mm (0.3 gpm). The gas-liquid mixture passes concurrently downward
through a 5.1 cm (2 in) diameter tower 66 cm (26 in) long for SO 2 absorption and then into
a separation chamber where the gas stream is withdrawn through a vacuum pump. The
liquid stream containing the dissolved sulfur oxides is then passed through a conductivity
cell that can measure SO 2 concentrations in ranges of either 0 to 1,000 or 0 to 10,000 ppm
by volume. A 10 my recorder is used. The system, as illustrated in Figure 17-4, has proved
successful over extended periods of operation.
Electrochemical transducer membrane cells can be employed for selective measurement of
SO 2 concentrations in combustion unit and process source flue gases. The principle of
operation was discussed in section 17.2.2.2.
Electrochemical membrane cells specific for SO 2 are relatively free from chemical
interferences from oxides of nitrogen, SO 3 , water vapor, and hydrocarbons. They maintain
a relatively stable calibration without substantial drift in response for extended periods with
minimum maintenance. The major operating difficulties are that the presence of particulate
matter tends to plug the membranes, water vapor condensation in the detector causes erratic
response, and the presence of 112 SO 4 mist can cause severe corrosion. The detection cells
must be replaced about once per year as the electrolyte solutions become depleted.
A continuous SO 2 monitoring system that uses an electrochemical transducer cell has been
described in the literature (11). The sample conditioning system consists of a ceramic heated
probe through which gas is withdrawn at a rate of 500 em 3 /min through a heated sample
WATER
OVERFLOW
WATER
IN LET
GAS
FLOW
II
RECORDER
FIGURE 17-4
CONTINUOUS CONDUCTIVITY MONITOR FOR MEASURING SULFUR
OXIDE EMISSIONS FROM SULFITE MILL SOURCES (10)
GAS FLOW
3 O I/mm.
PACKED
1tL iER
I—
WATER
rLow
17-12
-------
line. The gas stream is pulled by a leakproof stainless-steel vacuum pump with a Teflon
diaphragm located upstream of the detector because of the necessity of operating the
transducer cells under positive pressure. A refrigerator and condenser are used to remove
water vapor as it is necessary to maintain a strongly acidic medium to avoid absorportion of
SO 2 . The gas stream then flows into the detection cell for subsequent analysis,
amplification, and recording. The system is illustrated in Figure 17-5.
The system has proved successful during extensive field use but requires replacement of the
modular transducer detection cells at intervals ranging from 6 to 18 months, depending on
cell design and SO 2 level in the flue gas stream.
Coulometric titration can be used for monitoring SO 2 emissions from coal- and oil-fired
power boilers and sulfite pulp mill process sources where reduced sulfur compounds are not
present. The technique has been described in section 17.2.2.2. Coulometric titration is
nonselective for SO 2 and so organic olefins and other materials can interfere. It is, therefore,
necessary to add a combustion furnace upstream of the detector in order to oxidize the
olefins from oil-fired boiler flue gases to prevent their interference.
The method also requires removal of water upstream of the detector to av9id flooding the
cell.
Ultraviolet spectrophotometry is useful for measurement of SO 2 stack concentrations either
internally or externally. The method operates on the principle that the degree of ultraviolet
RECORDER L
FIGURE 17-5
ELECTROCHEMICAL TRANSDUCER MEMBRANE CELL FOR CONTINUOUS
SULFUR DIOXIDE MONITORING (11)
SAMPLE DETECTION
CONDITIONER CELL
HEATED
SAMPLE
LI NE
VACUUM
PUMP
WATER
17-13
-------
radiation absorbed at some characteristic wavelength by passage through a gas stream is
proportional to the SO 2 concentration. The method is relatively specific for SO 2 if the
proper wavelength ultraviolet light source is used, and it has only minimal interferences
from water and particulate matter. Three different modes of ultraviolet spectrophotometry
are used in commercially avaitable instrumentation for continuous monitoring of SO 2 levels
in flue gas streams.
Thoen, DeHaas, and Baumgartcl (12) describe the use of an internally located ultraviolet
emission spectrometer for continuous monitoring of SO 2 emissions from a magnesium base
sulfite recovery furnace following the absorption towers. The instrument employs a
detection system where the ultraviolet radiation from a single beam mercury vapor lamp at a
wavelength of 254 nm is passed across a duct through a cylindrical perforated fiber glass
tube 2 m (6 ft) long. Gas molecule interchange into and out of the fiber glass tube is
facilitated by a series of holes located at 90 degrees to the direction of flow; this orientation
inhibits large particles and water droplets from entering. The system is illustrated in Figure
17-6.
The system is suitable for measuring SO 2 concentrations from 80 to 4,000 ppm by volume
with an electronic output of 0 to ]0 volts DC. The system has a minimum of interferences,
is resistant to corrosion, and employs no moving parts. The device is relatively insensitive to
low SO 2 concentrations, displays sluggish response to rapid changes in concentration, and
the fiber glass is not suitable for high gas temperatures. The mercury vapor lamp at 254 nm
wavelength does not correspond to the SO 2 maximum absorbanec at 280 nm.
Control Recorder
::4 1::4L i I
Module 0-by D.C.
FIGURE 17-6
INTERNALLY LOCATED ULTRAVIOLET SPECTROMETER FOR
SULFUA DIOXIDE MONITORING IN FLUE GAS STREAMS (12)
Light
Source
Detector
Cel I
17-14
-------
Saltzman (13) describes the usc of an externally located dual beam ultraviolet
spectrophotometric analyzer for continuous monitoring of SO 2 levels in flue gas streams
from coal- and oil-fired power boilers and sulfite recovery furnaces. The gas stream is
withdrawn from the duct through a heated porous ceramic probe of 20 p (7.9 X iO in)
porosity to remove particulate matter, as shown in Figure 17-7. The gas sample passes
through a heated, electrically traced Teflon tube to prevent condensation of water vapor.
The gas stream then passes through the photometric detection cell at a rate of 1 1/mm and is
drawn through an air aspirator maintained at a constant vacuum. A compressed air purge
located upstream of the detection cell is used to remove particulate matter from the probe.
The detection system is a dual beam photometer in which ultraviolet light of 280 nm is
passed through the sample cell to provide for specific SO 2 absorbance. Visible light, with a
wavelength of 578 nm, is passed through the reference cell so that it is possible to minimize
the potential interference from NO 2 . The system is heated to prevent water condensation, is
rugged and durable under field conditions, and can maintain calibration on a stable basis for
extended periods. To prevent small particles from depositing on the detector cell surfaces, a
glass wool filter is used in the sample line to remove them (14).
FIGURE 17-7
INTERNALLY LOCATED ULTRAVIOLET SPECTROPHOTOMETER FOR
SULFUR DIOXIDE MONITORING IN FLUE GAS STREAMS (11)
Recorder
Liquid
Trap
Reference
Cell
¶
Flue
Gas
Air
Aspirator
17-15
-------
An ultraviolet correlation spectrometer located internally, is used for measuring SO 2 in flue
gas streams (15). The system employs a cylindrical slotted probe placed in the gas stream
perpendicular to the direction of flow. Ultraviolet light at a series of wavelengths
corresponding to the absorbance characteristics of SO 2 are passed from the light source into
the stack and reflected from a mirror back into the detector. The degree of absorption of
light is proportional to the SO 2 concentration in the flue gas and is indicated in ranges from
o to 1,000 or 0 to 5,000 ppm by volume.
1 7.2.4 Nitrogen Oxide Monitoring Systems
Oxides of nitrogen are released to the atmosphere from any combustion process because of
the reaction between 02 and N 2 at elevated temperatures. Approximately 90 percent of the
oxides of nitrogen formed is NO with the remainder being mostly NO 2 . The potential
sources of emissions of nitrogen oxides to the atmosphere are coal., wood., oil-, and
gas-fired power boilers, recovery furnaces and lime kilns in kraft pulp mills, and recovery
furnaces and sulfur burners in sulfite pulp mills.
The possible constituents to be monitored in flue gas streams include NO, NO 2 , and total
oxides of nitrogen. Continuous monitoring of oxides of nitrogen may be performed either
internally within the stack or externally from outside the stack by sample withdrawal.
Oxides of nitrogen emissions from kraft and sulfite pulp mill process sources are not
normally as significant as from power boilers because large quantities of water present in the
spent cooling liquors and lime mud inhibit the occurrence of high flame temperatures.
17.2.4.1 Sample Conditioning
The gas NO is relatively insoluble in water, but NO 2 can be removed during condensation of
water unless strongly acidic conditions are maintained. For certain types of detectors, a
heated, electrically-traced inert Teflon sampling line can be used if condensation is to be
prevented and the detector kept heated. Prime movers employed can be either
corrosion-resistant and leakproof vacuum pumps or air, steam or water aspirators. The
features for sample handling and conditioning systems for oxides of nitrogen measurements
arc similar to those employed for SO 2 systems, as described in the previous section.
17.2.4.2 Detection Systems
The major detection systems employed for continuous oxides of nitrogen measurements
include electrochemieal transducer membrane cells, ultraviolet spectrophotometry, infrared
spectrophotometry, and chemiluminescence. Most detection systems employed for
continuous oxides of nitrogen measurements are located external to the stack and,
therefore, require sample conditioning systems.
17-16
-------
Electrochemical transducer membrane cell detectors arc available for measuring either NO 2
or total oxides of nitrogen (NO plus NO 2 ) levels in flue gas streams. The sample
conditioning system requires withdrawal of the stack gas sample through a ceramic probe
for particulate removal, and a heated line to prevent moisture condensation. The gas is
drawn by a sealed, leakproof stainless steel vacuum pump via the sample conditioning
condenser into the electrochemical membrane detector. The concentration of NO 2 and/or
NO is taken as being proportional to the change in electrochemical potential across the cell
with a readout of 0 to 10 my. Readable concentration ranges arc zero to 500, 1,000, or
5,000 ppm by volume.
Ultraviolet spectrophotometry is useful for measurement of oxides of nitrogen levels in flue
gas streams. The detection principle is the same as for the SO 2 system described (section
17.2.2.2) except that the wavelengths for the light beams to the sample and reference cells
are 436 and 578 nm, respectively, where NO 2 is the chemical compound being measured. A
reactor also converts NO to NO 2 at elevated temperature and pressure. The concentrations
of NO 2 alone and NO 2 plus NO are read sequentially in a timed cycle.
Infrared spectrophotometry can also be used for measuring of oxides of nitrogen. The
principle is the same as for ultraviolet spectrophotometry except that characteristic infrared
absorption peaks for NO and NO 2 are used. There is only limited field experience with the
technique, to date.
Chemilumincscence is another technique for measuring oxides of nitrogen. Stack gas is
withdrawn from the duct, conditioned to remove particulate matter and water vapor, and
then passed to a catalytic chemical reactor to form NO 2 and 02. Light is produced by this
reaction, and the intensity of the light is proportional to the inlet concentration of NO. The
NO concentration alone is determined by letting the sample (containing both NO and NO 2 )
bypass the reduction chamber so that NO 2 is not reduced and, therefore, not detected.
17.2.5 Gas Chromatography
Gas chromatography separates constituents of a gaseous mixture by exploiting their
differences in relative affinity for a given packing material in a concurrent flow column.
These differences cause the various constituents of the mixture to pass through the column
at different rates. The gaseous components can then be individually analyzed as they pass
the column exit by means of a suitable detector. Gas chromatography provides a versatile
means of analyzing for a wide variety of compounds over a wide range of concentrations,
but only for discrete samples and not on a continuous basis.
The major elements of a gas chromatograph are the sample collection and handling system,
the sample injection system, the carrier gas flow system, the separation column, and the
detector. The primary variables in gas chromatography are the sample handling procedure,
17-17
-------
the sample size, the column dimensions, and packing, the column temperature, and the
detector. Compounds of primary interest for measurement by gas chromatography arc H 2 S
and the organic sulfur compounds, plus other organic compounds such as hydrocarbons,
terpenes, and alcohols.
17.2.5.1 Sample Handling
The major difficulties in sample handling and conditioning systems arc posed by the
presence of excessive quantities of particulate matter, water vapor, organic and aqueous
mists and droplets, and pulp fibers (15). A cyclone separator that uses an enlarged sample
probe with a splatter plate pointed downward in the gas stream is necessary to prevent
droplet entrainment in mist-laden gas streams. Porous ceramic or sintered stainless probes or
open tubes packed with glass wool can be used to remove particulate matter upstream of the
sampling lines for gas streams at temperatures above their dew points.
Water vapor is present in flue gas streams from most pulp mill sources in quantities ranging
from 20 Lo 95 percent by volume. The major methods for alleviating possible losses of
gaseous components to be analyzed are to heat the sample probe, sampling lines, and any
collection vessels by electric tracing to temperatures above the gas dew points, or to dilute
the source gas with dry gas of known composition Lo below the level at which water will
condense at the given temperature.
A possibility that must be anticipated and allowed for in the design and construction of gas
sample handling systems is chemical or physical reaction between the gases to be analyzed
and the wall materials of the sample lines or containers. Teflon and glass are the most nearly
inert wall materials readily available for sampling lines, while 316 stainless valves and fittings
are sufficiently inert and corrosion-resistant for normal use. Polyethylene and poly-
propylene can also be used, but are subject to melting at gas temperatures above 120° C
(2500 F).
The two major types of sample handling systems are the batch and continuous types. Batch
systems in use incluclc the use of cylindrical gas collection flasks and evacuated glass bottles
that normally must be heated to 150° C (300° F) or higher to prevent moisture
condensation. Sample injection into the chromatograph normally is made by glass syringes
of varying sizes. it is sometimes necessary to concentrate samples by freezeout, solid
adsorption, or liquid adsorption to have sufficient material to perform analyses.
Continuous sampling systems use withdrawal of the gas sample through a heated line from
the source at a relatively high flow rate. Gas samples can be injected into the chromatograph
at frequent intervals through a sample loop and port assembly. Dilution of a given sample
prior to injection into the chromatograph is sometimes necessary when using detectors, such
17-18
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as the flame photometric unit. Also, to minimize the retention time in the sampling lines,
withdrawal of a considerably larger volume of gas than that passed through the sample is, at
times, recommended.
17.2.5.2 Column Technology
Selecting the proper column and packing is important to successful analysis by gas
chromatography. Pertinent variables in column technology are the column length and
diameter, the solid support, the liquid used, and the tubing material. Selection of the proper
column is necessary to facilitate the separation of the gaseous components of interest;
different gases have different relative affinities for different packing materials.
Gas chromatographic column materials sufficiently inert and temperature resistant for sulfur
gas analyses include 316 stainless, glass, and Teflon (16). The sample injection sy6tem, the
separation column, and the detector must be heated to facilitate many of the sulfur gas
separations. Pressure and temperature limits, inertness, and durability must be considered in
selecting tubular column materials. The degree of separation between components that can
be achieved by a column increases with increasing length and decreasing diameter. Column
diameters normally vary from 0.3 cm (0.125 in) to 0.6 cm (0.250 in) with lengths ranging
from 3 to 30 m (10 to 100 feet). Carrier gas flow rates will vary from 30 to 150 cm 3 /min
(15). Teflon column temperatures should not exceed 100 to 1500 C (212 to 300° F).
Solid phase support materials must be of sufficient inertness, porosity, uniformity, strength,
and ease of packing for generalized use. The separation efficiency of the solid support is
directly proportional to its porosity and surface area, but inversely proportional to its
inertness. Noninert columns result in tailing of peaks. Normal column solid supports include
Chromosorb C, P, T, and W.
The liquid phase of the column separates the various components by either vapor pressure
or polarity; in either case molecules of greater molecular weight tend to remain in the
column for longer periods. Liquid phase materials in common use for sulfur gas separation
include the Carbowax 20X and 1540, polypropylene, glycol, binonylphthalate and
polyphenyl ether and Poropak Q (15).
17.2.5.3 Chromatographic Detectors
The major ehromatographic detection systems in use for sulfur gas analyses are thermal
conductivity, flame ionization, flame photometry, and microcoulometry. The detectors
used must have accuracy, stability, sensitivity, selectivity, durability, rapid response,
minimum maintenance, and freedom from interfering substances. Thermal conductivity
detectors are not specific or sensitive enough for most mill applications, and
17 19
-------
microcoulometric detectors are unsuitable because they can be easily overloaded at high
concentrations and require frequent maintenance.
A summary of detector characteristics is presented in ‘l’able 17-3.
The two major detectors in use in pulp mills include the flame ionization and flame
photometric units. Both systems arc free from water interference, have excellent stability
characteristics, and both require hydrogen flames and a nitrogen carrier gas. The flame
ionization detector is suitable for organic sulfur arid nonsulfur compounds, which can be
ionized in flames, but is not sensitive to the inorganic H 2 S or SO 2 . The flame photometric
detector is suitable for H 2 S and SO 2 , as well as organic sulfur compounds, over the range
from 5 ppb to 5 ppm by volumc (17). Samples of higher concentrations require either
dilution or the use of very small sample loops of 0.5 cm 3 or lower capacity.
17.2.6 Calibration Procedures
Periodic calibration of gaseous monitoring instruments is necessary to their continued
accuracy. The typical calibration procedure checks the instrument by using it to measure a
gas stream of known concentration. The difference between the known and the measured
values is the error of the instrument. Methods in use for preparing known gas concentrations
include rotating syringes, motor-driven syringes, flexible fabric bags, known cylinder
mixtures, and permeation tubes. Of these, all except the motor-driven syringes are in
common use.
17.2.6.1 Rotating Syringe
Rotating syringes can be used for calibration of gaseous monitoring instruments over both
ambient and source ranges of concentration. The technique employs dilution of pure gas
from a small syringe with air in a large syringe. The large syringe is placed in an upright
position and caused to rotate by the action of an airstream directed against its vanes. The
flow rate of gas mixture from the syringe is controlled by a calibrated limiting flow
capillary, which is usually a broken thermometer. The capillary is inserted into the air
stream to form the required gas concentration, which is then fed to the instrument, as
shown in Figure 17-8 (18).
The small syringes can range in size from 0.5 to 10.0 cm 3 while the large syringes can be 50
to 100 cm 3 . Adding pure gas to the large syringe at extremely high inlet concentrations is
sometimes necessary. The thermometer capillaries are individually calibrated by a soap
bubble flow meter; the flow rate remains constant because of the constant pressure exerted
by the rotating plunger. Flow rates for the capillaries normally range from 0.5 to
5.0 cm 3 1mm. Air flow rates for the dilution system can vary from 0.5 to 50 11mm (0.018 to
1.8 cfm).
17-20
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TABLE 17-3
OPERATING CHARACTERISTICS OF GAS CHROMATOGRAPHIC DETECTORS (15)
Gases Analyzed
Organ. Organ. Water
Detector Sensitivity Stability H 2 S SO 2 Sulfur Comp. Interference
ppm, by vol
Thermal Conductivity 10 Good Yes Yes Yes Yes Ycs
Flame Ionization 0.5 Excellent No No Yes Yes No
Flame Photometric 0.005 Excellent Yes Yes Yes No No
Bromine Coulometric 0.5 Poor Yes Yes Yes Yes* No
For oxidizable organic compounds only.
-------
The rotating syringe method provides a versatile and inexpensive means of providing gas
concentrations over a wide range with a minimum of equipment. But the method is not
without disadvantages. It does require prior calibration of the capillary flow rate. T he
syringes are subject to sticking in humid atmospheres. Careful loading is necessary to avoid
errors. Arid it is necessary to rotate the syringe plunger by directing an air stream against its
attached vertical rotor; otherwise there is no assurance that a constant delivery pressure is
maintained, which is necessary to assure a constant gas flow rate through the syringe
capillary.
Permeation tubes provide a versatile and accurate means of calibrating gaseous monitoring
instruments. The technique operates on the principle of a constant rate of diffusion of a
purc gas through a porous membrane of fixed cross-sectional area and thickness at a given
temperature (19). The permeation tube is placed in a chamber immersed in a constant
temperature bath, aiid dilution air is caused to pass through the system. The gas stream
containing the dilution air plus the pollutant gas at a specified concentration is then caused
to flow from the dilution chamber to the instrument for calibration purposes, as shown in
Figure 17-9.
Permeation tubes are made of cylindrical lengths of Teflon plastic filled with liquefied
pollutant gas to be measured. They can be fabricated or purchased commercially (20). Thc
constant temperature bath is normally kept at 25 to 27° C (77 to 80° F). The physical
FIGURE 17-8
ROTATING SYRINGE INSTRUMENT CALIBRATION PROCEDURE
To
Exhaust
Rototing
Syringe
To
Instrument
17-22
-------
Constant
Temp Bath
To Instrument
FIGURE 17-9
PERMEATION TUBE INSTRUMENT CALIBRATION PROCEDURE
To Exhaust
variables for the permeation tube are its diameter, wall thickness, and length. The primary
operating variables are the temperature and the air flow rate. Permeation rates for tubes arc
normally calibrated gravimetrically in terms of weight loss pcr unit time.
Permeation tubes require the use of a constant temperature bath and are more cumbersome
for field usc and more expensive than rotating syringe systems. Newly developed systems
now available commercially eliminate the use of liquid baths, resulting in a lighter weight,
more compact and less complex unit.
17.2.6.3 Gas Cylinders
Stainless gas cylinders used for calibration of gaseous monitoring instruments are cylinders
of given volume. Known amounts of gas are inserted in them and pressurized with either air
or nitrogen to produce known concentrations (21). The gas mixture can then be added at a
given flow rate to a stream of flowing air to produce a given concentration. The technique is
suitable for field use and is simple and inexpensive; however, sevcre losses can occur during
storage for certain gases by adsorption on or reaction with wall materials.
17.2.6.4 Calculation Procedures
1. Rotating Syringe
CLQL VsQL
QI VLQI
where: C 1 concentration fed to instrument, ppm by volume
CL concentration in large syringe, ppm by volume
Q 1 flow rate of dilution air, cm 3 /min
Dilution
Air
Chamber
17-23
-------
Q L = flow rate from large syringe, cm 3 /min
= volume of large syringe, cm 3
= volume of small syringe, cm 3
Note.—Because 1 ’ QL VL and V appear only in ratios, any consistent units of
volume and volume/time may be used.
2. Permeation Tube (20)
C —00623 ( R)(T )
— (MW) (Q 1 ) (F)
where: C 1 = concentration fed to instrument, ppm by volume
MW = molecular weight of gas, g/mole
P = pressure of calibration system, mm Hg
= dilution air flow rate, 1/mm
R = permeation rate, ng/min
T = water bath temperature, °K
17.3 Particulate Monitoring
Particulate sampling and analysis are necessary for recovery furnaces, smelt tanks, and lime
kilns in kraft pulp mills, for recovery furnaces in sulfite pulp mills, and for all coal-, wood-,
and oil-fired power boilers. Determination of particulate mass concentration is necessary to
meet current air pollution regulations, with increasing emphasis being placed on particle size
distribution. Batch techniques are specified for determining particulate mass concentrations,
emission rates, and size distributions. Continuous monitoring of particulate emissions is
gaining increasing emphasis as more accurate and reliable methods are developed (22).
17.3.1 Preliminary Considerations
Particulate matter is normally defined as any material emitted into the atmosphere in either
a solid or liquid state, including dusts, fumes, smoke, flyash, soot, tars, droplets, and mists.
Changes in physical state with temperature can cause confusion in defining particulate
matter for materials such as organic vapors and acid mists. Changes in chemical form, such
as oxidation of 502 to H 2 SO 4 in sampling train impingers after collection, can cause
especially serious difficulty in interpreting emission standards.
Particulate matter must be defined as such for either stack or standard conditions. The
present definition accepted by the U.S. Environmental Protection Agency is that particulate
matter is material collected on a filter of porosity 0.45 pm (1.8 X 1W 5 in) which has been
17-24
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heated to 1210 C (2500 F). Other definitions for particulate matter specified by local
agencics include, among others, material collected in liquid impingers at standard
conditions.
Particles in pulp and paper mill sources can vary from less than 0.01 to greater than 100 pm
(4 X 10 — 1 to 4 X iO in) in diameter, depending on the source and type of collector used.
Sampling at isokinetic conditions is normally necessary because of the inertial properties of
particles greater than 5 pm (2 X iO in) in diameter. The normal procedure employed for
maintaining isokinetic conditions within plus or minus 10 percent of the actual stack
velocity is to locate a pitot tube parallel to the sample probe to measure the gas velocity
continuously and to make periodic corrections as needed.
17.3.2 Batch Particulate Sampling
Batch particulate sampling is used to determine total particulate concentrations and
emission rates from flue gas streams and is necessary for calibration of continuous
particulate monitoring devices. Batch sampling provides an average value for particulate
concentration in a duct over a given time period, but does not provide real-time
instantaneous values. Collection methods for batch particulate sampling include filtration
and liquid impingement.
A low volume system (0.8 to 2.5 m 3 /h or 0.5 to 1.5 cfm), specified by the American
Society of Mechanical Engineers (ASME), relies on an internally located alundum thimble
for particulate collection, as shown in Figure 17.10 (23). The amount of material collected
is then analyzed gravimetrically. The standard specifics collection of particulate materials
larger than 1 pm (4 X jØ5 in) in diameter. The method is relatively simple and inexpensive,
but has several shortcomings. The thimbles tend to pass small particles, particularly during
the initial collection period, in direct proportion to increasing thimble porosity. The
thimbles are subject to leakage by improper sealing around their gaskets, to plugging and
washlhrough in wet gas streams, and to dust losses during handling following collection.
The Los Angeles County Air Pollution Control District employs a multiple.stage collection
train for particulate sampling in an arrangement similar to that employed by the National
Council for Air and Stream improvement (24). One sampling train employs an internally
located heated cyclone and alundum thimble as the primary collection stage, followed by a
series of liquid impingers, and a five grade thimble for removal of particles of diameter
greater than 0.3 pm (1.2 X 10—6 in).
The vacuum pump is placed downstream of the meter to preclude air dilution of the gas
stream. A sampling rate of 0.8 to 1.5 m 3 /h (0.5 to 0.9 cfm) is employed. Total particulate
matter is then determined by gravimetric analysis of the material collected. The method
requires several analyses, and the possibility of sulfate formation from oxidation of SO 2 in
the impinger liquid following collection introduces further uncertainty.
17.25
-------
ALUN DUM
THIMBLE
¶
II II
CONDENSER
PUMP
GAS
METER
FIGURE 17-10
ASME BATCH PARTICULATE SAMPLING TRAIN
The particulate sampling train, specified by the U.S. Environmental Protection Agency,
employs an externally located two-stage collection system as shown in Figure 17-11 (25).
The primary collection stage consists of a glass cyclone and a 5 cm (2 in) diameter glass fiber
filter of 0.30 to 0.45 pm (1.26 to 1.8 X 10—6 in) porosity that is heated to 1210 C (250° F)
to avoid condensation. The secondary collection stage consists of a series of
Greenburg-Smith liquid impingers containing water. The impingers are used to condense
water vapor present so as to prevent flooding the pump and meter and for absorption of
potentially corrosive gases. In addition, the water vapor content of the stack gases is
determined from the condensate collected in the impinger. The vacuum pump follows the
FIGURE 17-11
EPA BATCH PARTICULATE SAMPLING TRAIN
17-26
-------
impingers and is followed by the gas meter to measure the volume sampled. An S-type pitot
tube, located in parallel to the sample probe, facilitates the maintenance of isokinetic
sampling conditions. The system provides for efficient particulate collection, but is
cumbersome to handle. The small filters are subject to rapid plugging at high loadings, and
large pressure drops tend to develop after prolonged sampling periods.
17.3.3 Particle Size Distribution
A major problem in particulate sampling is the determination of particle size distribution in
flue gas streams. Particles can range in diameter from less than 0.01 to greater than 100 pm
(4 X iO to 4 X i0 in), with densities ranging approximately from 0.8 to 2.5 g/cm 3 .
The problem of determining particle size in gas streams is made especially difficult by the
high humidity of the gas, the presence of water droplets, and the tendency of particles to
agglomerate or coalesce. The major methods employed for particle size determinations in
emissions from pulp and paper mill sources are multistage cascade impaction and membrane
filtration.
Bosch, Pilat, and Hrutfiord (26) describe the use of a multistage cascade impactor for
particle size determination from kraft recovery furnaces over the range 0.5 to 20 pm
(2 X iO to 8 X 1ff 4 in) in diameter. The system is an internally located six-stage cascade
impactor in which particles of progressively smaller diameter are collected on successive
silicone-coated plates by passage through multihole plates with progressively smaller
diameter holes. Sampling times vary from 30 seconds to 30 minutes, depending on the
particulate loading of the source being measured.
The system must be calibrated for a given dust at a certain flow rate prior to collection to
determine the approximate mean particle diameters for individual stages. Particle size
determinations are made by gravimetric weighings of individual plates; the system of
classification of particles is as percentage by weight within given sizes. Particles smaller than
0.5 pm (2 X io in) in diameter are collected on a filter following the impaction plates, as
shown in Figure 17-12.
The National Council for Air and Stream Improvement has developed a method for
collection of particles on a membrane filter for subsequent visual counting and sizing by
means of an optical microscope (27). The system employs a parallel flow heated collection
train with the membrane filter on one branch and a valving system on the other, as shown in
Figure 17-13. The system must be hydrostatically balanced so that there is the same
resistance to flow in both branches. It is then placed in the stack and heated to 93° C
(20L1° F). This temperature is high enough to prevent condensation, but not so high as to
damage the filter. The gas stream is first directed through the bypass system and then
through the filter for 20 to 30 seconds to collect the particles. The technique is suitable
only for sizing particles on a count basis, which is a time consuming procedure, and does not
17-27
-------
I CASCADE IMPACTOR ASSEMBLY
2 WATER BATH
3 IMPINGERS
4 THERMOMETER
5 DRY GAS METER
6 VtCUUM GAUGE
7 VACUUM PUMP
8 GAS FLOW RATE REGULATOR
FIGURE 17-12
7
MULTISTAGE CASCADE IMPACTOR FOR PARTICLE SIZE
DISTRIBUTION DETERMINATION (26)
FIGURE 17-13
MEMBRANE FILTER SYSTEM FOR PARTICLE SIZE
DISTRIBUTION DETERMINATION (27)
8
‘I
17-28
-------
allow determination of particle size distribution on a weight basis. But, the sample
collection can be done in a short time interval, does not require the extensive precalibration
required for the cascade impactor devices, and does not require prior knowledge of the dust
density characteristics.
11.3.4 Continuous Monitoring
Continuous particulate monitoring is gaining increasing emphasis for determination of stack
concentrations and emission rates. Methods now in use for particulate monitoring in the
pulp and paper industry include wet chemical techniques, such as conductivity and specific
ion determinations, and optical bolometry and transmissometry. Techniques which may
become applicable in the future, include beta ray attenuation, piezoelectric crystallography,
electronic sensing, optical nephelometry, holography, electric ion capture, and optical lasers,
lidar, and radar (28).
17.3.4.1 Chemical Methods
Leonard (29) describes a wet chemical system for determining particulate emissions from a
kraft recovery furnace on a continuous basis. The system operates on the principle that the
increase in electrical conductivity of a liquid stream caused by the presence of sulfate ion is
proportional to the total particulate loading, which is primarily Na 2 SO 4 . Flue gas is
withdrawn from the duct at a predetermined average isokinetic velocity into a probe and
contacted concurrently with deionized water, as shown in Figure 17-14. The soluble
particulate matter is passed through a detection cell containing a conductivity probe and the
_______________ i i ___________________
DEIONIZED I I
[ ff]- j: [ fl [ 1 f l
VELOCITY CONCENTRATIOh
READOUT READOUT
FIGURE 17-14
CONTINUOUS MONITORING OF PARTICULATE EMISSIONS
WITH A CONDUCTIVITY CELL DETECTOR (29)
17-29
-------
gas is drawn off by a vacuum pump. The method is simple and inexpensive, but is
nonspecific because any conductivity producing substance can cause interference, e.g., SO 2
or CO 2 from the flue gas.
Tretter has developed a modification to the conductivity method for measuring particulate
concentrations which alleviates conductivity interferences caused by the presence of SO 2
and CO 2 in the flue gas (30). The method uses a sodium ion-specific electrode to measure
the sodium content of a water stream as an indicator of total particulate concentration. The
system withdraws the gas sample through a probe and allows the absorption into water in a
concurrent flow condenser by condensation of compressed steam, as shown in Figure 17-15.
The liquid stream is then directed to a detection cell containing the sodium electrode where
a liquid pH of 8.5 to 9.5 is maintained for maximum sensitivity. The system requires
periodic maintenance, and, for calibration, a correlation between sodium ion level and total
particulate concentration must be developed by using parallel batch tests.
17.3.4.2 Optical Determinations
Optical devices in use for particulate monitoring include internally located bolometers and
transmissometers, where the degree of light attenuation is a function of the particulate
concentration in the duct. The system directs a light beam across the duct to a detector
where the resulting electrical signal is amplified, transmitted, and printed out on a recorder.
FIGURE 17-15
CONTINUOUS MONITORING OF PARTICULATE EMISSIONS WITH A
• SODIUM ION-SPECIFIC ELECTRODE (30) -
COMPRESSED STEAM
EXHAUSTGA$
c5 ji
DETECTION TO DRAIN
CELL
17-30
-------
The sensitivity and accuracy of optical stack monitoring devices i affected by path length,
intensity, and wavelength of the light beam, moisture content as a function of temperature
of the stack gas, particle size distribution, and particle mass concentration of the flue gas
stream. (31).
Gansler describes the use of a bolometer for measuring particulate emissions from a kraft
recovery furnace flue gas (32). The system uses a tungsten lamp with an optical band from
approximately 500 to 2,500 nm. The device is suitable for measuring particle concentrations
at levels below 29/rn 3 (0.9 gr/cu ft), and provides warnings of precipitator malfunctions.
The device requires extensive prior calibration by parallel batch tests and also frequent lens
cleaning.
Beutner describes the use of optical transmissometers for continuous particulate monitoring
at a number of installations (33). The system employs a fixed length light beam with a
wavelength of 300 to 800 nm, about the visible region of the spectrum. The gas stream
passes between the beam and the detector. Fans are used to blow sufficient air across the
surface to keep the lens surfaces clean. The device can be either permanently mounted or
portable. The system is effective as an indication of particulate matter primarily of a size
range of 0.1.1.0 pm (4 to 40 X 10—6 in). The system must be calibrated for the specific
source by an extensive series of batch tests.
Optical particulate measurement devices are suitable for application to kraft recovery
furnaces, coal- and wood-fired power boilers, and kraft lime kilns, and possibly ammonium
and magnesium-base sulfite recovery furnaces.
They are particularly useful for providing warnings of possible particulate emission control
equipment malfunctions and as indicators of combustion unit efficiencies for power boilers.
They are also useful for emission monitoring for possible compliance with stack opacity
standards, because optical transmittance is the property they measure. Their use is limited
for mass emission monitoring because long-term correlations with batch tests must be
determined under stable operating conditions. The lenses of the detectors tend to become
obscure so that they can require frequent cleaning, and changes in particle size distribution
can alter readings. The devices are probably most suitable for stacks with low particulate
concentrations following high efficiency control devices where the particle size distribution
is relatively uniform.
17.4 Odor Measurements
Measurements of odors is a major problem in the pulp and paper industry. The major
chemical components that may cause a community odor nuisance are reduced sulfur
compounds, such as H 2 S, CH 3 SH, CH 3 SCH 3 , and CH 3 SSCH 3 . These malodorous sulfur
compounds and, possibly, other organic compounds are emitted from the digesters,
17-3 1
-------
evaporators, recovery furnace, smelt tank, lime kiln, and wastewatcr streams in kraft pulp
mills. Under certain circumstances, SO 2 and organic sulfur compounds from sulfite pulp
mills can also cause a community odor nuisance.
Odor measurement is an extremely complex task because of variations in the types and
amounts of odorous gases present, variations in response between individuals, and variations
within an individual with time because of hi physical condition, smoking history, time of
exposure, and prior history of exposures. Meteorological variables, such as temperature,
humidity, wind velocity, and turbulence, can affect odor responses of individuals (34).
Odorous gas molecules that are adsorbed on the surfaces of particles can travel for longer
distances in a concentrated form and bring sharper responses than the odorous gas molecules
alone (35).
17.4.1 Threshold Levels
The odor threshold levels for most malodorous sulfur compounds emitted from kraft pulp
mill sources arc between 1 and 10 ppb by volume, as shown in Table 17-4 (36):
TABLE 17-4
ODOR THRESHOLD LEVELS FOR
MALODOROUS SULFUR COMPOUNDS (36)
Sulfur Compound Threshold
ppm, by vol
H 2 S 4
CH 3 SH 2
CH 3 SCH 3 4
CH 3 SSCH 3 6
SO 2 3,000
The odorous gas levels for flue gas streams can be anywhere from 1,000 to 100,000 times as
high as those observed at distances of up to 10 km (6 miles) from the kraft pulp mill.
The evaluation of odors can be made in terms of their character, intensity, pervasiveness,
and acceptability. The intensity of an odor tends to increase with the logarithm of its
concentration according to the Weber.Fechner law (37). Odor intensity levels can be rated
on an arbitrary scale in values ranging from one to five, as shown in Table 17-5 (35).
17-32
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TABLE 17-5
ODOR INTENSITY LEVEL EVALUA-
TION SCALE
Arbitrary Intensity
N umber Level
0 No Odor
1 Detectable
2 Faint
3 Noticeable
4 Strong
5 Overcoming
These levels can be used as refercnce scales for people on a panel in an attcmpt to provide
some quantitative scale of odors.
17.4.2 Design Considerations
Major considerations in performing an odor threshold evaluation include thc selection of
human subjects, the sample collection methods, and the dilution techniques employed. The
two general approaches to odor level measurement are:
1. Direct subjective organoleptic determination of odor threshold levels, and
2. Nonselective chemical determination of individual odorant concentrations by gas
chromatography or other means.
Odor level measurements require careful collection and careful evaluation of samples.
An odor panel is normally selected as a group of people who have particular sensitivity,
accuracy, speed, and reproducibility for evaluating the odor threshold and intensity levels
for the gases being measured. The odor panel goes through a training program for a specific
period, and then is used to test stack gas mixtures. Panelists may undergo testing for 15
minutes at a time with rest periods of at least 30 minutes to avoid olfactory fatigue. It is
normally best to expose panelists to low concentrations before high concentrations to avoid
deadening their response. Prince and Ince find that responses can vary as much as 40 percent
for individual panelists on any given day and 20 percent for a panel of observers (38).
Gas sample collection may be by direct piping of the flue gas to a panel to minimize
potential losses, but this may not always be feasible. Samples of stack gas or ambient air
17-33
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may also be collected in plastic bags, heated evacuated bottles, or in syringes. Odorous gas
losses by solution in condensed water vapor, or physical absorption or chemical reaction
with container walls, are to be eliminated normally by the use of inert wall materials such as
Teflon.
1 7.4.3 Dilution Techniques
Two publications by the National Council for Air and Stream Improvement present
extensive discussions of systems for dilution of odorous gases to facilitate evaluation of
odorant intensity levels (39) (40). The major techniques for evaluation of odor threshold
levels are:
1. Static progressive dilution using glass syringes,
2. Dynamic dilution of odorous gases using odor-free air,
3. Dilution of odorous air with odor-free air by respiration, and
4. Vaporization of odorous compounds in a continuous flow system.
The American Society for Testiiig and Materials specifies progressive static dilutions in glass
syringes. The same is first collected in a 100 cm 3 glass syringe (41). The flue gas sample then
is diluted in the laboratory by adding aliquots of known volume to other syringes of this
same size. The aliquots then are given to an odor panel in order of decreasing dilution
(increasing odorant concentration) to evaluate odor threshold levels and odorant intensities.
The syringes must be cleared after use and dried to avoid contamination of future samples.
Several systems arc available for dynamic dilution of odorous gases by means of odor-free
air for subsequent observation by a panel. Normally, the odorous gas sample must be
collected in some type of container before being added to the dilution system in a
laboratory. Field exposure of subjects prior to odor level evaluations can lead to olfactory
fatigue and resultant insensitivity to odorous gas concentrations. Cederlof (42) describes a
system in which odorous gas samples are collected from kraft mill flue gases in flexible
plastic bags and returned to the laboratory where they are diluted with purified air for
exposure to a panel inside a hood, as shown in Figure 17-16. The inlet odorant gas flow is
controlled at a constant rate by a limiting flow capillary, while the air flow rate is measured
by a rotometer.
Direct observations of odor levels in the atmosphere or from flue gas streams by means of a
progressive dilution device known as a “Scentometer” have been made. The device is
powered by the observer’s respiration. The system draws in variable amounts of odorous gas
or air through an adjustable orifice followed by dilution with air which has been purified by
17-34
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ODOR
PAN EL
FIGURE 17-16
SWEDISH DYNAMIC DILUTION SYSTEM FOR ODOR LEVEL EVALUATION (42)
activated carbon, as shown in Figure 17-17 (43). The relative amounts of odorous gas and
purified air are governed by a series of four orifices which provide for dilution factors of 2,
7, 31, and 170 times, respectively. The procedure starts from the greatest degree of dilution
and works progressively toward lesser dilution until the odor threshold is observed.
17.5 Mobile Laboratories
Mobile laboratories provide the capability of making comprehensive evaluations of ambient
and source concentrations for gaseous and particulate materials in the field. Mobile
laboratories used for air pollution studies should be equipped with instrumentation for
continuous measurement of sulfur compounds, a gas chromatograph for analyses of specific
compounds, analytical equipment for additional constituents, such as oxides of nitrogen,
and particulate sampling equipment.
Waither and Amberg describe a mobile laboratory mounted in a van that has the capability
of monitoring malodorous sulfur compounds. It uses two gas chromatographs in parallel
equipped with thermal conductivity and flame ionization detectors, respectively (44). The
sample handling system consists of a heated ceramic probe and a Teflon electrically traced
DIAPHRAGM
PUMP
CAUBRATED
CAPILLARY TUBES
17-35
-------
NOSE PIECES
CHARCOAL
BED
PURIFIED PURIF IED
AIR FOR AIR FOR
DILUTiON DILUTION
GRADUATED
SERIES OF
ORIFICES
ODOROUS AIR
FIGURE 17-17
SCENTOMETER DILUTION SYSTEM FOR ODOR
THRESHOLD EVALUATION (43)
sampling line for carrying the flue gas from the stack to the instruments. The van proved too
small and had to be placed in a permanent location, thus negating its mobility.
Mulik, Stevens, and Baumgardner report on a mobile laboratory mounted in a trailer which
employed flame photometric detectors for measuring reduced sulfur compounds (45).
Samples were taken at a rate of 50 1/mm (1.8 cfm) from a source gas through a 0.63 cm
diameter and 76 m long (0.25 inch by 250 ft) electrically traced Teflon sampling line main-
tained at 182° C (3600 F) to prevent moisture condensation. The gas chromatograph em-
ployed a 10-port sample valve with a 10 cm 3 (0.6 cu in) sample loop actuated on a 10-
minute sequence for sample injection. A dynamic dilution system was used for sample dilu-
lion factors covering a range from 10 to 1,000,000 to 1, depending on the requirements
for optimum sensitivity with the flame photometric detector.
17.5.1 NCASI System
The National Council for Air and Stream Improvement employs a trailer for mobile emission
sampling of kraft pulp mills (46). The sample system consists of a heated ceramic sample
probe, a heated, electrically traced Teflon sampling line, and a heated conditioning box for
17-36
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dilution of the sample, as shown in Figure 17-18. The gas is withdrawn at a greater rate than
that required for the instruments in order to minimize retention time in the sample lines.
The instrumentation consists of two coulometric titrators with a furnace for sulfur gas
analysis and a gas chromatograph equipped with flame ionization arid flame photometric
detectors. Total cost of the entire system when constructed was approximately $30,000,
and it required two men for its operation.
The sampling handling system employs withdrawal of the sample gas at a rate of
approximately 1 1/mm (0.035 cfm), through a probe heated to 1500 C (300° F) to evaporate
water droplets to alleviate droplet entrainment in the sampling lines. The electrically traced
0.63 cm (0.25 inch) diameter Teflon sampling line 33 or 66 m (100 or 200 ft) length is
heated to 110 to 121° C (230 to 250° F) to prevent moisture condensation. The gas stream
then passes into a heated sample conditioning box, where portions of the gas stream are bled
off to either the continuous sulfur monitors or the gas chromatograph, while the remainder
passes to a large glass carboy acting as a condensate trap to avoid damage to the pump. The
sample gas stream leading to continuous sulfur analyzers can be diluted by factors from zero
to 25.
FIGURE 17-18
GASEOUS SAMPLING SYSTEM FOR SULFUR GAS ANALYSIS
IN NCASI MOBILE LABORATORY (46)
fJ
AM’UFIER RECORDER
SAMPLE CONDITIONING
BOX
GAS FI.OW —
ELECTRICAL -- --
qOMATOGRAPHIC DETEC
I. FLAME IONIZATION
2. FLAME P*IOTOMETRIC
GAS FLOW RATES
SAMPLE FLOW 10-100 ML/MIN.
DILUTION AIR 150-240
Q INSTRLIMENT FLOW 250
TOTAL FLOW 1000
17-37
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The continuous sulfur analysis system either splits the gas flow into two parallel streams for
simultaneous measurement of both total sulfur and TRS, or passes a single stream through
an 502 selective scrubber containing either KHC 3 H 4 O 4 or citric acid (C 6 H 8 0 7 ). The gas
then passes through parallel quartz combustion furnaces heated to 7600 C (1,4000 F) for
oxidation of the reduced sulfur compounds to SO 2 , which then is measured by means of
parallel bromine coulometric titrators. The difference between the signals for total sulfur
(502 + 1 - 125 + organic sulfur) and TRS (1-12 S + organic sulfur) is taken as being proportional
to the SO 2 concentration of the flue gas stream. Thc separation and identification of
specific sulfur compounds is made by periodic injection of samples of given volume through
the sample loop. The sample size is governed by the volume of the sample loop. The sample
gas then passes through a separation volume for analysis by either flame ionization or flame
photometric detection.
17.5.2 ltayonicr System
Waddington describes the mobile laboratory constructed by ITT.Rayonier, Inc., for
monitoring paper mill emissions and ambient air from a 10.7 m (35 foot) truck trailer (47).
The system has both particulate and gaseous sampling trains and sample handling systems, as
shown in Figure 17-19. The particulate systems consist of an EPA train for total particulate
analyses and an Andersen sampler for particle size determination. The gaseous sampling
system consists of parallel heated sampling lines and a sample conditioning system to
provide for dilutions with air by factors of 10 to 10,000. The gaseous instrumentation
consists of continuous monitors for reduced sulfur compounds, °2, CO, oxides of
nitrogen, and hydrocarbons. The gas chromatograph is equipped with flame ionization and
flame photometric detectors.
The sample handling system for source gas analyses employs withdrawal of the sample gas
through heated filters located outside the stack through an electrically traced heated Teflon
sampling line of 0.95 em (3/8 in) diameter 33 or 66 m (100 or 200 ft) in length. The gas
stream passes through at a rate of 100 to 160 1 /mm (3.5 to 5.7 cfm) and small portions then
pass out into a series of as many as three consecutive dilution stages in a heated sample
conditioning box for dilutions from 15 to 1,500 to one. Portions of the gas streams,
including SO 2 , NOx, CO, and TRS, are routed through three separate channels to several
detectors for measuring individual gaseous products.
Ambient monitors are used for SO 2 and NO 2 . The gas chromatograph employs parallel
flame photometric and flame ionization detectors with a 10-port sampling valve for sample
injection.
For particulate sampling, two parallel EPA-type sampling trains can be used on the inlets
and outlets of particulate emission control devices. Either can be fitted with Andersen
cascade impactors that can be internally or externally located for particle size
determinations. Ambient particulate measurements can be made with several high volume
17-38
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AMBIENT SAMPLING TOWER
(IS- 45 EXTENTION)
HI - VOLUME
CASCADE IM CTOR
AMBIENT TEFLON
SAMPLE LINES
FIGURE 17-19
PARTICULATE AND GASEOUS SAMPLE HANDLING SYSTEMS
FOR ITT-RAYONIER MOBILE LABORATORY (47)
suspended particulate samplers carried in the van. A weather station is also carried in the van
for meteorological studies of wind speed, wind direction, and air temperatures. Total cost of
the entire van was about $250,000, and its operation requires three men.
17.6 Economics
An important factor in the design and operation of continuous stack monitoring systems is
the cost of the necessary equipment. Auxiliary equipment, such as recorders, sampling lines,
fittings, and other appurtenances, must be included in the total cost. Failure to properly
account for the auxiliaries can cause variations in prices between different systems.
Appriximate ranges in capital cost for gaseous monitoring equipment are listed in Table
17-6, and for particulate monitoring equipment in Table 17-7. The figures listed are
expressed as approximate ranges only; exact figures must be determined by direct
quotations from specific manufacturers.
CONT
CLIMET kIETEROLOGICAL
SENSORS S AMBIENT
SAMPLE FILTERED
INTAKE (GLASPAC)
CALIBRATION GAS TO TIP OF GAS
SAMPUNG UMBILICAL
17-39
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TABLE 17-6
APPROXIMATE CAPITAL COSTS FOR CONTINUOUS GASEOUS STACK
MONITORING INSTRUMENTATION
Instrument Pollutants Capital
Type Measured Cost
dollars
Electrolytic Conductivity SOx 1,000-2,000
Ultraviolet Spectrophotometry 50 x NOR, TRS 2,500-10,000
Electrochernjcal Transducer SO, , NOR, CO, TRS 3,000-9,000
Coulometric Titration SO, , TRS 5,000-8,000
Flame Ionization Organic 2,000-7,000
Chemiluminescence NO 5,000-7,000
Flame Photometry SOx, TRS 4,000-8,000
TABLE 17-7
APPROXIMATE CAPITAL COSTS FOR CONTINUOUS PARTICULATE STACK
MONITORING INSTRUMENTATION
Instrument Detection Capital
Type Principle Cost
dollars
Optical Bolometer Optical 2,000-3,000
Optical Transmissometer Optical 5,000-7,000
Electronic Sensing Electrical 6,000-7,000
Beta-Ray Attenuation Radiation 10,000-20,000
17-40
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17.7 References
I. Cooper, H. B. H., and Rossano, A. T., Source Testing for Air Pollution ControL Wilton,
Connecticut. Environmental Science Service Corp., 1971.
2. Blosser, R. 0., Cooper, I-I. B. I-I., and Megy, J. A., Gaseous Emissions—Automatic
Techniques—Electrolytic Titration. NCASI Atmospheric Pollution Technical Bulletin
No. 38, National Council for Air and Stream Improvement, New York, New York,
December 1968.
3. Thoen, C. N., DcHaas, C. C., and Austin, R. R., Continuous Measurement of Sulfur
Compounds and their Relationship to Operating Kraft Mill Black Liquor Furnaces.
Tappi, 52:1485- [ 487, August 1969.
4. Thoen, G. N., DeHaas, C. C., and Austin, R. R., Instrumentation for Quantitative
Measurement of Sulfur Compounds in Kraft Gases. Tappi, 51:246.248, June 1969.
5. Canfield, J., Measurement of Odors and Sulfur Compounds. (Presented at the 12th
Conference on Methods in Air Pollution and Industrial Hygiene Studies, Los Angeles,
California, April 7, 1971.)
6. Altshuller, A. P., and Sleva, S. F., Vapor Phase Determination of Olefins by
Coulometric Method. Analytical Chemistry, 34:418-422, March 1962.
7. Austin, R. R., Sampling and Analysis of Pulp Mill Gases for Sulfur Compounds. Tappi,
54:977-980, June 1971.
8. The Barton Model 286 Sulfur Titrator. ITT Barton Instrument Co., Monterey Park,
California, 1967.
9. Cooper, H. B. H., and Rossano, A. T., Continuous Source Monitonng of Gaseous
Sulfur Compounds in the Paper Industry. (Presented at the 12th Conference on
Methods in Air Pollution and Industrial 1-lygiene Studies, Los Angeles, California, April
7, 1971.)
10. Miller, A. M., Brown, J., and Abrama, R., Applied Techniques of Analyses for Stack
Emissions. (Presented at the West Coast Regional Meeting of the National Council for
Air and Stream Improvement, Portland, Oregon, October 2, 1968.)
11. Mathis, G. V., Application of an Electrochemical Cell to NO and SO 2 Monitoring.
(Presented at the Miami University Symposium on Instrumentation for Continuous
Monitoring of Air and Water Quality, Oxford, Ohio, June 21, 1973.)
17-41
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12. Thoen, G. N., DcHaas, C. G., and Baumgartel, F. A., Continuous Sulfur Dioxide
Monitor and its Application to Sulfite Recovery Emissions. Tappi, 52:2304-2305,
December 1969.
13. Saltzman, R. S., Use of Photometric Analyses for Ultraviolet Analyzers for NO and
SON. (Presented at the Miami University Symposium on Instrumentation for
Continuous Monitoring of Air and Water Quality, Oxford, Ohio, June 21, 1973.)
14. Personal communication with Mr. Philip C. Stultz, Boise.Cascade Corporation, Salcm,
Oregon, .June 21, 1973.
15. A Guide to the Use of Gas Chromatography in Emission Analysts. Atmospheric Quality
Improvement Technical Bulletin No. 59. National Council for Air and Stream
Improvement, New York, New York, February 15, 1973.
16. Adams, D. F., and Koppc, R. K., Evaluation of Gas Chroinalographic Columns for
Analyses of Subparts per Million Concentrations of Gaseous Sulfur Compound.
Environmental Science and Technology, 1:479-483, June 1967.
17. Mulik, J. D., Stevens, R. K., and Baumgardner, R., An Analytical System Designed to
Measure Multiple Malodorous Compounds Related to Kraft Mill Activities. (Presented
at the TAPPI Water & Air Conference, Boston, Massachusetts, April 4, 1971.)
18. Rossano, A. T., and Cooper, H. B. H., Procedure for Calibrating a Continuous NO 2
Analyzer. Journal of the Air Pollution Control Association, 13:518-523, November
1963.
19. O’Keefc, A. E., Ortrnan, C. C., Primary Standards for Trace Gas Analysis. Analytical
Chemistry, 38:760-763, June 1966.
20. Duncan, L., aiid Tucker, T. W., A Guide to the Use of Permeation Tubes as Primary
Standards for Instrument Catibrat ion. Atmospheric Pollution Technical Bulletin
No. 47, National Council of the Paper Industry for Air and Stream Improvement, New
York, New York, May 1970.
21. Duckworth, S., Lcvaggi, D., and Urn, J., Field Dynamic Calibration of SO 2 Recording
Instruments. Journal of the Air Pollution Control Association, 13:429-434, September
1963.
22. Cooper, H. B. H., The Particulate Problem: Continuous Particulate Monitoring in the
Pulp and Paper Industry. (Presented at the iMiami University Symposium on
Continuous Monitoring of Air and Water Quality, Oxford, Ohio, June 20, 1973.)
17-42
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23. Determining Dust Concentrations in a Gas Stream. Performance Test Code 27-1957.
American Society of Mechanical Engineers, New York, New York, 1957.
24. Devorkin, H., Chass, R. L., and Fudvrich, A. P., Source Testing Manual. Air Pollution
Control District, Los Angeles, California, December 1972.
25. Code of Federal Regulations, Part 60, Chapter 1, Title 40. Standards of Performance
for New Stationary Sources. Method 5. December 23, 1971.
26. Bosch, J. C., Pilat, M. J., and Hrutfiord, B. F., Size Distribution of Aerosols from a
Kraft Mill Recovery Furnace. Tappi, 54:1871.1875, November 1971.
27. Walker, C. G., Manual for Counting and Sizing Particles from Kraft Recovery Furnaces.
Atmospheric Pollution Technical Bulletin No. 19. National Council of the Paper
industry for Air and Stream Improvement, New York, New York, July 1963.
28. Sem, G. J., et a!., State of the Art: 1971 Instrumentation for Measurement of
Particulate Emissions from Combustion Sources— Volumes I & II: Particulate Mass.
Reports APTD 0733 and 0734, Documents PB 202 665 and PB 202 666. U.s.
Environmental Protection Agency, Air Pollution Control Office, Durham, North
Carolina, April 1971.
29. Leonard, J. S., Continuous Kraft Mill Emission Monitoring. In: Blosser, R. 0., and
Cooper, H. B. H. (eds.)., Analytical Equipment and Monitoring Devices for Gases and
Particulates. Atmospheric Pollution Technical Bulletin No. 35, National Council for
the Paper Industry for Air and Stream improvement, New York, New York, March
1968.
30. Tretter, V. J., Use of Continuous Monitors of Soda Loss and Malodorous Sulfur Loss in
Process Control. Tappi, 52:2324-2326, December 1969.
31. Larssen, S., Ensor, D. S., and Pilat, M. J., Relationship of Plume Opacity to the
Properties of Particulate Emitted from Kraft Recovery Furnaces. Tappi, 55:88.92,
January 1972.
32. Gansler, N. R., The Use of a Bolometer for Continuous Measurement of Particulate
Losses from Kraft Recovery Furnaces. (Presented at the Annual Meeting of the Pacific
Northwest International Section of the Air Pollution Control Association, Vancouver,
British Columbia, November 22, 1968.)
33. Beutner, H. P., Measurement of Opacity and Particulate Emissions from Stocks.
(Presented at the Symposium on Instrumentation for Continuous Monitoring of Air
Water Quality, Miami University, Oxford, Ohio, June 20, 1973.’)
17-43
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34. Cooper, H. B. H., Ambient and Source Odor Measurements. (Presentcd at the Third
Annual Industrial Air Pollution Control Conference, University of Tennessee,
Knoxville, Tennessee, March 29, [ 973.)
35. Cooper, H. B. H., and Rossano, A. T., Particulate Matter and Odor Control.
Unpublished Special Report, University of Washington, Department of Civil
Engineering, Seattle, Washington, January 1971.
36. Physiological Effects. In: Air Pollution Abatement Manual. Washington, D.C.
Manufacturing Chemists Association, 1951.
37. Byrd, J. F., Phelps, A. A., Odor and Its Measurement. In: Stern, A. C. (cd.). Air
Pollution, Volume II, 2nd Edition. New York. Academic Press, Inc., 1968.
38. Prince, R. C. H., and Ince, J. I-I., The Measurement of Intensity of Odors. Journal of
Applied Chemistry, 8:314-32, May 1958.
39. Lindvall, T., On Sensory Evaluation of Odorous Air Pollutant Intensities. Atmospheric
Pollution Technical Bulletin No. 50, National Council of the Paper Industry for Air
and Stream Improvement, New York, New York, September 1970.
40. Caron, A. L., and Adams, D. F., Evaluation of the Use of Humans in Measuring the
Effectiveness of Odor Control Technology at the Source. Atmospheric Pollution
Technical Bulletin No. 36, National Council of the Paper Industry for Air and Stream
Improvement, New York, New York, September 1971.
41. Standard Method for Measurement of Odors in Atmospheres (Dilation Method). ASTM
Standard D 139-57. Philadelphia: American Society for Testing and Materials, 1957.
p. 185.188.
42. Cederlof, R., Edfors, M. L., Friberg, L., and Lindvall, T., Determination of Odor
Thresholds for Flue Gases from a Swedish Sulfate Cellulose Plant. Tappi, 48:405-411,
July 1965.
43. Huey, N. A., Broering, L. C., Jutze, G. A., and Guiber, C. W., Objective Odor Pollution
Control Investigations. Journal of the Air Pollution Control Association, 10:441.446,
December 1960.
44. Walther, J. E., and Amberg, H. R., Expenence with a Mobile Laboratory in Source
Sampling Kraft Mill Emissions. Tappi, 51:126A.129A, November 1968.
45. Mulik, J. D., Stevens, R. K., and Baumgardner, R. A., An Analytical System Designed
to Measure Multiple Malodorous Compounds Related to Kraft Mill Activities. In:
17.44
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Proceedings of the 12th Conference on Methods in Air Pollution and Industrial
Hygiene Studies. University of Southern California, Los Angeles, California, April 6-8,
1971.
46. Megy, J. A., Design, Operation, and Use of a Mobile Laboratory for Continuous
Monitoring of Kraft Mill Source Gases. (Presented at the West Coast Regional Meeting
of the National Council of the Paper Industry for Air and Stream Improvement,
Seattle, Washington, October 15, 1969.)
47. Waddington, G. E., A II’Iobtle Ambient and Stack Sampling System for the Pulp and
Paper Industry—A Unified Systems Approach. Pulp and Paper Magazine of Canada,
74:58-61, June 1973.
17-45
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APPENDIX A
GLOSSARY OF SYMBOLS
Symbol Definition
B&W Babcock & Wilcox
BOD biochemical oxygen demand
BOD, biochemical oxygen demand as determined by 7-day test
BLO black liquor oxidation
BLRBAC Black Liquor Recovery Boiler Advisory Committee
BP boiling point
BHV bomb heat values
BD bone dry
BCRC British Columbia Research Council
BTU British thermal unit
cm centimeter
CE Combustion Engineering
ACE CE system of contact evaporation using combustion air
LAH CE system of evaporation using Laminaire air heaters
cfm cubic feet per minute
cfs cubic feet per second
cu ft, ft 3 cubic foot
cubic meter
m 3 /h cubic meters per hour
°C degree Celsius
°F degree Fahrenheit
°K degree Kelvin
DO dissolved oxygen
DS dry solids
ft foot
fpm feet per minute
fps feet per second
oz fluid ounce
gal gallon
gpm gallons per minute
gr gram
g gram
h,hr hour
hp horsepower
in inch
A-i
-------
Symbol Definition
ID induced draft
SI International System of Units (“metric system”)
kcal kilocalorie
kg kilogram
kJ kilojoule
km kilometer
kW kilowatt
LK lime kiln
MJ megajoules
MPa megapascal ( 106 kg• rn/sec 2 per m 2 )
t metric ton
micrometer
mm millimeter
mg milligram
mV millivolt
mm minute
ng nanogram
nm nanometer
NCASI National Council for Air and Stream Improvement
NSSC neutral sulfite semi-chemical (process)
NCG noncondensable gases
R organic chemical radical
NO oxides of nitrogen
oxides of sulfur
ppb parts per billion
ppm parts per million
lb pound
pcf pounds per cubic foot
psia pounds per square inch, absolute pressure
psig pounds per square inch, gauge pressure
psi pounds per square inch
RF recovery furnace
s, sec second
ton short ton (2000 ib)
SSL spent sulfite liquor
std standard
scf standard cubic foot
scm standard cubic meter
sdcf standard dry cubic foot
SBLO strong black liquor oxidation
A-2
-------
Symbol Definition
TRS total reduced sulfur
w.g. water gauge
W watt
WBLO weak black liquor oxidation
yd yard
yr year
A-3
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APPENDIX B
CHEMICAL FORMULAS
Acetic Acid CH 3 COOH
Acetone CH 3 COCH 3
Acrolein CH 2 CHCHO
Ammonia NH 3
Ammonium bisulfite NH 4 HSO 3
Ammonium sulfate (NH 4 ) 2 SO 4
Ammonium sulfite (NH 4 ) 2 SO 3
Arsenic As
Beryllium Be
Boric acid H 3 B0 3
Bromine, molecular bromine Br, Br 2
Cadmium Cd
Cadmium sulfate CdSO 4
Calcium bisulfite Ca(HSO 3 ) 2
Calcium Oxide (lime) CaO
Calcium sulfite CASO 3
Carbon dioxide CO 2
Carbon monoxide CO
Carbonic Acid H 2 CO 3
Carbonyl hydrogen sulfide COSH
Carbonyl sulfide COS
Chlorine dioxide C10 2
Chlorine, molecular chlorine Cl, Cl 2
Citric Acid C 6 H 8 07
Dimethyl disulfide CH 3 SSCH 3
Dimethyl Sulfide CH 3 SCH 3
Ethanol, Ethyl alcohol CH 3 CH 2 OH
Ferric chloride Fe 2 Cl 6
Ferrous sulfide FeS
Fluorine Fl
Formaldehyde HCHO
Hydrobromic acid HBr
Hydtochloric acid HCI
Hydrogen, molecular hydrogen H, H 2
Hydrogen peroxide H 2 0 2
Hydrogen sulfide H 2 5
Lead. Pb
B-i
-------
Magnesium bisulfite Mg(HSO 3 ) 2
Magncsium hydroxide Mg(OH) 2
Magnesium oxide, magnesia MgO
Mercury Hg
Methane CH 4
Methanol, methyl alcohol CH 3 OH
Methyl mercaptan, methanethiol CH 3 SH
Nitric oxide NO
Nitrogen dioxide NO 2
Nitrogen, molecular nitrogen N, N 2
Nitrogen oxides NO
Oxygen, molecular oxygen, ozone 0, 02, 03
Phosphorus P
Potassium acid phthalate KHC 8 H 4 04
Selenium Se
Silver nitrate AgNO 3
Sodium bicarbonate NaHCO 3
Sodium hydrogen sulfide NaHS
Sodium bisulfite NaHSO 3
Sodium carbonate Na 2 CO 3
Sodium chloride NaC1
Sodium chlorite NaC IO 2
Sodium hydroxide NaOH
Sodium mercaptide CH 3 CH 2 SNa
Sodium sulfate Na 2 SO 4
Sodium sulfite Na 2 SO 3
Sodium thiosulfate Na 2 52 03
Sulfur S
Sulfur dioxide SO 2
Sulfur trioxide SO 3
Sulfuric acid H 2 SO 4
Vanadium V
Vanadium pentoxide V 2 O 5
Water H 2 0
Zinc Zn
B.2
* US GOVERNMENT PRINTING OFFICE 197 1-660-859
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