Biofilm Reactors
Prepared by the Biofilm Reactors Task Force of the Water Environment Federation® Rhodes R. Copithorn, P.E., BCEE, Chair Joshua P. Boltz, Ph.D., P.E. Christine deBarbadillo, P.E. Paul Dombrowski, P.E., BCEE John R. Harrison, P.E. Sarah Hubbell Joseph A. Husband, P.E., BCEE Samuel Jeyanayagam, Ph.D., P.E., BCEE Ishin Kaya, P.Eng. Carl M. Koch, Ph.D., P.E., BCEE David J. Krichten Frank M. Kulick III
James P. McQuarrie, P.E. Robert Nerenberg, Ph.D., P.E. Heather M. Phillips, P.E. Dana W. Rippon, P.E. Frank Rogalla Edward D. Schroeder, Ph.D. Dipankar Sen, Ph.D., P.E. Spencer Snowling George Sprouse, Ph.D., P.E. Christopher W. Tabor, P.E. Stephen Tarallo Ifetayo Venner, P.E. Thomas E. Wilson, Ph.D., P.E., BCEE Stefan Wuertz, Ph.D.
Under the Direction of the Municipal Design Subcommittee of the Technical Practice Committee 2010 Water Environment Federation® 601 Wythe Street Alexandria, VA 22314–1994 USA http://www.wef.org
Biofilm Reactors
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WEF Manual of Practice No. 35 Prepared by the Biofilm Reactors Task Force of the Water Environment Federation®
., BCEE
WEF Press Water Environment Federation Alexandria, Virginia
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About WEF Formed in 1928, the Water Environment Federation® (WEF ®) is a not-for-profit technical and educational organization with members from varied disciplines who work toward WEF’s vision to preserve and enhance the global water environment. For information on membership, publications, and conferences, contact: Water Environment Federation® 601 Wythe Street Alexandria, VA 22314-1994 USA (703) 684-2400 http://www.wef.org
Manuals of Practice of the Water Environment Federation® The WEF Technical Practice Committee (formerly the Committee on Sewage and Industrial Wastes Practice of the Federation of Sewage and Industrial Wastes Associations) was created by the Federation Board of Control on October 11, 1941. The primary function of the Committee is to originate and produce, through appropriate subcommittees, special publications dealing with technical aspects of the broad interests of the Federation. These publications are intended to provide background information through a review of technical practices and detailed procedures that research and experience have shown to be functional and practical. Water Environment Federation Technical Practice Committee Control Group R. Fernandez, Chair J. A. Brown, Vice-Chair B. G. Jones, Past Chair A. Babatola L. W. Casson K. Conway V. D’Amato R. P. Dominak A. Ekster R. C. Johnson S. Moisio T. Page-Bottorff S. J. Passaro R. C. Porter E. P. Rothstein A. T. Sandy A. Tyagi A. K. Umble
Contents List of Figures . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . xxv List of Tables . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . xxxiii Preface . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . xxxvii
Chapter 1 Introduction 1.0 Background and Purpose . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1 2.0 Characteristics of Fixed-Growth Processes . . . . . . . . . . . . . 2 3.0 History . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 3 3.1 Contact Beds. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4 3.2 Trickling Filters. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4 3.3 Rotating Biological Contactors. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5 3.4 Coupled Trickling Filter/Activated Sludge Process. . . . . . . . . . . . . . . . 6 3.5 Biological Filters. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7 3.6 Hybrid Processes . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7 4.0 Organization of Manual . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 10 5.0 References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 11 6.0 SUGGESTED READINGS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 14
Chapter 2 Biology of Fixed-Growth Process 1.0 INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 17 2.0 CLASSIFICATION OF LIVING ORGANISMS . . . . . . . . . . . . . . . . . . . . . . 18 3.0 MICROORGANISMS OTHER THAN BACTERIA . . . . . . . . . . . . . . . . . . 21 3.1 Fungi. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 22 3.2 Algae. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 22 3.3 Protozoa. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 23 3.4 Multicellular Invertebrates. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 23 vii
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3.5 Viruses. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 24 3.6 Consortia. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 26 4.0 CHARACTERISTICS OF BACTERIA . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 26 4.1 Structure of the Bacterial Cell. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 26 4.2 Chromosome and Plasmids . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 28 4.3 Cytoplasm. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 30 4.4 Cell Membrane. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 30 4.5 Cell Wall. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 30 4.6 Pili. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 31 4.7 Flagella. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 31 4.8 Extracellular Polymeric Substances. . . . . . . . . . . . . . . . . . . . . . . . . . . . 31 4.9 Chemical Composition of Cells . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 32 4.10 Example 1—Theoretical Oxygen Demand of Bacterial Cells. . . . . . . 32 4.10.1 Solution. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 32 4.10.2 Comment. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 33
5.0 B ACTERIAL METABOLISM, NUTRITION, AND RESPIRATION . . . . 33 5.1 Energy Source. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 36 5.2 Chemoheterotrophic Metabolism . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 36 5.3 Chemoautotrophic Metabolism. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 38 5.4 Photosynthetic Metabolism. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 39 5.5 Nutrient Requirements . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 40 5.6 Bacterial Energy Metabolism . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 41 5.7 Aerobic Growth and Respiration. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 42 5.8 Anoxic Conditions and Respiration. . . . . . . . . . . . . . . . . . . . . . . . . . . . 43 5.9 Anaerobic Respiration and Fermentative Metabolism. . . . . . . . . . . . 44 5.10 Energetics of Respiration. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 46 5.11 Example 2—Calculation of Electrode Reduction Potentials and ∆G for Half-Reactions ��������������������������������������������������������������������48 5.12 Solution . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 48 5.13 Co-Metabolism. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 50
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6.0 BACTERIAL GROWTH . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 51 6.1 The Bacterial Growth Curve. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 51 6.2 Growth in Mixed Cultures . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 53 6.3 Enrichment Cultures . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 53 6.4 Stability of Mixed Cultures. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 54 6.5 Effects of Environmental Variables. . . . . . . . . . . . . . . . . . . . . . . . . . . . 54 7.0 BACTERIAL GROWTH KINETICS IN BIOFILMS . . . . . . . . . . . . . . . . . 55 7.1 Rate of Bacterial Processes . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 56 7.2 Note . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 56 7.3 Physical and Chemical Changes in Biofilms Resulting from Growth��������������������������������������������������������������������������������������������59 7.4 Structured Models . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 59 7.5 Temperature Effects. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 60 7.6 Example 3—Effect of Temperature on Organic Removal . . . . . . . . . 61 7.7 Solution . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 61 7.8 Inhibition and Toxicity . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 63 7.9 Mass-Transfer-Rate Limitations. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 64 8.0 KEY TRANSFORMATIONS IN BIOFILMS . . . . . . . . . . . . . . . . . . . . . . . 65 8.1 Chemoheterotrophic Processes . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 66 8.2 Chemoautotrophic Processes. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 67 8.3 Biology of Nitrogen Transformations. . . . . . . . . . . . . . . . . . . . . . . . . . 67 8.4 Denitrification. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 67 8.5 Aerobic Nitrification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 68 8.6 Anoxic Nitrification/Denitrification. . . . . . . . . . . . . . . . . . . . . . . . . . . 69 8.7 Biological Phosphorus Removal. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 70 8.8 Sulfide and Sulfur Oxidation. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 71 8.9 Hydrogen Oxidation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 71 9.0 FEATURES OF MICROBIAL COMMUNITIES IN BIOFILMS . . . . . . . 72 10.0 REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 74
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Chapter 3 Trickling Filter and Combined Trickling Filter Suspended-Growth Process Design and Operation 1.0 INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 83 2.0 General Description . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 84 2.1 Distribution System. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 84 2.2 Biofilm Carriers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 87 2.3 Containment Structure . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 90 2.4 Underdrain System and Ventilation. . . . . . . . . . . . . . . . . . . . . . . . . . . . 90 2.5 Trickling Filter Pumping Stations: Influent and Recirculation. . . . . . 91 2.6 Hydraulic and Contaminant Loading. . . . . . . . . . . . . . . . . . . . . . . . . . . 92 3.0 Process Flow Sheets and Bioreactor Configuration . . 93 3.1 Standard Process Flow Diagrams . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 93 3.2 Bioreactor Classification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 96 3.3 Hydraulic Application: Effect on Media Wetting, Flow Distribution, and Control ��������������������������������������������������������������98 4.0 Ventilation and Air Supply Alternatives . . . . . . . . . . . . . . 100 4.1 Natural Draft. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 101 4.2 Mechanical Ventilation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 102 5.0 Trickling Filter Process Models . . . . . . . . . . . . . . . . . . . . . . . . 103 5.1 National Research Council. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 103 5.2 Galler and Gotaas. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 105 5.3 Kincannon and Stover. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 106 5.4 Velz. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 107 5.5 Schulze. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 107 5.6 Germain. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 108 5.7 Eckenfelder . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 109 5.8 Chartered Institution of Water and Environmental Management ����������������������������������������������������������������110 5.9 Logan Trickling Filter Model. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 111
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5.10 Selecting a Trickling Filter Model . . . . . . . . . . . . . . . . . . . . . . . . . . . . 112 5.11 Method for Combining Trickling Filter and Suspended-Growth Models . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 113 6.0 Process Design . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 115 6.1 Combined Carbon Oxidation and Nitrification. . . . . . . . . . . . . . . . . . 115 6.2 Nitrifying Trickling Filters . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 120 6.2.1 Gujer and Boller Nitrifying Trickling Filter Model. . . . . . . . . . . . . . . . . . . . . 122 6.2.2 Okey and Albertson Nitrifying Trickling Filter Model. . . . . . . . . . . . . . . . . . 124 6.2.2.1 Application of the Gujer and Boller Model. . . . . . . . . . . . . . . . . . . . . . 126 6.2.2.2 Application of the Albertson and Okey Model . . . . . . . . . . . . . . . . . . . 128
6.3 Temperature and Hydraulic Application Effects . . . . . . . . . . . . . . . . 131 7.0 Design Considerations . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 132 7.1 Distribution System. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 133 7.1.1 Hydraulic Drive Rotary Distributors.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 134 7.1.2 Electronic or Mechanical Drive Rotary Distributors. . . . . . . . . . . . . . . . . . . . 136 7.1.3 Optimizing Rotary Distributor Operation. . . . . . . . . . . . . . . . . . . . . . . . . . . . 138
7.2 Construction of Rotary Distributors . . . . . . . . . . . . . . . . . . . . . . . . . . . 138 7.3 Trickling Filter Media Selection. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 139 7.3.1 Depth. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 141 7.3.2 Structural Integrity . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 142
7.4 Trickling Filter Pumping Station or Dosing Siphon. . . . . . . . . . . . . . 144 7.5 Control Mechanisms for Trickling Filter Macro Fauna . . . . . . . . . . . 144 7.5.1 Operational Strategies and Facility Improvements for Macro Fauna Control . . . 145 7.5.2 Spülkraft. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 147 7.5.3 Flooding.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 148 7.5.4 Chemical Treatment.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 149 7.5.5 Physical Control����������������������������������������������������������������������������������������������152
7.6 Trickling Filter Startup . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 153 7.7 Combined Trickling Filter and Suspended-Growth Processes. . . . . 155 7.7.1 Activated Biofilter.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 156 7.7.2 Trickling Filter/Solids Contact. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 158
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7.7.3 Roughing Filter/Activated Sludge.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 162 7.7.4 Biofilter/Activated Sludge.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 162 7.7.5 Trickling Filter/Activated Sludge. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 162
8.0 REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 163
Chapter 4 Rotating Biological Contactors 1.0 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 174 2.0 Process design considerations . . . . . . . . . . . . . . . . . . . . . . . . . 178 2.1 Media Surface Area . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 179 2.2 pH and Nutrient Balance. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 180 2.3 Oxygen Transfer. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 180 2.4 Flow and Loading Variability. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 182 2.5 Operating Temperature. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 183 2.6 Solids Production. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 183 2.7 Toxic and Inhibitory Substances . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 184 3.0 Rotating biological contactor design methods . . . . . 184 3.1 Monod Kinetic Model . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 184 3.2 Second-Order Model. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 186 3.3 Empirical Model. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 187 3.4 Manufacturers’ Design Curves. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 189 3.5 Comparison of Model Predictions. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 190 3.6 Predicted Performance versus Full-Scale Data . . . . . . . . . . . . . . . . . . 191 3.7 Temperature Correction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 192 4.0 Rotating biological contactor nitrification models . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 194 5.0 Denitrification Application . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 197 6.0 Physical design features . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 198 6.1 Physical Layout. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 198 6.2 Tank Volume. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 198 6.3 Hydraulics and Flow Control. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 198 6.4 Media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 199
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6.5 Drive Systems. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 200 6.6 Covers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 201 6.7 Biomass Control. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 201 7.0 Rotating biological contactor design examples . . . 202 7.1 Secondary Treatment Design Example. . . . . . . . . . . . . . . . . . . . . . . . 202 7.2 Advanced Secondary Treatment Design Example. . . . . . . . . . . . . . 203 8.0 Problems and corrective actions . . . . . . . . . . . . . . . . . . . . . 204 8.1 Inadequate Treatment Capacity. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 205 8.2 Excessive First-Stage Loadings. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 205 8.3 Excessive Biomass Growth. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 206 8.4 Loping of Air-Drive Systems . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 206 8.5 High Clarifier Effluent Suspended Solids. . . . . . . . . . . . . . . . . . . . . . 207 8.6 Corrosion of Media Supports. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 207 9.0 Pilot-plant studies . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 207 10.0 References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 208
Chapter 5 Moving-Bed Biofilm Reactors 1.0 INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 212 2.0 Moving-Bed Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 213 3.0 Design Considerations for Moving-Bed Reactors . . . . . 217 3.1 Carrier Biofilms. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 218 3.1.1 Carbonaceous Matter Removal. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 220 3.1.2 High-Rate Designs. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 220 3.1.3 Normal-Rate Designs. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 221 3.1.4 Low-Rate Designs. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 221 3.1.5 Nitrification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 224 3.1.6 Denitrification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 229 3.1.6.1 Pre-Denitrification Moving-Bed Biofilm Reactors.. . . . . . . . . . . . . . . . 229 3.1.6.2 Post-Denitrification Moving-Bed Biofilm Reactors . . . . . . . . . . . . . . . 230 3.1.6.3 Combined Pre-/Post-Denitrification Moving-Bed Biofilm R eactors������������������������������������������������������������������������������������230
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3.2 Mixers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 231 3.3 Pretreatment. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 232 4.0 Solids Separation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 233 5.0 General considerations for moving-bed biofilm reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 234 5.1 Approach Velocity. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 234 5.2 Foaming. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 234 5.3 Media Transfer and Inventory Management. . . . . . . . . . . . . . . . . . . . 234 6.0 CASE STUDIES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 235 6.1 Moa Point Wastewater Treatment Plant, Wellington, New Zealand��������������������������������������������������������������������������������������������235 6.2 Harrisburg Wastewater Treatment Plant, Harrisburg, Pennsylvania��������������������������������������������������������������������������������������������238 6.3 Moorhead Wastewater Treatment Facility, Moorhead, Minnesota (Adapted from Zimmerman et al., 2004)��������������������������240 6.4 Williams Monaco Wastewater Treatment Plant, Henderson, Colorado ��������������������������������������������������������������������������������������������������241 6.5 Klagsham Wastewater Treatment Plant, Malmö, Sweden (Adapted from Taljemark et al., 2004)��������������������������������������������������246 6.6 Gardemoen Wastewater Treatment Plant, Gardemoen, Norway. . . . . 250 7.0 References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 253
Chapter 6 Hybrid Processes 1.0 OVERVIEW OF INTEGRATED FIXED-FILM ACTIVATED SLUDGE SYSTEMS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 260 1.1 Advantages. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 261 1.2 Disadvantages. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 262 2.0 Media Types . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 262 2.1 Fixed-Media Systems. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 263 2.2 Free-Floating Media Systems. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 264 2.2.1 Plastic. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 264 2.2.2 Sponge. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 264
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3.0 History of Process . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 265 4.0 Application of INTEGRATED FIXED-FILM ACTIVATED SLUDGE Systems . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 267 4.1 Fixed Media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 267 4.1.1 General Requirements. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 267 4.1.2 Growth on Media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 268 4.1.3 Kinetics. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 269 4.1.4 Worm Growth. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 270 4.1.5 Media Breakage. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 270 4.1.6 Dissolved Oxygen Level. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 270 4.1.7 Mixing . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 271 4.1.8 Access to Diffusers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 271 4.1.9 Odor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 271
4.2 Free-Floating Media—Sponge Media. . . . . . . . . . . . . . . . . . . . . . . . . . 272 4.2.1 General Requirements. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 272 4.2.2 Screen Clogging. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 272
4.3 Control of Biomass Growth . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 273 4.3.1 Loss of Sponges. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 273 4.3.2 Taking Tank Out-of-Service. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 274 4.3.3 Loss of Solids. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 274 4.3.4 Air Distribution System. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 274 4.3.5 Plastic Media . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 274 4.3.5.1 General Requirements.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 274 4.3.5.2 Biomass Growth.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 275 4.3.5.3 Media Mixing.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 276 4.3.5.4 Screens. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 276 4.3.5.5 Foaming. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 276 4.3.5.6 Media Replacement. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 276 4.3.5.7 Taking Tank Out-of-Service.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 276 4.3.5.8 Worm Growth.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 277 4.3.5.9 Startup.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 277
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5.0 Process Design . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 277 5.1 Introduction. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 277 5.2 Parameters Influencing Organics Removal in the Biofilm of Integrated Fixed-Film Activated Sludge Systems������������������������������278 5.2.1 Biofilm Flux Rates . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 278 5.2.2 Removals in Biofilm per Unit of Tank Volume . . . . . . . . . . . . . . . . . . . . . . . . 278
5.3 Parameters Influencing Removals in the Mixed-Liquor Suspended Solids������������������������������������������������������������������������������������281 5.4 Interaction Between the Mixed-Liquor Suspended Solids and the Biofilm����������������������������������������������������������������������������������������282 5.5 Interaction Between Heterotrophs and Nitrifiers. . . . . . . . . . . . . . . . 284 5.6 Design Tools/Procedures. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 284 5.6.1 Empirical Methods. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 285 5.6.1.1 Equivalent-Sludge-Age Approach. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 285 5.6.1.2 Quantity (Length or Web Surface Area) of Media Approach. . . . . . . . 286 5.6.2 Rates Based on Pilot Studies. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 286 5.6.3 Biofilm Rate Model. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 287 5.6.3.1 Define Range of Flux Rates. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 287 5.6.3.2 Quantify Removal at Different Mixed-Liquor Suspended Solids Mean Cell Residence Times ������������������������������������������������������287 5.6.3.3 Select Flux Rates Based on Location Along Aerobic Zone. . . . . . . . . . 287 5.6.3.4 Calculate the Quantity of Media Required. . . . . . . . . . . . . . . . . . . . . . 288 5.6.3.5 Additional Analysis to Finalize a Design. . . . . . . . . . . . . . . . . . . . . . . 288 5.6.3.6 Application of Kinetics-Based Approach with Integrated Fixed-Film Activated Sludge Design Software������������������������������������288
6.0 Case Studies . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 288 6.1 Annapolis Water Reclamation Facility, Anne Arundel County, Maryland��������������������������������������������������������������������������������������������������288 6.1.1 Original Wastewater Treatment Plant. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 289 6.1.2 Pilot Study (1993 to 1996). . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 289 6.1.3 Full-Scale Upgrade for Biological Nutrient Removal (1997 to 2000). . . . . . . 291 6.1.3.1 Pilot Study. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 291
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6.1.3.2 During Construction (1997 to 2000) .. . . . . . . . . . . . . . . . . . . . . . . . 293 6.1.3.3 Post-Construction (2000 to 2003). . . . . . . . . . . . . . . . . . . . . . . . . . . . . 299
6.2 Westerly Wastewater Treatment Plant, Westerly, Rhode Island . . . . . 299 6.2.1 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 299 6.2.2 Description of Original Facilities. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 299 6.2.3 Description of Upgrade. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 299 6.2.4 Design Criteria. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 302 6.2.5 Performance of the Integrated Fixed-Film Activated Sludge System. . . . . . . 302 6.2.6 Operational Issues . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 306 6.2.7 Costs. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 307
6.3 Broomfield Wastewater Treatment Plant, Broomfield, Colorado . . . 307 6.3.1 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 307 6.3.2 Full-Scale Plant Results. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 308
6.4 Colony Wastewater Treatment Plant, Colony, Texas. . . . . . . . . . . . . 308 6.4.1 Introduction and Background. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 308 6.4.2 Changing Design Conditions. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 316 6.4.3 Plant Construction and Operation. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 318 6.4.4 System Flexibility. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 320 6.4.5 Redworm Predation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 321
7.0 References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 321
Chapter 7 Biological Filters 1.0 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 327 2.0 Descriptions of BIOLOGICALLY ACTIVE FILTER Reactors and Equipment . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 329 2.1 Brief History of Biologically Active Filters. . . . . . . . . . . . . . . . . . . . . . 329 2.2 Downflow Biologically Active Filter with Sunken Media. . . . . . . . . 331 2.3 Upflow Biologically Active Filter with Sunken Media. . . . . . . . . . . . 334 2.4 Upflow Biologically Active Filter with Floating Media. . . . . . . . . . . 335 2.5 Moving-Bed, Continuous Backwash Filters. . . . . . . . . . . . . . . . . . . . . 337 2.6 Non-Backwashing, Open-Structure Media Filters . . . . . . . . . . . . . . . 339
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3.0 Media for Use in BIOLOGICALLY ACTIVE FILTERS . . . . . . . . . . . 341 3.1 Mineral Media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 341 3.2 Random Plastic Media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 343 3.3 Modular Plastic Media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 345 4.0 Backwashing and Air-Scouring . . . . . . . . . . . . . . . . . . . . . . . . . 345 5.0 BIOLOGICALLY ACTIVE FILTER Process Design . . . . . . . . . . . . . 349 5.1 Process Design for Secondary Treatment. . . . . . . . . . . . . . . . . . . . . . . 351 5.1.1 Volumetric Biochemical Oxygen Demand Loading. . . . . . . . . . . . . . . . . . . . . 351 5.1.2 Hydraulic Loading. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 351 5.1.3 Backwashing. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 351 5.1.4 Design Example: Design of a Submerged, Upflow Biological Aerated Filter System for Secondary Treatment (No Nitrification) ��������������353 5.1.5 Solution. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 353
5.2 Process Design for Nitrification. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 355 5.2.1 Influence of Hydraulic Filtration Rates. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 355 5.2.2 Effect of Process Air Velocity. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 357 5.2.3 Dependence on Loading Conditions. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 359 5.2.4 Temperature Effects . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 360 5.2.5 Design Example: Design of a Submerged, Upflow Biological Aerated Filter System for Nitrification Following Secondary Treatment��������������������360 5.2.6 Solution. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 361
5.3 Process Design for Combined Nitrification and Denitrification. . . . 362 5.4 Process Design for Tertiary Denitrification . . . . . . . . . . . . . . . . . . . . . 365 5.4.1 Volumetric Mass Loading . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 365 5.4.2 Half-Order Kinetic Model . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 367 5.4.3 Hydraulic Loading. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 368 5.4.4 Solids Removal and Sludge Production. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 368 5.4.5 Supplemental Carbon Requirements . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 369 5.4.6 Tertiary Denitrification Typical Operations Issues and Corrective Actions�������������������������������������������������������������������������������������������� 370 5.4.6.1 Excess Backwashing. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 370 5.4.6.2 Gas (Nitrogen) Accumulation. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 371
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5.4.6.3 Solids Breakthrough. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 371 5.4.6.4 Nitrate/Nitrite Breakthrough ��������������������������������������������������������������371 5.4.6.5 Carbon Breakthrough. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 372 5.4.6.6 Phosphorus Management.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 372 5.4.6.7 Operation During Peak Flow Events . . . . . . . . . . . . . . . . . . . . . . . . . . 372
5.5 Phosphorus Removal Considerations for Biologically Active Filter Processes����������������������������������������������������������������������������������������373 6.0 Design Considerations . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 374 6.1 Preliminary and Primary Treatment. . . . . . . . . . . . . . . . . . . . . . . . . . . 374 6.2 Backwash Handling Facilities. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 374 6.3 Biologically Active Filter Process Aeration . . . . . . . . . . . . . . . . . . . . . 375 6.3.1 Oxygen-Transfer Efficiency. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 375 6.3.2 Process Air Distribution Systems . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 377 6.3.3 Process Air Control . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 377
6.4 Supplemental Carbon Feed Requirements. . . . . . . . . . . . . . . . . . . . . . 378 7.0 Biologically active filter Case Studies . . . . . . . . . . . . . . . . 379 7.1 Chemically Enhanced Primary Treatment Followed by Two-Stage Biologically Active Filter for Total Nitrogen Removal: VEAS Wastewater Treatment Plant, Oslo, Norway ��������379 7.2 Chemically Enhanced Primary Treatment Followed by Three-Stage Biologically Active Filter for Total Nitrogen Removal: Siene Centre Wastewater Treatment Plant, Colombes, France��������������������������������381 7.3 Total Nitrogen Removal in a Single-Stage Biologically Active Filter: Frederikshavn Wastewater Treatment Plant, Denmark��������384 7.4 Nitrification and Denitrification: West Warwick, Rhode Island��������������������������������������������������������������������������������������������387 7.5 Post-Denitrification Sand Filters: Havelock, North Carolina. . . . . . 389 8.0 REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 391
Chapter 8 New and Emerging Fixed-Film Technologies 1.0 INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 401 2.0 BIOFILM REACTORS WITH SUSPENDED CARRIERS OR GRANULES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 402
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2.1 Biofilm Airlift Suspension Reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . 402 2.2 Upflow Anaerobic Sludge Blanket . . . . . . . . . . . . . . . . . . . . . . . . . . . 404 2.3 Expanded Granular Sludge Blanket . . . . . . . . . . . . . . . . . . . . . . . . . . 404 2.4 Internal Circulation Reactor. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 404 3.0 ANAMMOX BIOFILM REACTORS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 405 4.0 MEMBRANE BIOFILM REACTORS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 406 5.0 REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 408
Chapter 9 Clarification 1.0 Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 414 2.0 Solids-Separation Choices . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 416 3.0 Design Approach . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 417 3.1 Types of Settling Regimes. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 418 3.1.1 Type I . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 419 3.1.2 Type II. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 419 3.1.3 Type III. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 419 3.1.4 Type IV. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 419
3.2 Special Considerations for Nutrient Removal Sludges. . . . . . . . . . . 3.3 Clarifier Enhancements. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 3.4 Wastewater Flocculation. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 3.5 Flocculation Criteria. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 3.6 Clarifier Design Details. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
419 420 422 424 429
3.6.1 Influent Column. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 431 3.6.2 Energy-Dissipating Inlet . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 431 3.6.3 Feed Well (Flocculating Type). . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 433 3.6.4 Side Water Depth, Clear Water Zone, and Overflow Rate . . . . . . . . . . . . . . . 435 3.6.5 Floor Slope . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 438 3.6.6 Effluent Weir and Launder. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 440 3.6.7 Sludge Collectors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 443 3.6.8 Sludge Hopper . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 445
3.7 Rectangular versus Circular Clarifiers . . . . . . . . . . . . . . . . . . . . . . . . . 445
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3.8 Design Example . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 446 3.9 Clarifier Following Moving-Bed Biofilm Reactor, Trickling Filter, Rotating Biological Contactor, and Biotower��������������������������448 3.9.1 Secondary (Integrated Fixed-Film Activated Sludge) Clarifiers. . . . . . . . . . . 451 3.9.2 Sludge Hopper . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 454 3.9.3 Process Performance. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 455
3.10 Other Considerations. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 458 3.10.1 Modeling. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 458 3.10.2 Interaction with Other Facilities . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 458 3.10.3 International Practices. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 458
4.0 REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 459
Chapter 10 Effluent Filtration 1.0 INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 463 2.0 Process performance . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 465 3.0 References . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 469
Chapter 11 Development and Application of Models for Integrated Fixed-Film Activated Sludge, Moving-Bed Biofilm Reactors, Biological Aerated Filters, and Trickling Filters 1.0 INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 473 2.0 MODELING . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 478 2.1 Numerical Approach Using Semi-Empirical Equations for Biofilm (Steady-State and Dynamic Simulation)��������������������������������478 2.1.1 Ammonium-Nitrogen Uptake Rate . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 480 2.1.1.1 Ammonium-Nitrogen Uptake Rate by Nitrifiers in Biofilm. . . . . . . . . 480 2.1.1.2 Biofilm Nitrification Rates from Pilot Studies . . . . . . . . . . . . . . . . . . . 482 2.1.1.3 Ammonium-Nitrogen Uptake Rate by Nitrifiers in Mixed-Liquor Volatile Suspended Solids ��������������������������������������������485 2.1.1.4 Mass Balance for Ammonium-Nitrogen in Each Reactor. . . . . . . . . . . 490 2.1.2 Chemical Oxygen Demand Removal. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 494
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2.1.3 Biomass Production . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 499 2.1.3.1 Mixed-Liquor Volatile Suspended Solids . . . . . . . . . . . . . . . . . . . . . . . 500 2.1.3.2 Biofilm. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 501 2.1.4 Fraction of Nitrifiers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 502 2.1.5 Denitrification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 503 2.1.6 Oxygen. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 504
2.2 Numerical Approach to Solve One- and Two-Dimensional BiofilmDiffusion Models (Steady-State and Dynamic Simulation)��������������504 2.2.1 Ammonium-Nitrogen. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 506 2.2.2 Linkage to Equations 11.1 to 11.42 Presented Earlier. . . . . . . . . . . . . . . . . . . 508 2.2.3 Chemical Oxygen Demand, Biomass (Volatile and Total Suspended Solids), Dissolved Oxygen, and NOx-N ��������������������������������������������������������510 2.2.4 Biofilm Thickness, Growth, and Fraction Nitrifiers. . . . . . . . . . . . . . . . . . . . . 510
3.0 MODEL APPLICATIONS TO FULL-SCALE FACILITIES . . . . . . . . . . . 512 3.1 Integrated Fixed-Film Activated Sludge Plant Description and Modeling ������������������������������������������������������������������������������������������513 3.1.1 Integrated Fixed-Film Activated Sludge Plant Description. . . . . . . . . . . . . . 514 3.1.2 Integrated Fixed-Film Activated Sludge Plant Operation. . . . . . . . . . . . . . . . 515 3.1.2.1 Data from December 2006. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 515 3.1.2.2 Flow and Recycle. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 515 3.1.2.2.1 Primary Effluent.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 515 3.1.2.2.2 Aerobic Cells. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 518 3.1.2.2.3 Secondary/Plant Effluent. . . . . . . . . . . . . . . . . . . . . . . . . . . . 518 3.1.2.2.4 Discussion of the Data.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 518 3.1.3 Modeling Integrated Fixed-Film Activated Sludge in Aquifas . . . . . . . . . . . . 518 3.1.3.1 Results from Aquifas. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 519 3.1.3.2 Key Inputs to Aquifas Biofilm One-Dimensional Model . . . . . . . . . . . 526 3.1.3.3 Discussion of Aquifas Model and Accuracy of Results. . . . . . . . . . . . . 526 3.1.4 Modeling in BioWin. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 527 3.1.4.1 Framework. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 527 3.1.4.2 Results from BioWin.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 531 3.1.4.3 Discussion of Results from BioWin. . . . . . . . . . . . . . . . . . . . . . . . . . . . 531
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3.2 Moving-Bed Biofilm Reactor Plant Description and Modeling. . . . . 534 3.2.1 Moving-Bed Biofilm Reactor Modeling with GPS-X. . . . . . . . . . . . . . . . . . . . 537 3.2.1.1 Introduction. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 537 3.2.1.2 Example. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 542 3.2.2 Moving-Bed Biofilm Reactor Modeling with Aquifas. . . . . . . . . . . . . . . . . . . 542 3.2.3 Moving-Bed Biofilm Reactor Modeling—General Comments. . . . . . . . . . . . . 543 3.2.4 Integrated Fixed-Film Activated Sludge and Moving-Bed Biofilm Reactor Modeling—General Observations������������������������������������������������������552
4.0 REFERENCEs . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 553
Index . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 559
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List of Figures Figure 2.1 2.2 2.3 2.4 2.5 2.6
2.7 2.8
2.9
2.10
2.11 3.1 3.2
Page
Coenocytic cell structure typical of many fungi, in which cell contents, or cytoplasm, is multinucleate and continuous throughout the organism. . . . . . . . . . . . . 22 Schematic representation of the bacterial cell.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 29 Examples of organic compounds that are toxic and/or difficult to biooxidize.. . . . . . 37 Repeating glucose units of starch and glycogen connected by α-glycosidic bonds and cellulose connected by β-glycosidic bonds. . . . . . . . . . . . . . . . . . . . . . . . . . . 38 Schematic of reactions through which pyrite [FeS2(s)] is oxidized in mine drainage. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 39 Bacterial growth curve for a batch system in which nutrients initially are not limiting. As growth begins, the increase in cell number (or mass) is dependent only on the number (or mass) of cells present, and the system behaves in a first-order autocatalytic manner.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 52 Relationships between organic removal, cell mass production, and OUR.. . . . . . . . . . 58 Structured models of metabolism include reactions or steps in the metabolic process that can be identified or separated: (a) the rate of metabolism is separated from the rate of organic removal from the liquid and (b) a somewhat more complex storage mechanism is shown, which could be used to account for several observed rates of growth and oxygen uptake.. . . . . . . . . . . . . . . . . . . . . . . . . . . 60 Effect of temperature on organic removal rate and microbial growth in example 3. Initial BODU and suspended solids concentrations were 400 mg/L and 30 mg/L, respectively. The initial conditions provide for near-exponential growth because K << Co.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 62 Effect of inhibitant concentration on specific removal rate for two conditions: (a) substrate inhibition by a metabolizable inhibitant, represented by eq 2.22, and (b) a nonmetabolizable inhibitant, such as a heavy metal, represented by eq 2.23.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 64 Steps in the removal and breakdown of an organic particle in a biofilm.. . . . . . . . . . . 66 Typical trickling filter components and cross-section.. . . . . . . . . . . . . . . . . . . . . . . . . . . 85 Hydraulically driven rotary distributors use variable frequency drive controlled gates that either open or close distributor orifices, which adjust with varying pumped flowrates to maintain a constant preset rotational speed, and electrically driven rotary distributor. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 86
xxv
xxvi List of Figures
3.3
3.4
3.5
3.6 3.7
3.8
3.9 3.10 3.11 3.12 3.13 3.14 3.15 3.16
3.17 3.18
3.19 3.20 3.21 3.22
Adjustable plastic stanchions and FRP grating on the concrete floor of a bolted steel containment structure, and an HDPE mat used to support random synthetic media.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 91 Typical flow diagrams for the trickling filter process: (a) and (b) single-stage trickling filter process, (c) two-stage trickling filter process, and (d) two-stage trickling filter process with intermediate clarification. . . . . . . . . . . . . . . . . . . . . . . . . . . 94 Typical flow diagrams for trickling filter-suspended growth processes: (a) activated biofilter, (b) trickling filter/solids contact or trickling filter/activated sludge, (c) activated biofilter/solids contact or activated biofilter/activated sludge, and (d) trickling filter and activated sludge.. . . . . . . . . . . . 95 Three-dimensional representation of liquid flow distribution in cross-flow media, both (a) with and (b) without biofilm.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 98 (a) Trickling filter underdrain and media supports with air distribution piping throughout the underdrain, and (b) typical low-pressure fan applied to provide process air.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 102 Effect of SRT in a suspended-growth reactor and nitrification efficiency in an upstream trickling filter on effluent ammonia-nitrogen concentration in a combined TF/SG reactor. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 115 Nitrification efficiency versus organic loading in rock-media trickling filters. . . . . . 116 Nitrification efficiency as a function of filtered, or soluble, BOD5 in the effluent of a vertical-flow media trickling filter at Stockton, California.. . . . . . . . . . . 117 Nitrification efficiency versus organic loading in trickling filters.. . . . . . . . . . . . . . . . 118 Nitrification efficiency versus organic loading in trickling filters.. . . . . . . . . . . . . . . . 119 Effect of bulk liquid TSS concentration on nitrification in a pilot-scale NTF.. . . . . . . 122 Transition NH3-N concentrations as function of temperature.. . . . . . . . . . . . . . . . . . . 125 Actual and predicted effluent from an NTF. Predicted effluent was calculated using the modified Gujer and Boller (1986) model.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 131 (a) Typical hydraulically driven rotary distributor with reverse-thrusting jets and electric drive, and (b) modern hydraulically driven rotary distributor with variable-frequency controlled gates that close or open orifices (therefore, the flowrate contributes to reverse thrust).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 134 (a) Modern hydraulically driven rotary distributor, (b) gate controller, (c) orifices, and (d) full-length arm and distributor.. . . . . . . . . . . . . . . . . . . . . . . . . . . . 137 (a) “Pouch” snail typically found in trickling filters treating m unicipal wastewater, and (b) a modular plastic media sheet with biofilm and snail-grazing pathways.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 145 Snail shells covering fine-bubble diffusers in suspended-growth reactors that are located in (a) Pueblo, Colorado; and (b) Garland, Texas.. . . . . . . . . . . . . . . . . . . . . 146 Biofilm and filter fly control with low-frequency dosing.. . . . . . . . . . . . . . . . . . . . . . . 148 Nitrifying trickling filter operating modes for high-concentration undissociated aqueous ammonia dosing. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 151 Nitrification rate and biomass accumulation in pilot-scale NTFs after startup.. . . . . 155
List of Figures
3.23 4.1 4.2 4.3
4.4
4.5 4.6 4.7 4.8
4.9 4.10 5.1 5.2 5.3
5.4 5.5 5.6
5.7 5.8
Three modes of TF/SC process operation.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 160 General representation of RBC process.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 175 RBC process general flow arrangements.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 177 Clark RBC predictive model, Monod kinetics, BOD5 removal: (1) values are total BOD5; (2) 50% sBOD5 assumed; (3) >13°C (>55°F) temperature; (4) three stages, with 50% of media area in the first stage; and (5) standard-density media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 186 Opatken RBC predictive model, second-order kinetics, BOD5 removal: (1) values are total BOD5; (2) >13°C (>55°F) temperature; (3) 50% sBOD5 assumed; (4) three stages, with 50% of media area in the first stage; and (5) standard-density media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 187 Benjes RBC predictive model, BOD5 removal: (1) values are total BOD5; (2) >13°C (>55°F) temperature; and (3) standard-density media.. . . . . . . . . . . . . . . . . 188 Manufacturers’ RBC design basis, BOD5 removal: (1) values are total BOD5; (2) >13°C (>55°F) temperature; and (3) 50% sBOD5 assumed. . . . . . . . . . . . . . . . . . . . 189 Equipment manufacturers’ predictions versus full-scale plant data, BOD5 removal.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 191 Benjes RBC model predictions versus full-scale plant data, BOD5 removal: (1) values are total BOD5; (2) k = 0.30; (3) >13°C (>55°F) temperature; and (3) standard-density media. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 192 Benjes RBC predictive model, effect of k rate on predicted performance (standard-density media assumed).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 193 Manufacturers’ design basis, nitrification of domestic wastewater. . . . . . . . . . . . . . . 196 Vertical mounted flat-panel sieves with air sparge system and horizontal sieves located over aeration grid.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 216 (a) Aeration grid and distribution piping, and (b) stainless diffuser with 4-mm aeration-holes along underside.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 216 Representative carriers removed from each of four moving-bed reactors in series showing the variation in biofilm color (specialization and biofilm concentration (active biofilm) dependent on the operating condition and treatment function of each reactor. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 218 (a) COD removal efficiency at high loading conditions, and (b) poor settling character of biofilm slough under high loading rates. . . . . . . . . . . . . . . . . . . . . . . . . . . 221 Soluble BOD removal rate as a function of total BOD loading in roughing MBBR application. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 222 (a) Effect of BOD loading and dissolved oxygen on nitrification rates at 15°C (adapted from Hem et al. [1994]), and (b) difference in nitrification rate between multiple MBBRs in series.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 223 Influence of dissolved oxygen at low ammonia concentrations. . . . . . . . . . . . . . . . . . 225 (a) Seasonal biomass concentration and temperature in tertiary nitrifying MBBR, and (b) nitrification activity grouped by temperature condition as a function of dissolved oxygen level.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 228
xxvii
xxviii List of Figures
5.9 5.10
5.11 5.12 5.13 5.14 5.15 5.16 5.17 5.18 5.19 5.20 5.21 5.22 5.23 5.24 5.25 5.26 5.27 6.1 6.2
6.3 6.4 6.5 6.6 6.7 6.8 6.9 6.10 6.11 6.12 6.13
Denitrification rate as a function of temperature with different external carbon types.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 230 (a) ABS 123K mixers installed toward the surface of the reactor oriented with a 30-degree negative inclination to “push” the media down into the depths of the reactor, and (b) post-denitrification MBBR reactor in operation at the Sjölunda WWTP (Malmö, Sweden).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 232 Simplified liquid treatment schematic of Moa Point WWTP (New Zealand). . . . . . . 236 BOD profile results at Moa Point WWTP (New Zealand).. . . . . . . . . . . . . . . . . . . . . . . 237 (a) Test periods 1, 2, and 3, and (b) test periods 4 through 9. . . . . . . . . . . . . . . . . . . . . 238 Ammonia removal rate as function of loading rate in a tertiary MBBR.. . . . . . . . . . . 239 Simplified liquid treatment schematic of Moorhead WWTF (Minnesota).. . . . . . . . . 240 Ammonia profile results across nitrifying MBBR at the Moorhead WWTF (Minnesota).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 241 Simplified liquid treatment schematic of Williams Monaco WWTP (Colorado).. . . . 242 Two MBBR trains with four reactors each: anoxic and aerated.. . . . . . . . . . . . . . . . . . 243 BOD profile results across MBBRs at Williams Monaco WWTP (Colorado) (2004 to 2007). . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 244 Settled effluent TSS at Williams Monaco WWTP (Colorado) (2004 to 2007). . . . . . . . 245 Simplified liquid treatment schematic of Klagsham WWTP (Sweden).. . . . . . . . . . . . 247 Effluent BOD7 results at Klagsham WWTP (Sweden).. . . . . . . . . . . . . . . . . . . . . . . . . . 249 Effluent total nitrogen results at Klagsham WWTP (Sweden).. . . . . . . . . . . . . . . . . . . 249 Effluent total phosphorus results at Klagsham WWTP (Sweden).. . . . . . . . . . . . . . . . 249 Simplified liquid treatment schematic of Gardemoen WWTP (Norway).. . . . . . . . . . 251 Ammonia profile results across MBBRs at Gardemoen WWTP (Norway) (2001).. . . 252 Nitrogen profile results across MBBRs at Gardemoen WWTP (Norway) (2001).. . . . 252 IFAS process versus conventional activated sludge process. . . . . . . . . . . . . . . . . . . . . 260 Effect of biofilm thickness on locations of growth and surface area of biofilm: (a) magnitude of difference in biofilm surface area for “thin” and “thick” biofilms on the same carrier particle of moving-bed or IFAS system, and (b) magnitude of difference in biofilm surface area for “thin” and “thick” biofilms on the same carrier particle of moving-bed or IFAS system. . . . . . . . . . . . . . 279 Typical layout of plastic-carrier-media IFAS systems.. . . . . . . . . . . . . . . . . . . . . . . . . . 282 Typical layout of sponge-media IFAS system.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 283 Typical layout of cord or web fixed-bed-media IFAS systems.. . . . . . . . . . . . . . . . . . . 283 Pilot study configuration.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 290 Schematic of BNR upgrade to IFAS. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 292 Process configuration for BNR upgrade. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 292 Westerly, Rhode Island WWTF IFAS system flow diagram.. . . . . . . . . . . . . . . . . . . . . 300 Westerly, Rhode Island, WWTF IFAS system reactor train plan.. . . . . . . . . . . . . . . . . 301 Full-scale flow diagram. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 309 Monthly operating data. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 314 Polyester fabric media.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 316
List of Figures
6.14 6.15 6.16 6.17 6.18 7.1 7.2 7.3 7.4 7.5 7.6 7.7 7.8 7.9 7.10 7.11 7.12 7.13 7.14
7.15 7.16 7.17 7.18 7.19 7.20 8.1 9.1 9.2
9.3 9.4 9.5
Existing and proposed aeration basin configuration.. . . . . . . . . . . . . . . . . . . . . . . . . . . 317 Proposed process flow diagram. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 317 Installation of 10 IFAS media modules in existing basin.. . . . . . . . . . . . . . . . . . . . . . . . 319 Effluent cBOD concentration after train B startup.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 319 Effluent ammonia concentration (7-day average).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . .320 BAF process flow diagrams.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 328 BioCarbone® biological filter. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 332 Downflow denitrification filter.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 333 Biofor® upflow biological filter. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 334 Biostyr® process arrangement for nitrification and denitrification. . . . . . . . . . . . . . . . 336 Schematic of moving-bed denitrification filter.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 338 Schematic of non-backwashing, open-structure media filter.. . . . . . . . . . . . . . . . . . . . 339 Mineral media upflow SAF with block underdrain.. . . . . . . . . . . . . . . . . . . . . . . . . . . . 340 Carrier diameter versus media specific surface area.. . . . . . . . . . . . . . . . . . . . . . . . . . . 344 Relation between total solids of backwash liquor and time for different backwash airflow rates.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 347 Relationship between total solids of backwash liquor and time for d ifferent backwash water flowrates. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 347 Nitrification rate for differently pretreated raw wastewaters as a function of C:N ratio. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 356 Maximum nitrification versus long-term average loading.. . . . . . . . . . . . . . . . . . . . . . 359 BAF configurations for combined nitrification and denitrification: (a) pre-denitrification/nitrification and optional post-denitrification, and (b) pre-denitrification/nitrification in one reactor.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 363 Nitrification with flocculated raw wastewater.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 363 Denitrification filter design curves using EBDT with additional data points. . . . . . . 369 VEAS WWTP schematic.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 380 Seine Centre WWTP liquid treatment schematic.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 382 Frederikshavn Central WWTP liquid treatment schematic. . . . . . . . . . . . . . . . . . . . . . 386 Nitrification and denitrification BAF reactors at West Warwick, Rhode Island.. . . . 388 Reactor configurations: (a) UASB, (b) EGSB, (c) BAS, and (d) internal circulation. . . . . 403 BOD5 content of effluent versus TSS and trickling filter loading.. . . . . . . . . . . . . . . . . 415 (a) Suspended solids concentration and turbidity data collected throughout the study with a linear relationship function and 95% suspended solids concentration prediction interval estimate; (b) turbidity/time data collected from the Camp Creek plant on August 13, 1990, with batch flocculation curve fit; (c) turbidity/time data collected at the Utoy Creek plant on July 19, 1990, with batch flocculation curve fit; and (d) turbidity/time data collected at the Coneross plant on May 30, 1991, with batch flocculation curve fit.. . . . . . . . . . . . . . . 425 Effect of coarse-bubble aeration rate on effluent solids.. . . . . . . . . . . . . . . . . . . . . . . . . 426 Flocculation procedures, test number 1.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 427 Flocculation procedures, test number 2a.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 428
xxix
xxx List of Figures
9.6 9.7 9.8 9.9 9.10 9.11 9.12 9.13 9.14 9.15 10.1 10.2 10.3 10.4
EDI and hydraulic flocculating feed well. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 432 General arrangement of the improved clarifier inlet structure. . . . . . . . . . . . . . . . . . . 433 Dense sludge blanket profile: average and peak flows.. . . . . . . . . . . . . . . . . . . . . . . . . 435 Clarifier detention: efficiency versus CWD.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 436 Recommendations for primary clarifier floor configuration.. . . . . . . . . . . . . . . . . . . . . 439 Clarifier improvement with Stamford baffle.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 441 McKinney (Lincoln)/Stamford baffle arrangements.. . . . . . . . . . . . . . . . . . . . . . . . . . . 443 General arrangement of sludge hopper.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 446 General arrangement of improved wastewater clarifier.. . . . . . . . . . . . . . . . . . . . . . . . 447 EDI primary clarifier.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 450 Cloth filters.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 465 Traveling-bridge filter.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 465 Compressible media filters.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 466 Performance data for six different types of granular-medium filters tested on the same activated sludge effluent at a filtration rate of 2.7 L/m2 s (4 gpm/sq ft).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 467 11.1 Plant configurations. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 474 11.2 Schematic of IFAS and MBBR system with 10 cells (reactors) in series operating in enhanced nutrient removal configuration.. . . . . . . . . . . . . . . . . . . . . . . . . 475 11.3 Modeling biofilm in (a) trickling filter, (b) MBBRs, and (c) BAF.. . . . . . . . . . . . . . . . . 476 11.4 Nitrification rates for biofilm in IFAS and MBBR systems: (a) biodegradable COD limiting nitrifiers in biofilm, ammonium-nitrogen is not limiting, liquid temperature = 12°C, soluble biodegradable COD >10 mg/L, NH4-N > 3 mg/L, measured at a dissolved oxygen concentration of 8 to 9 mg/L; (b) liquid temperature = 12°C, soluble biodegradable COD < 10 mg/L, NH4-N < 5 mg/L, measured at dissolved oxygen concentration of 8 to 9 mg/L.. . . . . . . . . . . . . . . . . . . . 484 11.5 COD uptake rates for biofilm in IFAS and MBBR systems: liquid temperature = 12°C, measured at a dissolved oxygen concentration of 8 to 9 mg/L.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 496 11.6 Layout of Broomfield WWTP, Colorado.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 514 11.7 Aquifas output: The figures in the left column are from the semi-empirical model; the figures in the right column are from the biofilm one-dimensional model. The diffusion model can offer a higher degree of precision in its ability to predict day-to-day variations in a dynamic simulation, but takes substantially longer to run. Both models are able to predict the diurnal and 31-day average. . . . . . . . . . . 522 11.8 Substrate profiles for COD, NH4-N, and NO3N. The left column is from the semi-empirical model; the right column is from the biofilm one-dimensional model.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 523 11.9 COD uptake, NH4-N uptake, and denitrification in the biofilm and MLVSS. . . . . . . 524 11.10 Substrate profiles (dissolved oxygen, NH4-N, NOx-N, SCODbio, and percent VSS) inside the biofilm in aerobic cells 1 and 2.. . . . . . . . . . . . . . . . . . . . . . . . . 525
List of Figures
11.11 11.12 11.13 11.14 11.15 11.16 11.17 11.18
11.19 11.20 11.21 11.22 11.23 11.24
Typical raw influent COD fractions in BioWin. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 528 Screen shot of IFAS model in BioWin.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 531 BioWin IFAS calibration results for effluent ammonia and nitrate.. . . . . . . . . . . . . . . 532 BioWin IFAS calibration results for MLSS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 532 BioWin IFAS calibration results for effluent BOD and TSS. . . . . . . . . . . . . . . . . . . . . . 533 Schematic of one of two MBBR trains at the South Adams plant, Adams County, Colorado���������������������������������������������������������������������������������������������������������������������������������� 534 Layout of MBBR as modeled in GPS-X������������������������������������������������������������������������������ 543 Comparison of MBBR effluent NH3-N and NOx-N from GPS-X model, with actual data. The NH3-N represents ammonium-nitrogen + ammonia-nitrogen; at pH ~7, >99% is as ammonium-nitrogen. ���������������������������������������������������������������������� 546 Comparison of MBBR secondary effluent BOD5 and TSS model results from GPSX model with actual data �������������������������������������������������������������������������������������������������������� 546 Comparison of MBBR effluent MLSS (before settling in secondary clarifier) from GPSX model with actual data.������������������������������������������������������������������������������������ 547 Layout of South Adams MBBR as modeled in AquaNET (Windows.NET) version of Aquifas.������������������������������������������������������������������������������������������������������������������������������ 547 Comparison of MBBR effluent NH3-N and NOx-N from Aquifas model with actual data������������������������������������������������������������������������������������������������������������������������������ 550 Comparison of MBBR secondary effluent BOD5 from Aquifas model with actual data������������������������������������������������������������������������������������������������������������������������������������������ 551 Comparison of MBBR effluent MLSS (before settling in secondary clarifier) from Aquifas model with actual data�������������������������������������������������������������������������������� 551
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List of Tables Table 2.1 2.2 2.3 2.4 2.5 2.6 2.7 2.8 3.1 3.2 3.3 3.4 3.5 4.1 4.2 5.1 5.2 5.3 5.4 5.5 5.6 5.7 5.8 5.9 5.10 5.11 5.12 5.13 5.14 5.15
Page
Subdivisions of organisms and principal groups found in biological treatment systems. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 18 Common human diseases caused by viruses. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 25 Typical composition of Escherichia coli.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 27 Typical elemental composition of bacterial cells.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .28 Classification of living organisms by carbon and energy sources.. . . . . . . . . . . . . . . . . 34 Bacterial metabolism and respiration.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 35 Standard reduction potentials at 25°C for selected environmentally important redox couples.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 47 BODU and suspended solids concentrations for the three temperature values.. . . . . . 63 Properties of some trickling filter media.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 88 Trickling filter classification.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 97 Operating and flushing dosing rates for distributors.*. . . . . . . . . . . . . . . . . . . . . . . . . . 100 Reported zero-order nitrification rates for vertical- and cross-flow media. . . . . . . . . 121 Design criteria for selected combined TF/SG processes.. . . . . . . . . . . . . . . . . . . . . . . . 157 Comparison of organic loadings for models.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 190 Media surface adjustment factors.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 194 Plastic biofilm carriers.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 215 General summary of treatment schemes offered by MBBRs.. . . . . . . . . . . . . . . . . . . . .219 Typical BOD design loading criteria.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 220 Average results from normal-rate MBBRs with chemical phosphorus removal.. . . . 222 Examples of reported O2:NH4-N.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 226 Examples of reported temperature-dependency coefficients.. . . . . . . . . . . . . . . . . . . . 227 Examples of typical pre-denitrification rates observed with municipal wastewater.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 229 Screening examples at MBBR installations. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 233 Solids separation examples at MBBR installations.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 233 MBBR/solids contact process at Moa Point WWTP (New Zealand).. . . . . . . . . . . . . . 236 Tertiary MBBR process at Moorhead WWTF (Minnesota).. . . . . . . . . . . . . . . . . . . . . . 240 Multistage MBBR process at Williams Monaco WWTP (Colorado). . . . . . . . . . . . . . . 243 Post-denitrification MBBR process at Klagsham WWTP (Sweden).. . . . . . . . . . . . . . . 247 Multistage MBBRs at Gardemoen WWTP (Norway).. . . . . . . . . . . . . . . . . . . . . . . . . . . 251 Summary of annual performance showing that the process meets performance requirements.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 253
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xxxiv List of Tables
6.1 6.2 6.3 6.4 6.5 6.6 6.7 6.8 6.9 6.10 6.11 6.12 6.13 6.14 7.1 7.2 7.3 7.4 7.5 7.6 7.7 7.8 7.9 7.10 7.11 7.12 7.13 7.14 7.15 9.1 9.2 9.3 9.4 9.5 9.6 9.7
Summary of media characteristics. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 280 Biofilm SSA of various types of media.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 281 Design criteria, original Annapolis WRF.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 290 Design criteria, BNR upgrade.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 293 Operating MLSS MCRTs, HRTs, media, and average performance for IFAS. . . . . . . 293 Performance data.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 294 Summary of discharge permit limits. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 303 Design criteria for the Westerly LINPOR-CN system.. . . . . . . . . . . . . . . . . . . . . . . . . . 303 Summary of operating data.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 304 Effluent requirements.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 307 Specific design criteria for the upgraded plant.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 309 Influent data.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 310 Effluent data.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 312 Colony, Texas, 2004 historical influent loadings and effluent discharge. . . . . . . . . . . 315 Commercially available BAF reactor systems and media.. . . . . . . . . . . . . . . . . . . . . . . 342 Summary of baf backwashing requirements.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 348 Typical BAF loading rates for secondary treatment.. . . . . . . . . . . . . . . . . . . . . . . . . . . . 352 Effect of water velocity on nitrification rate for three types of biofilters. . . . . . . . . . . 356 Influence of grain size and shape on performance from testing in Oslo, Norway. . . 357 Typical BAF loading rates for nitrification. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 358 Typical biologically active filter loading rates (BAF) for pre-denitrification.. . . . . . . 365 Typical biologically active filters (BAF) loading rates for post-denitrification. . . . . . 366 Performance results for two-stage BAF in Oslo, Norway.. . . . . . . . . . . . . . . . . . . . . . . 381 Year 2000 average performance for three-stage BAF system at Seine Center WWTP.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 383 Year 2000 average performance for first-stage BAF system at Seine Center WWTP.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 384 Year 2000 average performance for second-stage BAF system at Seine Center WWTP.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 384 Year 2000 average performance for third stage BAF system at Seine C enter WWTP.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 385 Year 1998 average performance for the two-stage BAF at F rederikshavn Central WWTP.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 387 Havelock WWTP comparison of effluent BOD and total nitrogen average and variation. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 390 Clarification characteristics of trickling filter systems in municipal WWTPs. . . . . . . 421 Energy gradient (G) values as a function of aeration type. . . . . . . . . . . . . . . . . . . . . . . 427 Clarifier design features/interactions. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 430 Example minimum floor slopes versus diameter for clarifiers with scrapers.. . . . . . 440 Results of full-scale study of launder positioning. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 442 Design basis for primary and secondary clarifiers.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . 447 Review of secondary clarifier operating conditions.. . . . . . . . . . . . . . . . . . . . . . . . . . . . 452
List of Tables
9.8 9.9 10.1 11.1 11.2 11.3 11.4 11.5 11.6 11.7 11.8 11.9 11.10 11.11 11.12 11.13 11.14 11.15 11.16
Effluent quality of single-stage trickling filter nitrification plants.. . . . . . . . . . . . . . . . 456 Clarified/filtered effluent quality of Ohio nitrifying trickling filter plants.. . . . . . . . 457 Full-scale filtration of settled trickling filter effluent.. . . . . . . . . . . . . . . . . . . . . . . . . . . 468 Biofilm specific areas and modeling trickling filter, TF/AS, MBBR, and BAFs within models.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 478 Typical biofilm yields. Yields are generated by running a biofilm one-dimensional model or from measurements in pilot studies.. . . . . . . . . . . . . . . . . 480 Values of coefficients measured in pilot studies and the default values for the semi-empirical model.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 486 Matrix showing kinetic coefficients for MLVSS in semi-empirical and biofilm models.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 493 Semi-empirical modeling of the biofilm—equations for semi-empirical computation of biofilm flux.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 495 Actual plant data used for IFAS process model.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 516 Comparison of 31-day average from dynamic simulation against plant data Aquifas semi-empirical and biofilm one-dimensional models.. . . . . . . . . . . . . . . . . . . 520 Measured and computed values of biofilm growth (Aquifas biofilm one-dimensional model).. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 525 Summary of the BioWin influent fractions. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 529 Steady-state BioWin calibration results for a full-scale IFAS facility.. . . . . . . . . . . . . . 530 SA-T1: MBBR influent data and cell data for dissolved oxygen, January 2007.. . . . . 535 SA-T2: MBBR effluent data. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 538 SA-T3: Principal GPS-X biofilm parameters. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 540 SA-T4: Summary of GPS-X influent fractions (Mantis and ASM2d).. . . . . . . . . . . . . . 540 South Adams County (Colorado) steady-state results for January 2007 with GPS-X.. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 544 South Adams County—steady state results for January 2007 with Aquifas. . . . . . . . 548
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Preface The purpose of this Manual of Practice (MOP) is to provide a detailed understanding of all aspects of wastewater treatment using fixed-film processes. By way of background, an overview of the historical development of fixed-film treatment systems is provided. The microbiology of attached growth biofilms is presented as background to better understand the design and operation of the various fixed-film technologies. The following types of pure fixed-film and hybrid treatment systems are discussed in detail: • Trickling filters, • Rotating biological contactors, • Moving-bed biofilm reactors, • Integrated fixed-film activated sludge, and • Biological filters. Both aerated and anoxic applications of the biological filter systems are included. The chapter on trickling filters discusses combined fixed-film and suspended-growth processes. There is a chapter dedicated to new and emerging fixed-film treatment processes, including such processes as upflow anaerobic sludge blanket reactors and anammox biofilm reactors. Factors that affect the design of the various processes, appropriate design criteria and procedures, modelling techniques, equipment available, and construction features are discussed. Operational issues associated with each type of process are presented, including potential problems and corrective actions. Case studies are included to illustrate the application of these technologies. Finally, there is a detailed discussion of modeling techniques as applied to fixed-film and hybrid processes. The chapter discusses the use of several commercially available models.
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xxxviii
Preface
The overall objective of this book is to provide the reader with a complete understanding of the types of pure and hybrid fixed-growth treatment processes available and an understanding of their design, performance, and operational issues. This Manual of Practice was produced under the direction of Rhodes R. Copithorn, P.E., BCEE, Chair. The principal authors of this Manual of Practice are as follows: Chapter 1 Chapter 2 Chapter 3 Chapter 4 Chapter 5 Chapter 6 Chapter 7 Chapter 8 Chapter 9 Chapter 10 Chapter 11
Rhodes R. Copithorn, P.E., BCEE, Chair Stefan Wuertz, Ph.D. Edward D. Schroeder, Ph.D. Joshua P. Boltz, Ph.D., P.E. Joshua P. Boltz, Ph.D., P.E. James P. McQuarrie, P.E. Rhodes R. Copithorn, P.E., BCEE, Chair Dipankar Sen, Ph.D., P.E. Christine deBarbadillo, P.E. Joseph A. Husband, P.E., BCEE Frank Rogalla Christopher W. Tabor, P.E. Stephen Tarallo Robert Nerenberg, Ph.D., P.E. Thomas E. Wilson, Ph.D., P.E., BCEE Joseph A. Husband, P.E., BCEE Dipankar Sen, Ph.D., P.E. Heather Phillips, P.E.
Authors’ and reviewers’ efforts were supported by the following organizations: AECOM, Alexandria, Virginia Aqualia, Madrid, Spain Black & Veatch Corporation, Gaithersburg, Maryland, and Kansas City, Missouri Brentwood Industries, Reading, Pennsylvania CDM, Newport News, Virginia, and Portland, Oregon CH2M Hill, Philadelphia, Pennsylvania; Tampa, Florida; Redding, California Earth Tech, Alexandria, Virginia Entex Technology, Inc., Chapel Hill, North Carolina
Preface
GHD, Bowie, Maryland Greeley and Hansen, LLC, Wilmington, Delaware Hydromantis, Inc., Hamilton, Ontario, Canada Malcolm Pirnie, Inc., Columbus, Ohio, and White Plains, New York Regional Municipality of Waterloo, Ontario, Canada Thomas E. Wilson Environmental Engineers LLC, Barrington, Illinois University of California, Davis, California University of Notre Dame, Notre Dame, Indiana Woodard & Curran, Cheshire, Connecticut, and Dedham, Massachusetts
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Biofilm Reactors
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Chapter 1
Introduction 1.0 Background and Purpose
3.4 Coupled Trickling Filter/ Activated-Sludge Process 6 3.5 Biological Filters 7 3.6 Hybrid Processes 7
1
2.0 Characteristics of FixedGrowth Processes 2 3.0 History 3.1 Contact Beds 3.2 Trickling Filters 3.3 Rotating Biological Contactors
3 4 4
4.0 Organization of Manual
10
5.0 References
11
5
6.0 SUGGESTED READINGS
14
1.0 Background and Purpose This manual of practice is an update of the special publication titled Aerobic FixedGrowth Reactors, previously published by the Water Environment Federation® (Alexandria, Virginia) (WEF) in the year 2000. A number of changes in our understanding of biofilm processes and the technology and regulations that drive the wastewater treatment industry prompted this decision to rewrite the manual. Techniques developed in the field of molecular biology have expanded our understanding of biofilm kinetics and their microbial ecology. Regulations are imposing increasingly more stringent limits on the concentration of pollutants that may be discharged, including, most importantly, nutrients. The imposition of nitrogen and phosphorus limits has dramatically changed the design and operation of wastewater treatment facilities. These changes have brought about advances in the technology applied to 1
2
Biofilm Reactors
wastewater treatment, including many significant changes to the type of fixed-film treatment processes that are available today. The change in the title of this manual from Aerobic Fixed-Growth Reactors to Biofilm Reactors acknowledges some of these advances. The manual now incorporates both aerobic and anoxic fixed-film treatment processes that are applied for biological carbon and nitrogen removal.
2.0 Characteristics of Fixed-Growth Processes Although the basic metabolic processes that biological systems use to remove carbon and nutrients in wastewater treatment plants (WWTPs) are the same for fixed-film and suspended-growth systems, there are some inherent differences that provide several advantages and some challenges for the application of fixed-film processes. It is difficult to generalize, because each specific application has its own set of advantages and disadvantages, as will become apparent in the discussion of the various fixed-growth technologies discussed in this manual. However, a general summary of the differences would be useful at this point. Suspended-growth systems are comprised of biological flocs, but, theoretically, all of the dissolved substrate is available to all of the cells. With fixed-growth systems, the substrates must diffuse through the biofilm layers to become available; thus, the transport of substrate from the bulk liquid through the stagnant boundary layer and into the biofilm through the process of diffusion becomes another limiting factor. End products of the metabolic reactions must diffuse in the reverse direction. Thus, a cross-section through a fully developed biofilm will exhibit varying environmental and kinetic characteristics. A single biofilm may have aerobic, anoxic, and anaerobic processes occurring, and the substrate that becomes limiting will change through the depth of the biofilm. Thus, fixed-film processes are quite complex to model. Some advantages generally associated with fixed-film processes include the following: • Reduced operating and energy costs, • Smaller reactor volume, • Minimized need for settling capacity, and • Operational simplicity.
Introduction
Disadvantages are more specific to each technology, and they will be discussed in subsequent chapters. However, some general concerns associated with fixed-film processes include the following: • Potential clogging of the media system as a result of inadequate screening; • Excessive growth, which could plug the media system or cause free-floating media to sink; and • Inadequate mixing or short-circuiting, resulting in inefficient use of the media.
3.0 History Biological wastewater treatment technology has advanced tremendously since its early roots as a primitive science in the late 1800s. Much of the impetus was provided by increasingly crowded urban centers and the outbreak of typhoid resulting from a lack of understanding of public health issues. As the science developed, there was an early understanding that aeration of the wastewater was beneficial. An excellent presentation of the origins and development of fixed-film treatment processes has been provided by Alleman and Peters (1982). Initial experiments with this practice yielded nominal results, most likely because of a lack of understanding regarding the need for an active biomass. Aeration of fixed-film processes provided better results because of the presence of an established biomass. This was understood more fully later, as Gilbert J. Fowler of the University of Manchester, England, and his associates, Edward Ardern and William T. Lockett, in the first few decades of the 1900s, experimented with various processes, which led to our understanding of activated sludge (Cooper, 2001). Fixed-film processes continued to evolve as experimenters substituted wooden lathes in place of coarse rock, in an effort to increase the available surface area. Much of this work was advanced in the United Kingdom at the Manchester Sewage Works and at the Lawrence Experiment Station in Lawrence, Massachusetts (now known as the Senator William X. Wall Experiment Station). Biological filters originally were developed by the Lawrence Experiment Station in Massachusetts (Mills, 1890), using gravel for the biological media. These results proved that the process was not merely mechanical filtration, but also involved removal of pollutants by biological growth on the media. The results showed a correlation between the media volume and the volume of wastewater that could be treated.
3
4
Biofilm Reactors
From this beginning, the process evolved to be the dominant secondary wastewater treatment process in the United States and elsewhere by the 1950s. The introduction of synthetic media in the 1950s resulted in extending the use of fixed-growth concepts. Continued research led to the development of high-rate systems, such as biological aerated filters, moving-bed biofilm reactors, and various hybrid systems that combine the advantages of both suspended-growth and fixed-film processes.
3.1 Contact Beds The 1890s studies at the Lawrence Experiment Station in Massachusetts were viewed with interest in the United Kingdom. Corbett (1902) developed a trickling filter design that was the forerunner of today’s trickling filter systems. There was provision for distributing the influent across the surface of the gravel bed and ventilation of the underdrain. Independently, Crimp (1890) and Dibdin (1903) conducted studies that led to the development and use of a process known as a contact bed. Crimp and Dibdin constructed a basin, filled it with slag media, provided a means of flooding the basin, and then slowly drained it after approximately 1 hour of contact. The tank then was allowed to stand empty for 4 to 6 hours, which permitted the organics on the surface of the media to be oxidized. To prevent plugging of the slate media, the wastewater was screened or chemically precipitated before application. As a result, 75% removal of the oxidizable organic matter in the wastewater at a hydraulic loading rate of 1.2 m3/m2·d was obtained. The process evolved further by adding stages. However, clogging remained a continuing problem that was not resolved easily. The demise of the contact bed was initiated when the Royal Commission on Sewage Disposal report (1908) was published, which showed that a trickling filter could process twice as much wastewater per unit of volume as a two-stage contact bed.
3.2 Trickling Filters The emergence of the trickling filter was aided by the development of an effective means of distributing the flow to the filter media. Caink (1897) and Candy (1898) advanced Mills’ Lawrence Experiment Station concept into a rotating arm distribution system driven by water jets. Around the same time, a reciprocating distributor was developed for the rectangular filters, and mechanical drives (engine and electrically driven) were in use by 1904 (Stanbridge, 1972). The trickling filter process
Introduction
evolved further, and many different connotations were used to describe the process configurations. Following the initial evaluation by Dow Chemical Company (Midland, Michigan) from 1954 to 1955 (Bryan, 1955; Dow Chemical Company, 1955), the combined efforts of the Mead Corporation (Dayton, Ohio), Fluor Corporation (Irving, Texas), and Dow Chemical Company resulted in the development and application of both random and bundle synthetic media in the late 1950s. The history of the developmental stage of plastic media is described in reviews by Bryan (1982) and Peters and Alleman (1982). Different media configurations, producing different surface areas, continue to evolve today.
3.3 Rotating Biological Contactors Rotating biological contactors (RBCs) were an outgrowth of trickling filters and were influenced by the desire to reduce power consumption for wastewater treatment. Steels (1974) credits the first RBC concept to have been around 1900. A wooden cylinder with slatted walls was filled with brushwood and slowly rotated while approximately 50% submerged. Extensive testing of the RBC process was conducted in Germany. In 1900, Weigand patented a rotary cylinder made of wooden slats (Alleman and Peters, 1982). In the 1950s, researchers first used asbestos sheets, and, by the early 1960s, expanded polystyrene media came into use. The use of this lightweight material allowed the shafts holding the media to be lengthened. This resulted in a major growth of the RBC in the 1960s and 1970s, with more than 700 plants put into operation in Europe and the United States. This growth subsequently was stymied by several problems that developed, including performance of less than design expectations, excess biomass accumulations, shaft breakage, loping of disks (caused by unbalanced biomass weight), and undesirable biological growths. Most of the early developmental problems have been resolved, and there are many systems operating successfully, but the acceptance of RBCs as an effective treatment process has not returned to its former level among design engineers and owners. A submerged RBC, or submerged biological contactor (SBC), appeared in the 1980s, in which the discs are 70 to 90% submerged, and the shaft is driven by air. The intent was to decrease the loading on the shafts, improve biomass control, and provide an opportunity to retrofit existing activated sludge basins. The SBCs have been piloted for denitrification in an anoxic reactor. However, SBCs have seen only limited application.
5
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Biofilm Reactors
3.4 Coupled Trickling Filter/Activated-Sludge Process The use of coupled process designs likely precedes the 1950s, although definitive reports are lacking. The development of plastic media led to the use of trickling filters for highly loaded (roughing) modes of operation. These units were found to be suited particularly for strong municipal and industrial wastewater. Bryan and Moeller (1960) reported that the initial coupled process installation was designed to improve the quality of an activated sludge (AS) effluent during upset conditions. Soon after, the roughing trickling filter (RTF) found a better process location, which was before the activated sludge process. Here, the RTF smoothed the loading and stabilized the overall process. It also was used to protect activated sludge systems against toxic and shock loads from industrial contributors. Gehm and Gellman (1965) demonstrated that the coupled RTF/AS process could provide for control of bulking organisms and thus enhance the overall performance of WWTPs. The reduction of the sludge volume index (Gehm and Gellman, 1965) often was cited as the reason for using a trickling filter before an activated sludge system. Lower energy costs and the ability to handle shock loads also were cited as advantages of the coupled process. Many WWTPs receiving food, beverage, and other highcarbohydrate wastewater found it beneficial to provide partial treatment of the wastewater with trickling filters before the activated sludge process, to enhance the overall treatment performance. In the early 1970s, a modified trickling filter concept using wood slat media was marketed. This concept was called the activated biofilter (ABF) process. The first reference to this concept was reported by Egan and Sandlin (1960), where the settled sludge from a plastic media effluent was recycled over the tower. Bryan (1962) reported on the design of the Saginaw Bay and Cities Service facilities in Saginaw, Michigan, where provision was made to recirculate the return activated sludge through a media filter. In this process, the sloughed trickling filter biomass was recirculated at a high rate through the filter. The filter used an open media, and wood stacks were the most common media. This trickling filter concept was later coupled with a short-term (15 to 30 minutes) activated sludge system (ABF/AS) (Slechta and Owen, 1974), at Corvallis, Oregon, and found wide acceptance for treating municipal and industrial wastewater. Research in Corvallis, Oregon, during the late 1970s led to the development of the trickling filter/solids contact process in 1979 (Norris et al., 1982). This process uses a trickling filter sized for the removal of the majority of the soluble organic matter
Introduction
followed by an aerated solids contact channel or basin and a secondary clarifier with recycle of return activated sludge to the aerated solids channel. The primary purpose of the aerated solids contact channel is to flocculate suspended solids contained in trickling filter effluent for removal in the secondary clarifier.
3.5 Biological Filters The basic concept of the biological filter is to provide a high specific surface area for the growth of biomass in a compact volume. The biological filter provides for both the removal of substrates biologically and the retention of solids, which are removed periodically through a backwash cycle. The basic technology has been applied to aerobic, anoxic, and anaerobic processes. Biological aerated filters (BAFs) were developed in the 1980s in Europe and have found extensive application since then for carbon and nitrogen removal. Since 1982, over 500 installations have been built using a variety of technologies (Stephenson et al., 2004). A wide variety of configurations and media systems have been developed, but, generally, the advantages that have come to be recognized for this type of technology include the following: • Relatively small space requirement, because they are typically high-rate systems; • Ability to treat dilute wastewaters; • No concern for sludge settleability; and • Relatively easy to contain potential odors. The technologies may be categorized according to whether they are upflow or downflow processes; packed bed or fluidized bed; and aerobic, anoxic, or anaerobic.
3.6 Hybrid Processes The use of fixed-film media in aerated reactors is an old concept. In recent years, the concept has been referred to as integrated fixed-film activated sludge (IFAS) and has been used increasingly as a means of improving the treatment capacity and nutrient removal capabilities of existing activated sludge facilities. Before the 1940s, the Hays and Griffith processes (Wilford and Conlon, 1957) used baffles, cement asbestos, wood, or other construction materials in the aeration tank to enhance performance. The Hays and Griffith processes were improved over
7
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Biofilm Reactors
the original process (Hays, 1931), in which submerged rock filters were force-aerated. These processes were called submerged contact aeration processes. Approximately 60 Hays contact plants were constructed for United States Army camps between 1940 and 1945 (Packham, 1988). Wood, asbestos, and plastic sheeting were fixed in the aeration basin to provide a surface for bacterial growth. The plant treatment trains consisted of primary clarification followed by two equally sized aeration basins with intermediate and final clarification. Sludge return was not practiced. Asbestos sheets were suspended vertically in the aeration basins, such that they extended from 10 cm below the liquid surface to just above the horizontal grid of aeration pipes. The asbestos sheets were spaced 3.8 cm apart throughout the length of the aeration tanks and permitted a serpentine pathway for floating material. The aerators provided oxygen transfer and mixing between the asbestos sheets, and treatment was affected by the attached fixed-film growth on the sheets. Typical combined hydraulic retention times for the two aeration basins ranged from 1.7 to 3.0 hours. The National Research Council Report (1946) developed design criteria for the contact aeration process simultaneously with the development of design criteria for activated sludge and trickling filter processes. The contact aeration process eventually failed because of the additional cost of aeration and because it produced a lower quality effluent compared with the activated sludge process. The process essentially ceased to be used in the 1960s. Steels (1974) reported on previous attempts in the 1920s to improve plant operations by retaining various types of small particulates in the aeration basins. These particulates included grit, brushwood, wood, and cork and were effective in treating strong wastewater. The Hays and Griffith processes resurfaced in a number of variant forms, including a similar concept (Hamoda and Abd-El-Bary, 1987) referred to as an activated submerged fixed-growth bioreactor. The use of submerged RBCs in aeration basins is another form of this concept. The basic concept of these systems, as currently applied, is to increase the amount of biomass available for treatment per unit volume. By retaining the biomass on a fixed-film media within the suspended-growth reactor, treatment performance can be enhanced without increasing the solids loading on the secondary clarifiers. The IFAS systems, in their various forms, combine the apparent advantages of both fixedfilm and suspended-growth biomass systems. Another major effort to use fixed-film media in aeration tanks originated in Japan during the 1960s. Kato and Sekikawa (1967) developed a process that they called fixed
Introduction
activated sludge and applied primarily to the treatment of industrial wastes. This process vertically suspended an open plastic matrix in an aerated reactor and typically was operated without a return sludge. More than 60 such installations were in place in Japan during the 1960s. Both fixed and free-floating types of media systems have been developed. A rope-type media, which is installed on racks placed in the activated sludge basins, originally was developed in Japan for the purpose of achieving greater levels of carbonaceous biochemical oxygen demand removal within the volume of an existing activated sludge basin (Iwai et al., 1990). The product then was applied in Germany to upgrade treatment plants for nitrification and subsequently in North America for both purposes. The majority of applications have been in aerobic basins or in basins that alternate between aerobic and anoxic conditions. Pilot testing of the rope-type media began in the United States in the early 1990s, followed by full-scale installations (Randall and Sen, 1996). Several different types of rope media have been developed and are available commercially. Free-floating types of systems use the sponge of plastic media. Processes using sponge-type media originally were developed in Europe in the late 1970s, and two basic systems emerged. One system, the Captor process (Atkinson et al., 1979), was developed by Simon-Hartley through work completed at the University of Manchester in the United Kingdom. The concept was commercialized by Simon-Hartley in conjunction with several universities, Severn Trent Water Authority (Birmingham, United Kingdom), and the Water Research Center in Swindon, Wiltshire, United Kingdom. A second sponge-type system was developed by the LindeAG Corporation in the mid-1970s (Hegemann, 1984) and was introduced commercially in Europe and North America as the Linpor System by the Lotepro Corporation (Mount Kisco, New York) (Morper and Wildmoser, 1990; Reimann, 1990). A plastic media system using small cylindrical biofilm carrier elements originally was developed by Kaldnes Miljoteknologi in Norway (Odegaard and Rusten, 1990; Odegaard et al., 1994). There currently are several manufacturers of this type of media system. The free-floating media systems have been applied in both IFAS and movingbed biofilm reactor (MBBR) configurations. The MBBR process differs from IFAS in that there is no return activated sludge; thus, the MBBR is a pure fixed-film process and not a hybrid. The plastic media by Kaldnes Mitjoteknologi originally was developed as an MBBR process and was patented (AnoxKaldnes MBBR) (Odegaard, 2006).
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Biofilm Reactors
4.0 Organization of Manual A brief description of the organization of this manual and the material contained in each chapter is presented in this section. Chapter 2 presents the microbiology of wastewater treatment as applied to biofilms and is intended to provide a foundation for understanding the discussion of kinetics presented in the individual technology chapters and in the final chapter on modeling. Trickling filter processes and RBC processes are discussed in Chapters 3 and 4, respectively. Moving-bed biofilm reactors are discussed in Chapter 5. In each case, the process and mechanical design considerations and the operational issues are presented. Various design approaches are presented, including both empirical and kinetic-based models specific to the technology. Application of the technologies to carbon removal and nutrient removal are presented, and case studies are included. The focus of Chapter 6, Hybrid Processes, is on IFAS processes. The various types of media incorporated to IFAS systems are discussed with design considerations relevant to each type of media. Case studies are presented to illustrate their application. Chapter 7, Biological Filters, presents the design, construction, and operational issues associated with aerobic and anoxic biological filters in their various configurations, including downflow and upflow BAFs with sunken media, upflow with floating media, open-structure media filters, and continuous backwash filters. The application of biological filters for carbon removal, nitrification, and denitrification is discussed. New and emerging technologies that show promise in the laboratory or through small pilot-scale studies, but that are not yet commercially available, are discussed in Chapter 8. Clarification, as it relates to meeting advanced levels of treatment following fixed-film systems, is discussed in the following two chapters. Chapter 9 discusses gravity settling and membrane separation. Chapter 10 discusses effluent filtration technologies. Finally, a comprehensive discussion of modeling concepts is presented in Chapter 11. Each technology chapter presents modeling techniques specific to that process, but this chapter covers the kinetics in general, modeling techniques, and application of models to various types of pure fixed-film and hybrid systems.
Introduction
5.0 References Alleman, J.; Peters, R (1982) The History of Fixed Film Wastewater Treatment Systems. Proceedings of the International Conference of Fixed Film Biological Processes, Kings Island, Ohio. http://web.deu.edu.tr/atiksu/ana52/biofilm4. pdf (accessed March 2010). Atkinson, B.; Black, G. M.; Lewis, P. J. S.; Pinches, A. (1979) Biological Particles of Given Size, Shape and Density for Use in Biological Reactors. Biotechnol. Bioeng., 21 (2), 193–200. Bryan, E. H. (1982) Development of Synthetic Media for Biological Treatment of Municipal and Industrial Wastewater. Paper presented at the 1st International Conference on Fixed-Film Biological. Processes, Vol. 1, Kings Island, Ohio; Sponsored by University of Pittsburg, U.S. Army Corps of Engineers, U.S. Environmental Protection Agency, and U.S. National Science Foundation), 89. Bryan, E. H. (1955) Molded Polystyrene Media for Trickling Filters. Proceedings of the 10th Purdue Industrial Waste Conference, West Lafayette, Indiana, May 9–11; Purdue University: West Lafayette, Indiana, 164. Bryan, E. H. (1962) Two-Stage Biological Treatment: Industrial Experience. Proceedings of the 11th South Municipal Industrial Waste Conference; North Carolina State University, North Carolina. Bryan, E. H.; Moeller, D. H. (1960) Aerobic Biological Oxidation Using Dowpac. Proceedings of the Conference on Biological Waste Treatment; Manhattan College: Riverdale, New York. Caink, T. (1897) Specifications of Inventions. Br. Patent 19153. Candy, F. P. (1898) Specifications of Inventions. Br. Patent 2749. Cooper, P. F. (2001) Historical Aspects of Wastewater Treatment, Decentralized Sanitation and Reuse: Concepts, Systems and Implementation, Chapter 2; International Water Association Publishing: London, United Kingdom. Corbett, J. (1902) Some Sewage Purification Treatment Experiments. J. Sanit. Inst., 23, 601–602. Crimp, S. (1890) The Construction of Works for the Prevention of Pollution by Sewage of Rivers and Estuaries; Charles Griffin & Company: London, United Kingdom.
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Biofilm Reactors
Dibdin, W. J. (1903) The Purification of Sewage and Water, 3rd ed.; The Sanitary Publishing Company: London, United Kingdom. Dow Chemical Company (1955) Dowpac™ FN-90 and Dow HCS™. Plastics Technical Services., Dow Chemical Company: Midland, Michigan. Egan, J. T.; Sandlin, M. (1960) The Evaluation of Plastic Trickling Filter Media. Proceedings of the 15th Purdue Industrial Waste Conference, West Lafayette, Indiana; Purdue University: West Lafayette, Indiana, 107–-115. Gehm, H. W.; Gellman, I. (1965) Practice, Research and Development in Biological Oxidation of Pump and Paper Effluents. J. Water Pollut. Control Fed., 57, 1392–1398. Hamoda, M. F.; Abd-El-Bary, M. F. (1987) Operating Characteristics of the Aerated Submerged Fixed-Film (ASFF) Bioreactor. Water Res., 21, 939–947. Hays, C. C. (1931) Sewage Treatment Process. U.S. Patent 1,991,896. Hegemann, W. (1984) A Combination of the Activated Sludge Process with FixedFilm Bio-Mass to Increase the Capacity of Wastewater Treatment Plants. Water Sci. Technol., 16, 119–130. Iwai, S.; Oshino, Y.; Tsukada, T. (1990) Design Operation of Small Wastewater Treatment Plants by the Microbial Film Process. Water Sci. Technol., 22, 139–144. Kato, K.; Sekikawa, Y. (1967) FAS (Fixed Activated Sludge) Process for Industrial Waste Treatment. Proceedings of the 22nd Purdue Industrial Waste Conference, West Lafayette, Indiana; Purdue University: West Lafayette, Indiana, 926–949. Mills, H. F. (1890) Purification of Sewage and Water. Special report to the Massachusetts State Board of Health: Boston, Massachusetts, 25. Morper, M.; Wildmoser, A. (1990) Improvement of Existing Wastewater Treatment Plant Efficiencies Without Enlargement of Tankage by Application of the Linpor Process—Case Studies. Water Sci. Technol., 37, 207–215. Norris, D. P.; Parker, D. S.; Daniels, M. L.; Owens, E. L. (1982) High Quality Trickling Filter Effluent Without Tertiary Treatment. J. Water Pollut. Control Fed., 54, 1087–1098. Odegaard, H. (2006) Innovations in Wastewater Treatment: the Moving Bed Biofilm Process. Water Sci. Technol., 53, 17–33.
Introduction
Odegaard, H.; Rusten, B. (1990) Upgrading of Small Municipal Wastewater Treatment Plants with Heavy Dairy Loadings by Introduction of Aerated Submerged Biofilters. Water Sci. Technol., 22 (7/8), 191–198. Odegaard, H.; Rusten, B.; Westrum, T. (1994) A New Moving Bed Biofilm Reactor—Application and Results. Proceedings of the 2nd International Specialized Conference on Biofilm Reactors, Paris, France, Sep 29–Oct 1; International Association on Water Quality: London, United Kingdom, 221–229. Packham, R. F. (1988) Biological Filtration. Manuals of British Practice in Water ollution. Control, Institute of Water Pollution Control: London, United Kingdom. Peters, R. W.; Alleman, J. E. (1982) The History of Wastewater Treatment Systems. Paper presented at 1st International Conerence on Fixed-Film Biological Processes, Vol. 1, Kings Island, Ohio (sponsored by University of Pittsburg, U.S. Army Corps of Engineers, U.S. Environmental Protection Agency, and U.S. National Science Foundation), 60. Randall, C.; Sen, D. (1996) Full-Scale Evaluation of an Integrated Fixed-Film Activated Sludge (IFAS) Process for Enhanced Nitrogen Removal. Water Sci. Technol., 33 (12), 155–162. Reimann, H. (1990) The Linpor Process for Nitrification and Denitrification. Water Sci. Technol., 22, 297–298. Royal Commission on Sewage Disposal (1908) Fourth Report. Royal Commission on Sewage Disposal: London, United Kingdom. Slechta, A. E.; Owen, W. F. (1974) ABF Short-Term Aeration-Pilot Plant Results, Corvallis, OR. Technical Bulletin, Neptune, Microfloc, Inc.: Corvallis, Oregon. Stanbridge, H. H. (1972) The Introduction of Rotating and Traveling Distributors for Biological Filters. Water Pollut. Control, 44, 573. Steels, I. H. (1974) Design Basis for the Rotating Disc Process. Effluent Water Treat. J., 14 (9), 434–445. Stephenson, T.; Cornel, P.; Rogalla, F. (2004) Biological Aerated Filters (BAF) in Europe: 21 Years of Full Scale Experience. Proceedings of the 77th Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Oct 2–6; Water Environment Federation: Alexandria, Virginia. Wilford, J.; Conlon, T. P. (1957) Contact Aeration Sewage Treatment Plants in New Jersey. Sew. Ind. Wastes, 29, 845–855.
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6.0 SUGGESTED READINGS Sen, D.; Randall, C. W. (1996) Mathematical Model for a Multi-CSTR Integrated Fixed Film Activated Sludge (IFAS) System. Proceedings of the 69th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 5–9; Water Environment Federation: Alexandria, Virginia. Stensel, H. D.; Brenner, R.; Lee, K.; Melcer, H.; Rakness, K. (1988) Biological Aerated Filter Evaluation. J. Environ, Eng., 114 (6), 1352–1358.
Chapter 2
Biology of Fixed-Growth Process 1.0 INTRODUCTION
17
4.7 Flagella 4.8 Extracellular Polymeric Substances 4.9 Chemical Composition of Cells 4.10 Example 1—Theoretical Oxygen Demand of Bacterial Cells 4.10.1 Solution 4.10.2 Comment
2.0 CLASSIFICATION OF LIVING ORGANISMS 18 3.0 MICROORGANISMS OTHER THAN BACTERIA 21 3.1 Fungi 22 3.2 Algae 22 3.3 Protozoa 23 3.4 Multicellular Invertebrates 23 3.5 Viruses 24 3.6 Consortia 26 4.0 CHARACTERISTICS OF BACTERIA 4.1 Structure of the Bacterial Cell 4.2 Chromosome and Plasmids 4.3 Cytoplasm 4.4 Cell Membrane 4.5 Cell Wall 4.6 Pili
5.0 BACTERIAL METABOLISM, NUTRITION, AND RESPIRATION 5.1 Energy Source 5.2 Chemoheterotrophic Metabolism 5.3 Chemoautotrophic Metabolism 5.4 Photosynthetic Metabolism 5.5 Nutrient Requirements 5.6 Bacterial Energy Metabolism
26 26 28 30 30 30 31
31 31 32
32 32 33
33 36 36 38 39 40 41
(continued) 15
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Biofilm Reactors
5.7 Aerobic Growth and Respiration 5.8 Anoxic Conditions and Respiration 5.9 Anaerobic Respiration and Fermentative Metabolism 5.10 Energetics of Respiration 5.11 Example 2—Calculation of Electrode Reduction Potentials and ∆G for Half-Reactions 5.12 Solution 5.13 Co-Metabolism 6.0 BACTERIAL GROWTH 6.1 The Bacterial Growth Curve 6.2 Growth in Mixed Cultures 6.3 Enrichment Cultures 6.4 Stability of Mixed Cultures
42 43
44 46
48 48 50 51 51 53 53 54
6.5 Effects of Environmental Variables 54 7.0 BACTERIAL GROWTH KINETICS IN BIOFILMS 7.1 Rate of Bacterial Processes 7.2 Note 7.3 Physical and Chemical Changes in Biofilms Resulting from Growth
55 56 56
59
7.4 Structured Models 7.5 Temperature Effects 7.6 Example 3—Effect of Temperature on Organic Removal
59 60
7.7 Solution 7.8 Inhibition and Toxicity 7.9 Mass-Transfer-Rate Limitations
61
61
63 64
8.0 KEY TRANSFORMATIONS IN BIOFILMS 65 8.1 Chemoheterotrophic Processes 66 8.2 Chemoautotrophic Processes 67 8.3 Biology of Nitrogen Transformations 67 8.4 Denitrification 67 8.5 Aerobic Nitrification 68 8.6 Anoxic Nitrification/ Denitrification 69 8.7 Biological Phosphorus Removal 70 8.8 Sulfide and Sulfur Oxidation 71 8.9 Hydrogen Oxidation 71 9.0 FEATURES OF MICROBIAL COMMUNITIES IN BIOFILMS 72 10.0 REFERENCES
74
Biology of Fixed-Growth Process
1.0 INTRODUCTION The purpose of this chapter is to provide a general introduction to the biology of fixedgrowth processes used in water and wastewater treatment. Although the fundamental principles of biological processes are the same for all process configurations, the constraints of fixed-growth processes result in particular characteristics of microbial communities and advantages and disadvantages relative to alternative systems. In fixed-growth processes, fluid containing necessary nutrients passes over the microorganisms growing in a biofilm on a support surface. Nutrients diffuse into the biofilm and are metabolized by the immobilized microorganisms. A significant result is that microbial populations can function well, even when fluid-phase nutrient concentrations are extremely low (Schroeder, 2002; Schroeder et al., 2000). Correspondingly, at high fluid-phase nutrient concentrations, oxygen transfer may become limiting, and anoxic zones may develop within the biofilm. The most basic considerations in microbial treatment process design and operation are that microorganisms use contaminants in water or wastewater as sources of carbon and energy for growth, as nutrients required to sustain growth, or as electron acceptors in the respiratory process. In all cases, microorganisms remove contaminants from the dissolved state; that is, particulate material must first be dissolved and then metabolized. Engineering of microbial systems requires consideration of the requirements for microbial growth and determination of the conditions necessary to carry out the desired biochemical reactions. Particular constraints encountered in biological wastewater treatment include the necessity to remove a wide range of organic and inorganic materials, consistent production of finished waters having extraordinarily low organic and suspended solids concentrations, and the presence of organic materials that are highly toxic to most forms of life. The environmental requirements for biological growth generally are inconsistent with the production of potable waters. If biological treatment is a step in the production of potable waters, a number of abiotic treatment steps must be incorporated downstream. In a sense, most rivers form minimally controlled treatment links between wastewater and potable water treatment systems. Treated wastewater discharged to rivers is diluted and contaminants are further stabilized before extraction and treatment of water for use by municipalities.
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2.0 CLASSIFICATION OF LIVING ORGANISMS Living organisms currently are classified into three general domains—Eucarya, Bacteria, and Archaea (see Table 2.1). Viruses—nonliving obligate infectious particles—are significant factors in microbial processes and also are discussed in this chapter. Two types of cell organization exist—eukaryotic and prokaryotic. Plants; animals; and the algae, protozoa, and fungi have eukaryotic cells, while members of the Bacteria and Archaea domains have the structurally simpler prokaryotic cell type. Viruses do not have a cellular structure, being composed of one or two strands of genetic material (deoxyribonucleic acid or ribonucleic acid) and a protein coat or
Table 2.1 Subdivisions of organisms and principal groups found in biological treatment systems. Domain or group
Cell structure Properties
Constituent groups
Eukaryotes
Eukaryotic
Nematodes, Plants (seed plants, rotifers ferns, mosses), algae, fungi Animals (invertebrates, vertebrates)
Multicellular, extensive differentiation of cells and tissues
Groups commonly found in biological treatment
Algae, fungi, protozoa Unicellular, mycelial or filamentous: little or no differentiation of cell type
All groups
Bacteria
Prokaryotic
Cell chemistry similar to Eukaryotes
Most bacteria
Most Gram-positive genera, some Gram-negative genera
Archaea
Prokaryotic
Distinctive cell chemistry
Methanogens, halophiles, thermoacidophiles
Methanogens
Viruses
None
Non-living, obligate parasites composed of nucleic acid strand(s) and protein coat or capsid
Biology of Fixed-Growth Process
capsid. Eukaryotes are divided into two general groups—multicellular organisms, in which cells have specific functions, and unicellular organisms, in which all cells carry out the same range of functions. Multicellular organisms are classified into two general categories—plants and animals—according to their energy and carbon sources, structure, type of growth, and movement. Unicellular organisms are classified into three general categories—protozoa, fungi, and algae—in a manner analogous to that for multicellular eukaryotic organisms. The two domains characterized by the prokaryotic cell have distinctive differences in cellular chemistry. All of the organisms in these two divisions are microscopic, and most of the individual cells are smaller than 5 µm. Bacteria are divided into 17 major lineages (or phylums) of cultured organisms and are estimated to contain many more phylums if uncultured microbial communities are included. Classification is based on the similarity of ribosomal ribonucleic acid (RNA) sequences that have been retrieved from various environments. As will be explained later, RNA comprises approximately 20% of the dry weight of cells. The most commonly used method of categorizing bacteria by genus and species is to establish similarities in a specific type of RNA, 16S ribosomal RNA, which has a molecular weight of approximately 500 000. Bacteria include organisms of relevance in biological treatment of water and wastewater, most of the organisms of importance in soils, and most of the organisms pathogenic to humans and other mammals. A relatively small fraction of known bacterial species are agents of disease (Schroeder and Wuertz, 2003), and pathogens may be harbored within biofilms in drinking water distribution systems (Szewzyk et al., 2000) and wastewater treatment systems (Skraber et al., 2007). Three phylums have been classified as Archaea. Euryarchaeota include Extremely Halophilic Archaea in environments having salinities of 1 to 5.5 M; Methanogens, which produce methane (CH4); Thermoplasmatales, which include thermophilic and extremely acidophilic genera that require low-pH (0.5 to 4.5) and high-temperature (55 to 85°C) environments; and Hyperthermophilic Euryarchaeota, which thrive at temperatures ranging from 80 to 110 °C. Crenarchaeota, the second phylum, contains members capable of existing at either very hot or very cold temperatures. A presumptive third phylum is Korachaeota; it consists of members found at high-temperature hydrothermal environments. In most cases, organisms are identified by their genus and species, with the genus capitalized and the species non-capitalized, and both genus and species in italics or underlined. For example, organisms of the genus Pseudomonas are extremely common in both soil and biological treatment processes. Species of Pseudomonas frequently
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Biofilm Reactors
found are P. aeruginosa, P. cepacia, P. putida, and P. stutzeri. Within a species, strains often are identified that behave slightly differently. For example, some strains of Pseudomonas putida may metabolize toluene, and others may not (metabolism is the sum of the processes in a living organism; materials that are metabolized are used to fuel cellular activities and incorporate to cellular molecular components). The differences between strains might be considered analogous to characteristics in humans such as height, left- or right-handedness, or ability to distinguish color. Microorganisms are defined loosely as organisms that cannot be seen by the naked eye, a definition that includes members of all of the divisions of organisms, as shown in Table 2.1. Biological treatment involves the use of microbial communities to remove or modify contaminants in water and wastewater. In most cases, the microbial communities are dominated by mixed (heterogeneous) bacterial populations (cultures made up of a number of bacterial species). A few species generally may make up the bulk of the bacterial population on either a number or mass basis. The reason that predominant species develop is the existence of a competitive advantage over less numerous groups, as a result of the particular environmental conditions or the available nutrient sources. For example, a biological process designed to remove a chlorinated aromatic compound and operating at 10 °C would almost always have at least one strain of the genus Pseudomonas among the dominant species, because this genus of bacteria is competitive at relatively low temperatures, and some species have a particular capacity to degrade aromatic compounds (Evans and Fuchs, 1988; Levin and Shapiro, 1967; Reineke and Knackmuss, 1988). Mixed bacterial populations also should be thought of in terms of interacting communities and symbiotic relationships. Overall growth and contaminant removal rates are enhanced by the interaction of the various species making up the population. In addition, interactions can be exploited further by controlling specific operational parameters in reactors. For example, some genera responsible for biological phosphorus removal require low-molecular-weight (volatile) organic acids that would not occur in significant concentrations in most wastewaters. Operating biological processes with an anaerobic segment results in fermentation (production) of volatile acids by acid-fermenting bacteria. In this manner, the substrates required by the phosphorus-removing bacteria are produced (substrate is a general term for microbial energy and/or carbon sources; the term comes from cultures grown on the surface of solid growth media, because nutrients are extracted from the substratum) (Fuhs and Chen, 1975; Levin and Shapiro, 1967, Shapiro et al., 1967). In a mixed-culture system that has not been engineered to remove nutrients, it would be expected that
Biology of Fixed-Growth Process
the various species in a mixed bacterial population produce necessary growth factors for other species as a matter of course, but this might not result in the desired reactions (e.g., phosphorus removal) taking place. In such mixed bacterial populations, species that grow most rapidly are better adapted to the particular environment (e.g., temperature, pH, or salinity) and are most efficient in energy use predominate. However, a type of hierarchy develops in which species depend on each other and in which species occupy ecological niches. For example, in a system treating gasoline production wastewaters, species that break down aromatic molecules may leave residues that serve as carbon and energy sources for species unable to break the aromatic ring. The predominant use of biological treatment is for the removal of organic compounds from water or wastewater, but removal of inorganic ions, such as ammonium, nitrate, nitrite, and phosphate, or even materials that are present in trace concentrations, such as selenium, arsenic, and mercury, is both possible and increasingly common. In addition, mixed microbial cultures can be used to produce desired end products, such as CH4. Descriptions of the structure and growth characteristics of bacterial cells and the other organisms found in biological treatment processes can be found elsewhere (Madigan et al., 2003).
3.0 MICROORGANISMS OTHER THAN BACTERIA Both microscopic multicellular and unicellular organisms have roles in fixed-growth biological treatment processes. Microscopic plants serve as surfaces on which smaller organisms grow and, through photosynthesis, as sources of oxygen. Microscopic animals serve as scavengers that remove floating debris. A few biological treatment processes have been designed that take advantage of plant growth characteristics, but the role of microscopic animals is largely uncontrolled. Unicellular eukaryotic organisms are present in most biological treatment processes. Protozoans serve a particulate scavenging role similar to multicellular animals and also may have a role in making particulate material more bioavailable to bacteria. Fungi have roles similar to the bacteria, but rarely compete well in treatment systems. Algae, like plants, can be used to provide oxygen to microbial systems and to remove inorganic nutrients, although most fixed-growth processes do not have enough light to maintain significant photosynthesis. A few biological treatment processes that take advantage of the potential contributions of algae and fungal treatment processes are under development for use with some hazardous materials (Woertz, Kinney, McIntosh, and Szaniszlo, 2001; Woertz, Kinney, and Szaniszlo, 2001).
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Figure 2.1 Coenocytic cell structure typical of many fungi, in which cell contents, or cytoplasm, is multinucleate and continuous throughout the organism.
3.1 Fungi Fungi have cell walls, typically are coenocytic (multinucleate with no cellular subunits [see Figure 2.1]), are immotile, and use organic material for both energy and carbon sources. Inorganic contaminants, such as nitrogen, phosphorus, and other nutrients, are incorporated to cell tissue in stoichiometric amounts, as in the case of bacterial growth. Fungi grow at considerably lower rates than bacteria and do not compete well in most treatment process environments. A fungus that appears to have considerable potential in the treatment of hazardous organic compounds is Phanerochaete chrysporium, a white rot fungus. This organism produces an extracellular peroxidase that degrades lignin in the presence of peroxide. The reaction is relatively nonspecific, with respect to the type of carbon–carbon bond attacked, and has been found to be effective in initiating the degradation of a variety of highly chlorinated compounds, including the dioxins (Aust et al., 1988; Hackett et al., 1977). The use of Phanerochaete chrysporium is limited to conditions in which nitrogen is limiting, because the peroxidase is not produced otherwise. A possible application of genetic engineering would be in the transfer of the gene responsible for the synthesis of the peroxidase to bacteria that grow well in wastewater treatment processes.
3.2 Algae Algae, like the fungi, are immotile eukaryotes with cell walls. Most species are unicellular, and some form aggregates of cells, including filaments. Not all algae are microscopic. Carbon dioxide (CO2) serves as the carbon source for algal growth, and energy is derived from the absorption of light by photosynthetic pigments, with oxygen being produced as a byproduct. Because the algae are aerobic organisms, a portion of the oxygen produced is used in their metabolism. Excess oxygen accumulates in the surrounding water, as long as enough light is available. The principal wavelength range in which algae absorb light is between 300 and 700 nm.
Biology of Fixed-Growth Process
Although algae are sometimes used in nutrient removal systems, they are extremely difficult to separate from water and often are troublesome contaminants themselves. Nutrient-rich ponds, sloughs, and lakes often have algal “blooms” over short periods of time, which result in unaesthetic green mats on the water surface. The mats may be blown onto beaches, where they decay, or are problems for people using the waters for recreation. Some algal species produce organic compounds that cause taste and odors in drinking water supplies. Certain marine flagellated algae (called dinoflagellates) produce toxins that are harmful to humans and fish.
3.3 Protozoa The protozoa do not contain cell walls and use bacteria, and presumably other organisms, and particulate and soluble organic materials for food. Thus, protozoa are heterotrophic predators and, in a sense, are higher on the trophic ladder than bacteria or algae. A large number of protozoan species exist, and a number typically are seen in the microbial communities found in fixed-growth processes. Protozoa ingest some organic contaminants and use inorganic ions also. However, the numbers of protozoa present are relatively small, their growth rates are substantially lower than those of bacteria, and the amount of contaminants removed by protozoa is negligible relative to that removed by bacteria, fungi, and algae. However, they may serve to pre-process particulates for bacteria, by producing exudates of increased solubility.
3.4 Multicellular Invertebrates Higher forms of microorganisms—those with complex, invertebrate, multicellular bodies, such as rotifers and nematodes—generally act as a higher step in the food chain, but overall perform a role that is similar to the protozoa. It is quite possible for a fixed-growth process to operate successfully without the presence of protozoa or higher forms, but the presence of scavengers generally is believed to be beneficial. These organisms are present because the environmental conditions necessary for satisfactory contaminant removals often are suitable for protozoa and higher forms. Their absence in fixed-film processes often is associated with sloughing events—that is, the uncontrollable detachment of whole segments of biofilm from the support medium. The upper limit on the size of the higher forms (e.g., nematodes rather than carp) is the result of several factors, including sensitivity to oxygen concentration, physical space requirements, and growth/replication rates.
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3.5 Viruses A group that stands alone as being nonliving, but having an effect on the growth of organisms, is the viruses. The viruses are particles composed of one or two strands of genetic material and a protective protein coat, or capsid. All viruses are obligate parasites; they are unable to carry out any form of metabolism and are completely dependent on host cells for replication. The mechanism of viral action is composed of the following four steps:
(1) Adsorption of the viral particle (viron) onto a receptor site on the cell membrane; (2) Infection—the insertion of the genetic material to the host cell; (3) Redirection of host cell functions to replication of viral particles; and (4) Release of the viral particles into the environment (often by disruption of the cell membrane, with consequent death of the cell).
Viral particles typically are 0.01 to 0.2 µm in size and are species-specific, with respect to infection. Most viral groups attack only one host species, and individuals within a species may become resistant. Viruses that attack bacteria are termed bacteriophage (or simply phage). A number of important human diseases are caused by viruses, and examples are listed in Table 2.2. Modes of transmission of hepatitis B viruses, polioviruses, and picornaviruses are known to include water, although the principal pathways typically are through food, personal contact, or exchange of body fluids. Some viruses, such as the retroviruses (including the HIV group), appear to be too fragile for water transmission to be a significant danger to public health. Viruses are important in water and wastewater treatment as contaminants to be removed. Discharge of human viruses into drinking water supplies is completely unacceptable, and the discharge of significant concentrations of human viruses with treated wastewaters is a matter of increasing concern. The role of bacteriophages in biological treatment processes is not currently understood, although phages are known to influence bacterial population dynamics in marine systems (Steward et al., 1996). Methods for monitoring the presence of human viruses are not commonly applied because of the small size of the particles and the difficulty in culturing. Some important environmental viruses cannot be grown by cell culture, at present. Increasingly, molecular biology methods based on the amplification of deoxyribonucleic acid (DNA) and, indirectly, RNA using the polymerase chain reaction (PCR) are applied to water samples (Rajal, McSwain, Thompson, Leutenegger, Kildare, and
Biology of Fixed-Growth Process
Table 2.2 Common human diseases caused by viruses (adapted from Stanier et al., 1986). Viral group and type
Disease
Herpesviruses Cytomeglovirus
Respiratory infections
Epstein–Barr virus
Mononucleosis
Herpes simplex viruses
Oral and genital cold sores
Varicella virus
Chickenpox, shingles
Hepatitus B virus
Serum hepatitus
Influenza viruses
Viral influenza and viral pneumonia
Polioviruses
Poliomyelitis
Poxviruses Orf virus
Contagious pustular dermatitis
Variola virus
Smallpox
Picornaviruses Coxsackie viruses
Herpangina
Hepatitis A virus
Infectious hepatitis
Poliomyelitis virus
Poliomyelitis
Rhinoviruses
Most colds
Parainfluenza viruses
Measles, mumps, rubella
Rhabdoviruses
Rabies
Reoviruses
Diarrheal diseases
Retroviruses Human T-cell leukemia virus
T-cell leukemia
Human immunodeficiency viruses
Acquired immunodeficiency syndrome (AIDS)
Rotaviruses
Diarrhea
Wuertz, 2007). This research now allows a more rapid and quantitative detection of viruses. A differentiation between viable or infectious and inactivated viruses by quantitative PCR is possible, in principle, but is difficult to perform at this time. At present, disinfection methods are used that have been demonstrated to be effective in
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Biofilm Reactors
laboratory and field experiments. Regular monitoring of process effluents for viruses is not a conventional practice, but may become so once molecular detection methods have been validated sufficiently.
3.6 Consortia It is important to see the microorganisms used in biological treatment as a community living in consortia. Each of the organisms present occupies an ecological niche, and the relative numbers of the species will change as conditions change. For example, increases or decreases in temperature will modify the competitive advantages among the species present and result in changes in species predominance. It must be assumed that cyclic operations are more favorable to some groups than others. The ability to use particular materials, such as benzene rings or ammonia, as energy sources is not widespread among microbial species, and the presence of these materials may give some species either a niche, allowing them to survive, or a competitive advantage over other species.
4.0 CHARACTERISTICS OF BACTERIA Because bacteria are the dominant organisms in biological treatment systems, they will be discussed here in greater detail. Additional information on the other organisms commonly present in fixed-growth processes will be given as appropriate. General structural characteristics of the bacterial cell are presented here, and more detailed information can be found in other references (Ingraham et al., 1983; Madigan et al., 2003; Stanier et al., 1986).
4.1 Structure of the Bacterial Cell An overall understanding of the typical composition of bacterial cells can be obtained from Tables 2.3 and 2.4. Although the composition given is for a particular bacterial species, Escherichia coli, grown under specific conditions, the relative makeup of most bacteria is similar. The physical structure of the bacterial cell can be characterized by shape (i.e., spherical, rod, and spiral) and components, including their chemical makeup, size, and the manner in which they grow (i.e., individual cells, colonies, and filaments). Shape and size are somewhat variable with stage of growth. A species listed as a rod may appear to be spherical under some growth conditions, and cell size changes during the growth cycle. Bacteria that typically are found in groups, such as the filamentous species (which often are characterized by coenocytic cell structure),
Biology of Fixed-Growth Process
27
Table 2.3 Typical composition of Escherichia coli (adapted from Ingraham et al., 1983). Macromolecule
Percentage of Mass/cell* Molecular Number of Different kinds total dry weight grams (×1015) mass molecules per cell of molecules
Protein
55.0
155.0
RNA (ribonucleic acid)
20.5
59.0
4.0 x 104
2 360 000
1050
23 S rRNA
(31.0)
1.0 x 106
18 700
1
16 S rRNA
(16.0)
5.0 x 10
18 700
1
5 S rRNA
(1.0)
3.9 x 10
18 700
1
transfer
(8.6)
2.5 x 10
205 000
60
messenger
(2.4)
1.0 x 10
1380
400
5 4 4 6
2.13
1
DNA (deoxyribonucleic acid)
3.1
9.0
2.5 x 10
Lipid
9.1
26.0
705
22 000 000
4
Lipopolysaccharide
3.4
10.0
4346
1 200 000
1
Peptidoglycan
2.5
7.0
(904)n
1
1
Glycogen
2.5
7.0
1.0 x 10
4360
1
96.1
273.0
Soluble pool
2.9
8.0
Inorganic ions
1.0
3.0
Total macromolecules
Total dry mass/cell
284.0
Water
670.0
Total mass/cell
9
6
954.0
*Note that grams times 10 means that the actual mass has been multiplied by 10n. For example, the protein mass of an E. coli cell is approximately 155 x 10-15 g. n
grow as single cells also. Thus, visual observation does not provide a method of species identification. Current molecular methods of identification and classification are based primarily on similarity of 16S ribosomal RNA or detection of specific genes. For example, intact cells belonging to the Bacteria and Archaea domains can be tested with fluorescently labeled genetic probes (short sequences of DNA that can penetrate most microbial cells). The probes are single-stranded and bind to 16S rRNA in the cell. Only perfect matches of sequences will lead to retention of the probe in the cell after a sequence of hybridization and washing steps. With the aid of a good fluorescent
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Table 2.4 Typical elemental composition of bacterial cells (Stanier et al., 1986). Element Carbon Oxygen Nitrogen Hydrogen Phosphorus Sulfur Potassium Sodium Calcium Magnesium Chlorine Iron ΣTrace elements
Percent of dry weight General physiological function 50 20 14 8 3 1 1 1 0.5 0.5 0.5 0.2 0.3
Constituent of organic cell materials Constituent of organic cell materials and cellular water Constituent of proteins, nucleic acids, coenzymes Constituent of cellular water and organic cell materials Constituent of nucleic acids, phospholipids, coenzymes Constituent of proteins and coenzymes Major cation in cell processes Major cation in cell processes Major cation in cell processes and enzyme cofactor Major cation in cell processes, cofactor in ATP reactions Major anion in cell processes Constituent of cytochromes and other proteins, enzyme cofactor Inorganic contituents of special enzymes
microscope, it is possible to verify in situ the presence or absence of certain groups of prokaryotes in a biofilm (Wilderer et al, 2002). Other methods include immunoassays and fatty acid methyl ester analysis and phospholipid fatty acid analysis, which are used to develop unique “fingerprints” of species and consortia, bacterial metabolism, resistance to antibiotics, and chemical characteristics of the cell wall. The nature of the structural components of the cell is of considerable interest in biological treatment. Reference can be made to Figure 2.2 in discussing the most important components relative to contaminant removal—the genetic component, enzymes, storage bodies, cytoplasmic (cell) membrane, cell wall, and bound extracellular polymeric substances (EPSs, often referred to as capsule). Bound EPS is associated closely with cells, but its polymers are not anchored in the cell envelope (Nielsen and Jahn, 1999). Strictly speaking, EPSs are not part of the cell structure but are included here because of their importance in biofilms and the difficulty of physically separating bound EPSs from cells (Spaeth and Wuertz, 2000).
4.2 Chromosome and Plasmids The genetic component of bacterial cells (the genome) includes both the single DNA molecule, called chromosome, and relatively small DNA circlets, called plasmids, located in the cytoplasm. The chromosome is circular and double-stranded, with a length greater than 1000 µm and a molecular weight of approximately 109. Because bacterial
Plasmid
Ce ll Membrane
Cell Wall Nucleoid DNA
Capsule Pilus
Biology of Fixed-Growth Process
Flagellum
Cytoplasm Includes: RNA, protein, plasmids (DNA), volutin granules (polyphosphates, sulfur), storage products (glycogen, lipids)
Figure 2.2 Schematic representation of the bacterial cell. Note that particular bacterial species are characterized by general shape and the tendency to occur singly, in clumps, or as chains or filaments. Typical cell sizes, excluding the capsule, range from 0.5 to 2 µm. The capsule is considered part of the EPS, and its size tends to be very small during rapid growth.
cells are approximately 1 to 2 µm in length or diameter, and the nuclear region is only a small part of the cell, the chromosome must be folded tightly. The chromosomal DNA is essential to the life of a cell. Without it, the information required for the production of necessary enzymes and other structures necessary for growth is missing. Damage to the DNA can result in the loss of an essential activity and death. Plasmids consist of extrachromosomal, typically circular, double-stranded DNA and typically are not a required component of the cell, but they seem to provide particular capabilities that make a strain more or less competitive in a given environment. These include resistance to specific antibiotics, toxin production, ability to metabolize unusual compounds or ions, and possibly resistance to attack by bacteriophage.
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4.3 Cytoplasm The tightly packed, granular region within the cytoplasmic membrane is the cytoplasm. Principal components of the cytoplasm are the ribosomes, made up of RNA and proteins, where other cell components are synthesized; enzymes (proteins), which catalyze necessary chemical reactions; plasmids, organic and inorganic compounds involved in metabolism; and granular inclusions containing storage materials, such as glycogen and poly-β-hydroxybuterate (PHB), polyphosphate (volutin granules), and sulfur. Carbon storage as glycogen or PHB appears to be speciesdependent. Enteric bacteria, cyanobacteria (blue-green algae), and spore-forming bacteria tend to store the glucose polymer, glycogen, while groups commonly found in soil, such as pseudomonads and rhizobia, tend to store PHB. Some groups (e.g., Acinetobacter and possibly the filamentous bacteria) do not store carbon, which is a factor in their competitiveness under feast/famine conditions. Glycogen granules are detectable only using electron microscopy (Ingraham et al., 1983) and are relatively evenly distributed. Poly-β-hydroxybuterate, polyphosphate (or volutin), and sulfur granules often can be seen, with special stains applied to slides, as refractile bodies under conventional light or fluorescent microscopy (Serafim et al., 2002). The intracellular accumulation of large amounts of polyphosphate granules in some bacterial species under cyclic anaerobic or anoxic/aerobic operation is used in enhanced biological phosphorus removal processes (Kong et al., 2005).
4.4 Cell Membrane Surrounding the cytoplasm is a bilayer unit membrane composed primarily of phospholipids and proteins. The cell membrane acts as an osmotic barrier (phospholipid function) and has specific transport functions (protein function), in which transferase enzymes carry out transport reactions for specific molecules. In aerobic bacterial cells—those that use oxygen as a terminal electron acceptor—the electron transport system of enzymes, in which the energy-rich compound ATP is produced and oxygen is consumed, is attached to the cell membrane. The typical thickness of the cell membrane is approximately 7 to 8 nm.
4.5 Cell Wall Structural strength is provided by the cell wall. However, this structure also acts as a molecular sieve, which screens out toxic molecules and antibiotics, and contains binding and hydrolytic enzymes, which aid in nutrient gathering and transport.
Biology of Fixed-Growth Process
In bacteria, two general types of cell envelopes are identified by the Gram-stain process, in which an applied dye is found to be permanently fixed (positive) or washed out of the cell (negative). Both types have cell walls made of peptidoglycan. Gram-positive bacteria have a thicker cell wall, and many species also have acidic polysaccharides, called teichoic acids. They often are resistant to desiccation and are found more commonly in soil and other oligotrophic environments. Many bacteria found in wastewater reactors are Gram-negative. Their cell envelope includes a thinner cell wall made of peptidoglycan and a second (outer) membrane, in addition to the cell membrane. Archaea have cell walls made up of a variety of components, but not peptidoglycan.
4.6 Pili Single strands of protein originating in the cell membrane and extending approximately 10 µm from the cell wall are the pili. These hairlike apparati appear to have a function of binding the cell to specific structures. For example, pili are required to establish cell-to-cell contact in the transfer of plasmids from one cell to another, in the process of bacterial conjugation. The presence of pili on the cell surface can aid the initial cell adhesion to interfaces. A specific type of pilus, type IV, also has been shown to be involved in cell movement (Mattick, 2002) and movement of biofilms across solid surfaces by retracting the pilus and pulling the cell forward (HallStoodley et al., 2004).
4.7 Flagella The long (15 to 20 µm) filament called the flagella moves the bacterial cell with a whiplike rotational motion. Movement of motile bacteria (not all species have such structures) toward favorable environments is a complex response to chemical gradients and “attractant solutes” in the cell’s surroundings.
4.8 Extracellular Polymeric Substances A polysaccharide layer is secreted by many bacteria. The size of this amorphous capsule or glycocalyx may be larger than the cell, in some cases. In addition, there are many other macromolecules that are either actively secreted or released through cell leakage and cell death. They include proteins and nucleic acids (DNA and RNA). The EPS can act as a binding agent, attaching cells onto surfaces, such as fixed-growth packing media, rocks, pipes, and teeth, or to other cells to form microbial aggregates
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in the form of biofilms and flocs. Initial adhesion to interfaces also is facilitated by the cell appendages, such as the pili discussed above. Once a biofilm is developing, the cellular EPS production is increased (Venugopalan et al., 2005). An operational definition of EPS distinguishes between bound and soluble EPS (Nielsen and Jahn, 1999).
4.9 Chemical Composition of Cells In most engineering problems, the fundamental composition and structure of microbial cells is of less importance than overall size (typically 1 to 5 µm for bacterial cells) and elemental chemical composition. The principal elements of all living cells, on a mass basis, are carbon, oxygen, nitrogen, hydrogen, phosphorus, and sulfur. A number of metals (iron, manganese, potassium, cobalt, calcium, copper, and zinc) are essential to life, because they serve as cofactors (or mediators) to the electron transport that takes place in specific enyzyme-catalyzed reactions. Empirical cell formulas, or molar ratios, of the chemical constituents of microbial cells are used to estimate nutrient requirements for growth and convert gravimetric cell mass measurements into theoretical oxygen demand of cell tissue. The most widely used empirical cell formula, C5H7NO2 (Porges et al., 1953, Rittman and McCarty, 2001), omits the essential nutrient phosphorus. Inclusion of phosphorus results in considerably more complex formulas, such as C42H100N11O13P (McCarty, 1965). Empirical cell formulas are representative of growth under specific environmental conditions, and care must be used in general application of the relationships, as can be seen in the following example.
4.10 Example 1—Theoretical Oxygen Demand of Bacterial Cells Determine the theoretical oxygen demand of 1 g (dry) of microbial cells using the two empirical cell formulas given above. Assume that organic nitrogen in the cells is not oxidized and remains in the -3 oxidation state.
4.10.1 Solution The solution can be obtained as follows:
(1) Write a stoichiometric equation for the oxidation of C5H7NO2. C5H7NO2 + 5O2 → 5CO2 + NH3 + 2H2O (2) Determine the ThOD of 1 g of C5H7NO2 Empirical molecular weight = 113
Biology of Fixed-Growth Process
1 g = 0.008 85 mol 5 mol of O2 is required per empirical mole of cells (0.008 85 mol cells)(5 moles O2/mol cells)(32 g O2/mol O2) = 1.42 g
(3) Write a stoichiometric equation for the oxidation of C42H100N11O13P
2C42H100N11O13P + 107 O2 = 84CO2 + 22NH3 + 64H2O + 6H+ + 2PO4– 3 Empirical molecular weight = 997 1 g = 0.001 00 mol 53.5 mol of O2 are required per empirical mole of cells (0.001 mol cells)(53.5 mol O2/mol cells)(32 g O2/mol O2) = 1.71 g
4.10.2 Comment Note the large difference in oxygen demand that would be predicted for the same mass of cells from application of the two empirical formulas. The empirical formula used most commonly, C5H7NO2, was developed based on cells grown on a case in medium, and the general validity is not well-established.
5.0 B ACTERIAL METABOLISM, NUTRITION, AND RESPIRATION Microbial metabolism of interest in water and wastewater treatment is divided into the following two major categories: • Catabolism—the breakdown or degradation of materials with the concomitant release of energy, and • Anabolism—the synthesis of new cellular materials using the energy released in catabolism and the breakdown products of catabolism. Metabolism also can be described by the type of energy source used (e.g., organic compounds, inorganic ions, and light). Emphasis in this section will be placed on catabolism, while anabolism will be discussed in the Viruses and Consortia sections, dealing specifically with microbial growth. Nutrition, as defined here, is related to the chemical requirements associated with microbial metabolism. Most of the bacteria found in soils and natural waters and those used in wastewater treatment are not very fastidious, with respect to nutrition. Inorganic nutrient sources, such as ammonia and nitrate for nitrogen or phosphate for phosphorus, are satisfactory, and a relatively
33
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wide variety of organic materials can be used as energy sources. A few microorganisms found in natural systems, such as the nitrifiers and the methanogenic bacteria, have tightly constrained nutrient requirements. Microbial respiration is defined by the type of terminal electron acceptor used in the energy-yielding reaction processes and can be classified generally as aerobic, anoxic, or anaerobic and by the electron acceptor used (aerobic, anoxic, or anaerobic). A summary of bacterial metabolism and respiration is given in Tables 2.5 and 2.6, and more detailed discussions can be found in Lengeler et al. (Eds., 1999) and Madigan et al. (2003).
Table 2.5 Classification of living organisms by carbon and energy sources. Classification
Carbon source
Energy source
Autotroph
CO2
—
Heterotroph
Organic compounds
—
Chemotroph
—
Chemical compounds
Chemolithotroph
—
Inorganic compounds
Chemoorganotroph
—
Organic compounds
Phototroph
—
Light
Chemoautotroph
CO2
Chemical compounds
Photoautotroph
CO2
Light
Chemoheterotroph
Organic compounds
Chemical compounds
Photoheterotroph
Organic compounds
Light
Methylotroph
1-carbon compounds
1-carbon compounds
General
Other terms Eutroph
Use high concentrations for carbon and energy
Oligotroph
Use low concentrations for carbon and energy
Zymogenous
Grow rapidly when carbon and energy source added
Saprophyte
Live off of dead organic matter
Table 2.6 Bacterial metabolism and respiration. Electron acceptor
Principal non-cell product
Microbial group
O2
CO2, H2O
NO3–, NO2– SO4–2 Fe3+ , ClO4–
CO2, , N2, N2O CO2, H2S, S0 CO2, Fe+2 CO2, Cl-
CO2 Organic metabolites
CO2, CH4 Volatile acids, alcohols
Aerobic heterotrophs Denitrifiers Sulfate reducers Iron reducers Perchlorate reducers Methanogens Fermenters
O2
CO2, NO2
NH3
NO2–
N2, H2O, NO3–
NO2– H2 H2S H2 H2 Light (≈870 nm)*, H2S
O2 NO3–, NO2– O2 CO2 SO4–2 CO2
CO2, NO3– CO2, , N2, N2O CO2, SO3–2 CH4 H2S, H2O
Type of metabolism Electron donor Chemoheterotrophic
35
Chemoautotrophic
–
Photoheterotrophic Photoautotrophic *Energy source
Organic compounds
Inorganic compounds NH3
Light (450 to 650 nm)*, H2O Light (≈870 nm)*, H2S
Carbon source Organic compounds
CO2
Organic compounds CO2 O2 CO2 CO2
SO3–2, H2O O2, H2O SO3–2, H2O
Ammonia oxidizers Anaerobic ammonium oxidizers Nitrite oxidizers Denitrifiers Sulfide oxidizers Methanogens Sulfate reducers Phototrophs Phototrophs Phototrophs
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Biofilm Reactors
5.1 Energy Source Three types of energy source are used by bacteria—organic compounds (chemoheterotrophic metabolism), inorganic compounds (chemoautotrophic metabolism), and radiant energy (photosynthetic metabolism). In general, a single species does not use more than one type of energy source, and, for this reason, bacteria often are classified as chemoheterotrophic, chemoautotrophic, or photosynthetic. For example, the Pseudomonads, a common soil bacterial genus, are all chemoheterotrophic. In wastewater treatment, the most important groups of bacteria are chemoheterotrophic. However, the chemoautotrophic nitrifying bacteria are essential in processes where ammonia-nitrogen removal is required. The photosynthetic Cyanobacteria (formerly blue-green algae) are essential components of wastewater treatment processes, such as oxidation ponds and artificial marshes.
5.2 Chemoheterotrophic Metabolism Organic material serves as both an energy and carbon source in chemoheterotrophic (typically shortened to heterotrophic) metabolism. Bacteria are able to use a wide variety of organic materials as food or “substrate,” including complex molecules, such as protein and starches, and many anthropogenic compounds, such as chlorobenzoate and pentachlorophenol. Most bacterial species are capable of metabolizing, or partially metabolizing, a relatively large number of compounds. However, the rates of metabolism of specific compounds vary significantly among species, and, in cases where a mixture of organics are available, a large number of species will grow as a consortium that, in some manner, appears to provide the most efficient biodegradative community. Nearly all naturally occurring organic compounds can be degraded by at least one species of bacteria. However, a number of synthetic organic compounds that have no naturally occurring counterparts (termed xenobiotics after the Greek word xeno, or stranger) have proven difficult to degrade biologically (Madigan et al., 2003; Reineke and Knackmuss, 1988; Rittman and McCarty, 2001). Examples include multi-ring compounds, such as the pesticides Dieldrin, a chlorinated hydrocarbon originally produced in 1948 by J. Hyman & Co., Denver, Colorado, and toxephene; the polychlorinated biphenyls; and chlorinated volatile organic compounds, such as the solvents trichloroethylene (TCE), tetrachloroethylene, and 1,1,2-trichloroethane, which are shown in Figure 2.3. Chlorendic acid is used as a “builder” for other chemicals, many of which are pesticides. Over 200 polychlorinated biphenyls exist. The compounds were used principally as insulators in electrical transformers, but currently are not manufactured in the
Biology of Fixed-Growth Process
Cl
Cl
Cl
Cl
Cl Cl
Cl
COOH Cl Cl
COOH
Chlorendic Acid
Cl Cl A Polychlorinated Biphenyl
Benzene
OH
Napthalene
Phenol
Phenanthrene
Cl
Cl
Cl
Cl
C
C
C
C
Cl Cl Cl H Trichloroethylene Tetrachloroethylene
H
Cl
Cl
C
C
H
Cl H 1,1,2-Trichloroethane
Figure 2.3 Examples of organic compounds that are toxic and/or difficult to biooxidize.
United States because of their carcinogenicity. Benzene is easily degraded biologically, but is highly carcinogenic. Phenol is biodegradable at low concentrations. Napthalene and Phenanthrene are among the most common polynuclear aromatic hydrocarbons. The chlorinated aliphatic solvents TCE, tetrachloroethylene (perchloroethylene, perc, or PCE), and 1,1,2-trichoroethane (1,1,2-TCA) are commonly used chlorinated solvents. Non-degradable xenobiotic compounds are a matter of current concern because of potential toxicity to living organisms. Many have been accumulating in the food chain, and effects of the accumulation, such as decreases in bird reproduction rates, loss of fisheries, and increases in specific malignant tumors, are being discovered continually. Materials that are slow or impossible to break down biologically are referred to as recalcitrant or refractive. Some naturally occurring compounds are refractive. These include cellulose and lignin, both of which are associated with plant fiber. Cellulose is a particularly interesting compound because it is, like starch, a polymer of glucose. Both starch and glucose are easily broken down by a wide variety of bacteria and most other heterotrophic organisms. Starch is broken up into small subunits in reactions catalyzed by enzymes, such as α-amylase and maltase, which break the α-glycosidic bond (Figure 2.4)
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Biofilm Reactors
HOCH2 H
H
H
OH
O
HOCH O
H
H O
H
2
H
OH
O H
OH
H
H
OH
O
α − glycosidic bond a. Repeating unit of starch and glycogen
HOCH
O
H
2
H
HOCH O
OH
H
H
OH
H
O
H
2
H OH H
O
H
H
O
OH
β − glycosidic bond
b. Repeating unit of cellulose
Figure 2.4 Repeating glucose units of starch and glycogen connected by α-glycosidic bonds and cellulose connected by β-glycosidic bonds.
between the glucose molecules. The bond between the glucose molecules of cellulose is a β-glycosidic bond (a mirror image of the α-glycosidic bond), and the enzymes that catalyze the cleavage reaction are unusual in bacteria but common in the fungi. Bacteria in the aerobic Actinomycete group and a few anaerobic heterotrophic bacteria, notably some Clostridia, found in such places as the intestines of termites and certain anaerobic environments, produce the enzyme, but it is lacking in most aerobic species.
5.3 Chemoautotrophic Metabolism Organisms capable of using CO2 or bicarbonate (HCO3–) as a carbon source and a reduced inorganic chemical, such as hydrogen (H2), ammonia (NH3) (or ammonium [NH4+]), nitrite (NO2–), Fe2+, hydrogen sulfide (H2S), or elemental sulfur (S0), as an energy source are referred to as chemolithotrophic, chemoautotrophic, or, more commonly, autotrophic. Specific groups of bacteria oxidize each of the above materials. The ammonia oxidizers (Nitosomonas, Nitrosospira, and Nitrosococcus are the predominant genera) and nitrite oxidizers (Nitrobacter and Nitrospira are the predominant
Biology of Fixed-Growth Process
H O + O2 + FeS 2(s) 2
-2 + H+ 2
Fe +2 + S
_ +2 2 Fe + SO4
+3 Fe + H2 O O 2 + H+
H 2O
Fe(OH)
3(s) +
H+
Figure 2.5 Schematic of reactions through which pyrite [FeS2(s)] is oxidized in mine drainage. Note that the stoichiometry is not shown. genera) are the autotrophs of major interest in biological treatment. Sulfur-oxidizing bacteria (the aerobic Thiobacillus, Thiotrix, and Beggiatoa and the anaerobic photosynthetic green and purple bacteria) are important in many environments, including wastewater collection systems. Sulfides and thiosulfate are oxidized by these bacteria in the aerobic liquid film that forms on the upper surface of open-channel pipes, and the oxidation product is sulfuric acid. The resulting corrosion is difficult to control and results in severe maintenance problems. Iron bacteria cause corrosion in many industrial systems, and species such as Thiobacillus ferrooxidans, Thiobacillis thiooxidans, and Ferobacillis ferrooxidans are responsible for the production of acid mine drainage through catalysis of the reaction Fe+2 → Fe+3 + e-. Oxidation of Fe+2 in mine drainage occurs spontaneously under some conditions, but the bacterially mediated corrosion reaction is 106 times faster (Singer and Stumm, 1970). Acid mine drainage is a major problem in most coal and copper mining areas. The reactions are summarized in Figure 2.5. Molecular hydrogen is a common product in the anaerobic bacterial breakdown of organic materials. If H2 accumulates in a reaction system, organic breakdown will cease and, therefore, a mechanism for removal must be present. Autotrophic conversion of molecular hydrogen and CO2 to CH4 by anaerobic methane-fermenting Archaea is an important process in the treatment of organic sludges and is essential to the functioning of ruminant animals.
5.4 Photosynthetic Metabolism A small number of bacterial species obtain their energy for growth from light and their carbon for growth from CO2 or HCO3–. One group, the Cyanobacteria, formerly known as blue-green algae, carries out oxygenic photosynthesis, in which molecular
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Biofilm Reactors
oxygen is a product. Blue-green bacteria are important in oxidation ponds, a type of biological wastewater treatment process. Other photosynthetic bacteria may occur in overloaded oxidation ponds but are not of importance in biological treatment systems under normal operating conditions. Cyanobacteria have chlorophyll that absorbs light in the 800- to 900-nm wavelength range, and the principal mode of growth is light adsorption with concomitant lysis of H2O. Oxygen is a byproduct, and CO2 is assimilated as the carbon source. Ammonia and nitrate both are used as nitrogen sources. A number of Cyanobacteria species can fix nitrogen. Simple organic compounds can be used as carbon, but not as energy sources, by a number of Cyanobacteria species, and a few species can grow heterotrophically. Thus, the Cyanobacteria are extraordinarily adaptable microorganisms, in terms of their nutrition, and it is not surprising that they predominate in many natural environments. Fixed-growth biological treatment processes do not provide suitable conditions for their growth, however, and the use of Cyanobacteria essentially is limited to faculative and aerobic oxidation ponds. Anoxygenic photosynthesis is carried out by two other groups of phototrophic bacteria—the Rhodospirillaceae, or purple nonsulfur bacteria, and the Chlorobiaceae, or green bacteria. Anoxygenic photosynthesis does not result in oxygen generation and does not occur in the presence of oxygen. The principal wavelengths absorbed by the phototrophic bacteria are between 400 and 500 nm and 700 and 1000 nm. A number of species can grow heterotrophically on simple sugars, amino acids, and organic acids. Although these bacteria typically are considered anaerobes, most are not extremely sensitive to oxygen, and many can grow in micro-aerophilic environments. Neither the green or purple bacteria are of great importance in fixed-growth systems.
5.5 Nutrient Requirements Microbial metabolism is directed toward reproduction of the organisms, and this requires that chemical constituents of cell components are available for assimilation and incorporation to new cell components. The elements listed in Table 2.4 are required in the approximate proportions given. Note that the two empirical formulas given above in the Chemical Composition of Cells section have approximately the same percentages of the principal elements. For example, in the empirical formula C5H7NO2, the percentages of carbon, hydrogen, nitrogen, and oxygen on a mass basis are 53, 6, 12, and 28, respectively. Nutrients typically are assimilated from a limited number of elemental states; therefore, both the amount of nutrients present and the state of the nutrients are
Biology of Fixed-Growth Process
important. For example, heterotrophic bacteria require carbon in the organic form. Many bacterial species are able to use only a limited range of organic compounds. Only a few species are capable of metabolizing the 5-carbon sugar lactose in the absence of oxygen, and this property is used to determine the presence of coliform organisms (species that are common in the gut of warm-blooded animals) in water samples. The most probable number test for coliform bacteria is constructed on the basis of this fact. Most bacteria can assimilate nitrogen in the ammonia (-3), nitrite (+3), and nitrate (+5) oxidation states. Sulfur assimilation typically is from the sulfate (+6) oxidation state. Metals are assimilated from the ionic state, in almost all cases. Often one or two nutrients in the environment limit microbial growth. The limiting nutrient concept is useful in predicting the effects of pollutants on receiving waters and in designing and operating biological treatment processes. The Great Lakes in the United States and Canada are phosphorus-limited, and the increases in phosphorus inflow rates that resulted from the introduction of biodegradable detergents in the late 1960s greatly increased the rates of eutrophication. The removal of phosphorus from wastewater discharges to tributaries of the Great Lakes was an application of the limiting-nutrient concept. Many industrial wastewaters are unbalanced, with respect to nutrients, and stoichiometric additions of the growth-limiting nutrients (typically nitrogen and/or phosphorus) must be made.
5.6 Bacterial Energy Metabolism Respiration is the mechanism through which the electrons removed from energy sources are discharged to the cell’s external environment and when use of a terminal electron acceptor is coupled to energy production. Typically, respiration involves a series of oxidation–reduction (redox) reactions, but the final step involving a terminal electron acceptor defines the type of respiration occurring. Two types of respiration occur—aerobic, in which molecular oxygen serves as a terminal electron acceptor, and anaerobic, in which inorganic compounds or ions, such as nitrate (NO3–), sulfate (SO4–2), S0, CO2, Fe3+, Mn4+, SeO4–2, AsO33–, and ClO4–, or organic compounds, such as fumarate, chlorobenzoate, and dimethyl sulfoxide, serve as the terminal electron acceptor. The terminal electron acceptor is discharged by the cell to the environment; thus, the use of organic compounds as electron acceptors results in the discharge of organics to the liquid outside the cell. Fermentation reactions are used by bacteria to catabolize organic compounds when there is no external electron acceptor available. The organic substrate is used as the electron donor, and the electron acceptor is a high-energy intermediate product
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of the transformation of the donor that can yield ATP (the energy currency of life systems) via substrate-level phosphorylation. There are several types of fermentations, and some, such as the production of CH4 from acetate, are of importance in fixed-film processes.
5.7 Aerobic Growth and Respiration The use of oxygen as a required terminal electron acceptor is the principal characteristic of aerobic metabolism. Heterotrophic metabolism will be used as the model in this section, although important autotrophic organisms, such as the nitrifiers, are also aerobes. Biological treatment systems nearly always are made up of mixed microbial populations, and, because of the metabolic interactions of the various species, the principal end-products of heterotrophic aerobic growth and respiration are new microbial cells, CO2, and water (H2O). Other products may be produced in significant quantities, but this is not the general case. Models of aerobic metabolism generally are based on the conceptual expression given in eq 2.1.
New Organic + O 2 + Nutrients → bacterial + CO 2 + H 2 O matter cells
(2.1)
The new microbial mass term in eq 2.1 includes all synthesized materials, including microbial exopolymers. Experimentally, the value of the new microbial mass term is measured as the increase in either suspended or volatile suspended solids (VSS). Volatile suspended solids measure only organic solids and are better suited to describe biomass in a reactor. Writing eq 2.1 in the form of a stoichiometric balance is not possible in the general sense, because conversions are related to the characteristics of the organic material, the species of organisms present, and the treatment system operating characteristics. However, development of stoichiometric expressions is a necessary step in treatment process design, because estimates of both oxygen requirements and excess solids production must be made. In nearly all applications of biological treatment, a mixture of organic compounds is present, some of which are soluble and some of which are in the particulate form. This situation complicates the application of stoichiometric relationships, because the mixed microbial populations generally attack a wide variety of soluble and particulate materials simultaneously. The result is that surrogate variables for the organic reactant concentration must be used.
Biology of Fixed-Growth Process
The most common surrogate variable is the biochemical oxygen demand (BOD), although chemical oxygen demand (COD) and total organic carbon (TOC) are used, to some extent. In most cases, the BOD value reported in laboratory analyses is the 5-day, 20 °C value. However, the BOD value that represents the concentration of biodegradable organic material present is the ultimate BOD (UBOD). If the systems used in biological treatment were mathematically linear, the BOD5 value could be used, but this is not the case. Thus, care should be taken in process design to ensure that UBOD values are estimated and used in calculations. Models used to evaluate and design wastewater treatment processes generally use COD as a basis for quantifying the strength of the wastewater. This is because COD provides a consistent and convenient method for computing mass balances, which are fundamental to calculating, for example, the carbon entering and leaving a process, waste sludge produced, and oxygen consumed. The COD measures the electron-donating capacity of organic material, and, because electrons cannot be created or destroyed, it provides a conservative basis for these calculations. The issue of how much of the COD represents biodegradable organic matter is dealt with by characterizing the COD in terms of its different fractions. This is discussed further in Chapter 11 on modeling.
5.8 Anoxic Conditions and Respiration Anoxic conditions will be defined to occur when oxygen is absent from the liquid environment, or present at very low levels, and either NO3– or NO2– are available as the exogenous (from outside the cell) terminal electron acceptor. Anaerobic respiration is applied to situations where CO2, SO4–2, or certain organic compounds are available as terminal electron acceptors. While commonly used in environmental engineering practice, this differentiation is not made in the discipline of microbiology, and metabolism is termed aerobic, anaerobic, or fermentative. The term anoxic then is reserved to describe an environment that is free of molecular oxygen. We will adhere to the usage in environmental and wastewater engineering, but caution the reader about conflicting definitions and uses in the literature. A large number of heterotrophic bacterial species are capable of reducing NO3– to NO2– in a manner analogous to the use of oxygen as a terminal electron acceptor. These bacteria are all members of a group termed facultative anaerobes; that is, they can metabolize and grow under aerobic or anaerobic environments. A smaller number of heterotrophic facultative anaerobes are able to reduce NO2– to gaseous N2O and N2, with the end product distribution determined by the pH of the liquid. The organic
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Biofilm Reactors
materials used to drive the reduction processes (termed denitrification or dissimilatory nitrate reduction) are quite varied. Single carbon compounds, such as methanol, CH3OH, can be used, in addition to the more conventional types of organics found in domestic wastes. A second type of NO3– and NO2– reduction, termed assimilatory reduction, is used by many species of bacteria to produce a –3 oxidation state nitrogen for incorporation to organic molecules, such as the amino acids that make up protein molecules. Assimilatory reduction is not tied to energy metabolism, is not analogous to the use of oxygen in aerobic respiration, and occurs under both aerobic and anaerobic conditions. Because energy is expended in reducing NO3– and NO2–, the mass yield of microorganisms will be lower than for systems where NH3 is the nitrogen source. Both obligate aerobes and facultative anaerobes can carry out assimilatory reduction. Nitrogen compounds also can be removed from wastewater using anoxic ammonia oxidation (anammox), with NH4+ acting as an electron donor and NO2– as a terminal electron acceptor. This process is carried out by slow-growing bacteria of the Planctomycetes phylum of Bacteria, which are autotrophic. In biofilms, there exist both oxic and anoxic zones; hence, aerobic nitrifiers produce the NO2– that becomes the electron acceptor necessary for the anammox process.
5.9 Anaerobic Respiration and Fermentative Metabolism The common anaerobic terminal electron acceptors of importance are SO4–2 and CO2. Ferric (Fe+3) ion also is used by some bacteria, including many of those that reduce NO3–. However, the insolubility of Fe+3 makes the ion an insignificant factor as an electron acceptor in the biological treatment of water. The bacterial species capable of dissimilatory SO4–2 reduction are not numerous (limited to approximately 8 genera), and the reaction product, hydrogen sulfide (H2S) typically is noticeable because of its odor. All of the SO4–2 reducers are obligate anaerobes; that is, they do not grow in the presence of oxygen. Common energy sources for dissimilatory SO4–2 reduction are H2, lactic acid, and pyruvic acid. As in the case of NO3–, a large number of bacterial species, both aerobic and anaerobic, can reduce SO4–2 in an assimilatory manner. The ability to reduce SO4–2 is almost essential for organisms growing in soil and water, because nearly all of the available sulfur is in the SO4–2 form. Carbonate respiration involves H 2 as an electron donor and CO 2 as a terminal electron acceptor; it can lead to the formation of CH4 (in methanogenesis),
Biology of Fixed-Growth Process
carried out by a limited number of genera belonging to the Archaea, or acetate (in acetogenesis), which is performed by homoacetogens belonging to the Bacteria. Both groups are strictly anaerobic. Similar to NO3– and SO4–2, there also are a great number of species belonging to all three domains of life that can assimilate CO 2 as a nutrient source without concurrent use as an electron acceptor under aerobic or anaerobic conditions. The use of electron acceptors—those produced within the cells or by other cells— results in the fermentation of a variety of relatively low-molecular-weight organic compounds. Typical products are low-molecular-weight (volatile) organic acids, alcohols, and aldehydes. The principal volatile acids produced are formic, acetic, propionic, and butyric, with acetic being the most important. Methanogenic members of the Archaea, such as Methanosaeta and Methanosarcina, oxidize acetic and formic acids with the concomitant production of CH4 and CO2. Hence, they are distinct from autotrophic methanogens using H2 as an energy source and reducing CO2 (or HCO3–) as a terminal electron acceptor to CH4. Anaerobic growth by fermentation and respiration can be used to completely remove organic materials from water and wastewater only if methane fermentation is included as a step in the process. Methanogens are the final link in a series of interdependent biochemical reactions involving different organisms (termed syntrophic reactions), which characterize anaerobic environments in fixed-growth reactors treating biological waste. The concept is attractive, because the methane has a very low solubility and can be collected and used as a fuel. However, the energy for growth derived from methane fermentation is relatively small, and the rates associated with growth are low. Economic operating conditions generally require that operating temperatures be in the range 30 to 35 °C, which requires that the reactors be heated. Generation of sufficient methane to heat the systems requires that strong, or high-concentration, organic feed solutions are used. Organic sludges produced in wastewater treatment plants (WWTPs) and industrial wastewaters, such as those from food-processing industries, are the principal candidates for application of anaerobic fermentation treatment processes. Anaerobic sludge digestion is one of the most common methods of stabilizing sludges resulting from wastewater treatment, and sufficient excess methane often is produced to allow heating of buildings and, in some cases, production of electricity. In addition, low-temperature (<20 °C) anaerobic treatment of wastewater is possible using fixed-growth processes, but more insights to microbial community changes are needed to ensure stable maintenance of reactor operations (Enright et al., 2007).
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Biofilm Reactors
5.10 Energetics of Respiration The molar free energy released or consumed in a reversible reaction is given by eq 2.2.
∆G = ∆Go + RTlnQ
(2.2)
Where ∆G = free energy change, kJ/mol; ∆Go = standard free energy change, kJ/mol; T = absolute temperature, Kelvin (K); R = universal gas constant, J/K mol; and Q = reaction quotient (Q ≈ K, the equilibrium coefficient in dilute solutions). Electrical potential (in volts) can be related to the free energy change by the Nernst equation, as follows:
∆G = –nFE
(2.3)
Where n = number of electrons transferred in the reaction, and F = Faraday’s constant (9.649 × 104 J/mol∙V). In practice, ∆E is referenced to the standard hydrogen electrode for which a standard free energy change, ∆Go, of 0 is assigned to the half-reaction when the reaction quotient, Q, is equal to 1. 1/ 2 [H2 ] 1 + H + e– = 2 H2(gas) Q = = 1 (2.4) H + Note that, by convention, the activity of a gas is 1 and, therefore, [H+] = 1. Electrical potentials relative to the standard hydrogen electrode are designated by EH and, therefore,
∆G = -nFEH
(2.5)
Substitution of eq 2.5 into eq 2.2 gives the following:
–nFEH = –nF EHo + RTlnQ EH = EHo –
RT ln Q nF
(2.6) (2.7)
For the hydrogen ion reaction of eq 2.2, the standard free energy, ∆Go, is set equal to zero by convention; therefore, the standard electrode potential is also zero. Reactions in the environment typically occur under nonstandard conditions, that is, when the reaction quotient for the hydrogen ion reaction of eq 2.2 is not equal to 1
Biology of Fixed-Growth Process
and the temperature is not equal to 0. Comparisons typically are made for pH = 7 and 25 °C, where the standard potential is as follows: 1/ 2 RT RT [1] EH(W) = EHo – ln Q = 0 – ln −7 nF nF 10 (8.31J / mol ⋅ K )(293K ) ln 107 (1)(9.649x10 4 J / mol ⋅ V )
= −
= 0.0257(16.12) = –0.407 V
Standard potentials for selected environmentally important half-reactions are given in Table 2.7. In each case, the value of EH(W) is obtained by subtracting the functional term RTlnQ/nF from the standard potential. Table 2.7 Standard reduction potentials at 25°C for selected environmentally important redox couples. E (V)
E(W) (V)
∆G(W)/na (kJ/mol)
1.59
1.18
–113.5
1.22
0.81
–78.3
1.24
0.74
–71.4
0.52b
–50.2b
0.83
0.42
–40.5
NO3–+ 10H+ + 8 e– = NH4+ + 3H2O FeOOH(s) + HCO3– + 2H+ + e– = FeCO3(s) + 2H2O Pyruvate + 2H+ + 2e– = lactate SO –2 + 9H+ + 8e– = HS– + 4H O
0.88
0.36
–34.7
S(s) + 2H+ + 2e– = H2S(g) CO2(g) + 8H+ + 8e– = CH4(g) + 2H2O CO2 + HCO3– + H+ + e– = CH3COO– + 3H2O 2H+ + 2e– = H
Half-reaction 2HOCL + 2H+ + 2e– = Cl + H O 2
2
O2(g) + 4H+ + 4 e– = 2H2O 2NO3– + 12H+ + 10 e– = N2 + 6H2O Cl + 2e– = 2Cl– 2
MnO2(s) + HCO3– + 3H + 2e– = MnCO3(s) + 2H2O NO – + 2H+ + 2 e– = NO – + H O 3
2
4
2
2
2
6CO2 + 24 H+ + 24 e– = C6H12O6 + 6H2O a
Based on [HCO3–] = 10 M. Number of moles transferred.
b
–3
1.40
–0.05b
4.6b
–0.19
18.3
0.25
–0.22
21.3
0.17
–0.24
23.5
0.17
–0.25
23.5
–0.29
27.65
0.00
–0.41
39.6
–0.01
–0.43
41.0
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Biofilm Reactors
5.11 Example 2—Calculation of Electrode Reduction Potentials and ∆G for Half-Reactions Determine the electrode potential for the reduction of NO3– to NO2– at temperatures of 15 and 25 °C and pH values of 7 and 10.
5.12 Solution The solution can be obtained as follows:
(1) The half-reaction and standard electrode potential are NO3– + 2H+ + 2 e– = NO2– + H2O EH(W) = 0.83
(2) Determine the electrode potential at 15°C. pH = 7 NO −2 }{H 2 O} { RT EH(W) = 0.83 − ln 2 nF {NO −3 }{H+ }
For equal molar concentrations of NO3– and NO2–, that is {NO3–} = {NO2–}, and remembering that the activity of water is 1.0,
EH(W) = 0.83 −
EH(W) = 0.83 −
pH = 10
RT 1 RT ln = 0.83 − In (107)2 2 + nF nF H
(8.314J / mol ⋅ K)(285K) ln 1014= 0.83 − 0.40 = 0.43 V (2)(9.649 × 104 J / mol ⋅ V) EH(W) = 0.83 −
EH(W) = 0.83 −
RT ln(1010)2 nF
(8.314J / mol ⋅ K)(285K) ln 1020 = 0.83 − 0.57 = 0.26 V (2)(9.649 × 104 J / mol ⋅ V)
(3) Determine the electrode potential at 25 °C.
pH = 7
EH(W) = 0.83 −
(8.314J / mol ⋅ K)(285K) ln 1014 = 0.83 − 0.41 = 0.42 V (2)(9.649 × 104 J / mol ⋅ V)
Biology of Fixed-Growth Process
pH = 10
EH(W) = 0.83 −
(8.314J / mol ⋅ K)(285K) ln 1020 = 0.83 − 0.59 = 0.24 V (2)(9.649 × 104 J / mol ⋅ V)
Thus, the electrode potential is not very sensitive to the temperature value (within the typical environmental operating range). However, the potential value is quite sensitive to the operating pH. A sense of the relative attractiveness of electron acceptors in biochemical reactions can be attained by comparing the free energy change of the overall redox equations. An example is the oxidation of glucose (C6H12O6) using oxygen (O2) and NO3– as the electron acceptors. Determination of the free energy change is accomplished by application of the Nernst equation (eq 2.3) using the standard reduction potentials available in the literature. For the glucose –O2 and glucose NO3– reaction systems, the relevant half-reactions are as follows:
6CO2 + 24 H+ + 24 e– = C6H12O6 + 6H2O
(2.8)
O2(gas) + 4H+ + 4 e– = 2H2O
(2.9)
2NO3– + 12H+ + 10 e– = N2 + 6H2O
(2.10)
The free energy change for the two reaction systems is compared by writing the half-reactions on a per-electron basis, as follows: 1 1 1 CO2 + H+ + e– = C6H12O6 + H2O E(W) = − 0.43 (2.11) 4 24 4
1 1 O (gas) + H+ + e– = H2O E(W) = 0.81 4 2 2
(2.12)
1 6 1 3 NO3– + H+ + e– = N2 + H2O E(W) = 0.74 5 5 10 5
(2.13)
The ∆G values for the total reactions then are calculated, as follows: Glucose/O2
1 C H O + O2(gas) = CO2 + H2O 6 6 12 6
∆E(W) = E(W)reductant – E(W)oxidant = 0.81 − (−0.43) = 1.24 V ∆G(W) = -F∆E(W) = –119.6 kJ/mol
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Glucose/NO3–
C6H12O6 + NO3– + H+ = CO2 + N2 + H2O
∆E(W) = E(W)reductant – E(W)oxidant = 0.74 − (−0.43) = 1.17 V
∆G(W) = –F∆E(W) = –112.9 kJ/mol
Oxygen is a more attractive electron acceptor than nitrate, on the basis of available energy from the reaction, and this is one reason that oxygen dominates among the range of electron acceptors used by bacteria.
5.13 Co-Metabolism Organic compounds that do not serve as growth substrates are sometimes broken down through a process commonly called co-metabolism or incidental metabolism. The biodegradation process typically is initiated through the action of an enzyme having little specificity, which typically attacks growth substrates. In some cases, carbon from the breakdown of materials co-metabolized has been shown to be incorporated to cells, which may indicate that products of the initial reactions may be growth substrates. However, often, the product of co-metabolism is a dead end, that is, another non-growth substrate. The most widely studied co-metabolic reactions are associated with the breakdown of chlorinated organic compounds and, in particular, the widely used solvent TCE. Aerobic biodegradation of TCE is initiated by methane monooxygenase enzymes used to break down methane (methanotrophic bacteria) (AlvarezCohen and McCarty, 1991); ammonia monooxygenase (Hyman et al., 1988), which catalyzes the conversion of ammonia to hydroxyl ammine in members of the genus Nitrosomonas; and phenol monooxygenase and toluene dioxygenase, which catalyze the first steps in the breakdown of phenol and toluene, respectively, in members of the genus Pseudomonas (Fan and Scow, 1993). Co-metabolism of TCE presents the following two interesting problems:
(1) TCE strongly competes for the active site with the growth substrate (see the Consortia section); and (2) The reaction product, TCE epoxide, apparently sorbs onto the oxygenase enzyme and destroys the catalytic activity.
Incorporation of co-metabolic processes to treatment systems presents real challenges.
Biology of Fixed-Growth Process
6.0 BACTERIAL GROWTH The objective of biological treatment is the development of a process that will transform contaminants in a water or wastewater into an acceptable form or state. However, from the viewpoint of the microorganisms used in treatment processes, the objective is simply to replicate themselves. The result of replication is an increase in total microbial mass, or mass concentration, and this increase is defined as growth. In some cases, the microbial objective is modified to surviving austerity until replication is possible again. Examples of austerity include dry and low-temperature periods in soil and the segments of biological wastewater treatment systems in which the primary organic substrates are exhausted. Bacterial replication, or growth, requires energy; carbon (to build the organic compounds making up the cell); and inorganic nutrients, such as nitrogen, phosphorus, sulfur, and iron, which are components of organic compounds or serve in other ways in the chemical reaction processes carried out within the cell. Heterotrophic metabolism, where organic compounds serve as both the energy source and the carbon source, is the most important type of metabolism in biological treatment and will be used as a model.
6.1 The Bacterial Growth Curve When bacteria are grown in a batch culture with the initial condition that nutrients are not limiting, a characteristic pattern of growth occurs, which is commonly divided into the six segments shown in Figure 2.6—lag, increasing growth, logarithmic growth, declining growth, stationary, and death. The bacterial growth curve shown in Figure 2.6 is representative of pure (i.e., single species) cultures only, but the concepts generally are applied to other systems in a modified form. In the experimental process of producing a growth curve, the number of cells is determined by sampling at selected time intervals and using standard microbial plate-count techniques to estimate the population. A similar curve can be generated by measuring the mass concentration of cells present, and, because each cell can be presumed to have approximately equal mass, the two curves should be proportional. The limiting nutrient typically is the energy source (an organic compound for heterotrophic species), but could be an inorganic nutrient, such as nitrogen or phosphorus. Typically, the cells used in growth-curve experiments are taken from cultures in the stationary phase, and a period of time, or lag, is required before growth is initiated. The lag can be eliminated by transferring cells from a culture in the logarithmic
51
Stationary
Decreasing Growth
Lag No
Log Growth
Co
Increasing Growth
Limiting Nutrient Conc., mg/L
Biofilm Reactors
Log10 Number of Cells
52
Death
Figure 2.6 Bacterial growth curve for a batch system in which nutrients initially are not limiting. As growth begins, the increase in cell number (or mass) is dependent only on the number (or mass) of cells present, and the system behaves in a firstorder autocatalytic manner.
growth phase. A short period of time, in which the growth rate changes from zero to the maximum value, exists (the increasing growth phase), which is probably the result of an uneven conversion of the population from non-growth to growth conditions. During the logarithmic growth phase, the rate-limiting variable is the number of cells present. Because the cells are a reaction product, the overall rate typically is observed to be first-order autocatalytic, with respect to the cell number or mass concentration. As the limiting nutrient concentration decreases to the point where the individual cells must compete for the material, the overall growth rate slows (decreasing growth), and, when the limiting nutrient concentration approaches zero, the culture begins to convert to a resting state, in which cell division essentially stops, and cell functions are maintained by catabolizing unnecessary materials and storage compounds. The unnecessary materials include transferase enzymes necessary for transporting nutrients across the cell membrane and internal metabolic enzymes
Biology of Fixed-Growth Process
used in processing externally derived nutrients. Often, the capsular layer increases in size, possibly to limit desiccation in dry soils. After a period of time extending from several hours to days, the culture begins to die. Upon death, cells typically breakup (lyse), and many cell components become nutrient sources for surviving cells.
6.2 Growth in Mixed Cultures Mixed microbial cultures typically are dominated by a few species of bacteria. However, a large number of microbial groups typically are present, and all of the groups present occupy ecological niches. In most cases, several species present will be able to at least partially metabolize organic compounds initially present. As breakdown progresses, other species may be able to metabolize partially degraded materials or excreted waste products of the first metabolizing groups. Higher forms of organisms will begin to graze on the heterotrophs as the population increases. An interlocking set of growth curves is produced that cannot adequately be described by cell counts of a limited number of species. Total mass concentration commonly is used to follow the process, but the values measured are the sum of all species present, and, as time increases, they increasingly represent the growth of secondary organisms. This interaction of microbial groups in mixed cultures is one reason that kinetic and stoichiometric parameters measured under one set of conditions often are difficult to extrapolate to other conditions.
6.3 Enrichment Cultures When mixed cultures are operated under constant conditions, either steady-state feed and physical environment or repeated identical cycles, species with the highest growth rates will have a competitive advantage and will gradually predominate. In some cases, the procedure results in cultures limited to one or two species. Such operating conditions are called enrichment cultures, and the technique is used in the laboratory to isolate and grow species having particular desired characteristics. Enrichment culture conditions often exist in industrial wastewater treatment because of the limited number of biodegradable compounds present, the high temperature, or the low pH of the wastewater. In biological water and wastewater treatment, enrichment culture principles are used to develop flocculant cultures for suspended-growth systems and polyphosphate-storing cultures for phosphate-removal systems and also to control the growth of filamentous bacteria.
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Biofilm Reactors
6.4 Stability of Mixed Cultures The use of mixed cultures for water and wastewater treatment results in more stable process operation. In most treatment operations, characteristics of the influent stream are changing constantly. Flowrate, chemical constituents concentrations, and temperature change on a continuous basis. Such varying conditions help maintain the mixed culture, by providing environmental conditions in which a broad spectrum of bacterial species can compete. The principal advantages of maintaining mixed cultures are that a more complete breakdown of organic compounds present occurs and that changes in influent characteristics typically do not greatly affect effluent quality. Bio-oxidation by single species of bacteria often results in some compounds being left unoxidized. In many cases, certain branched aliphatic compounds or substituted rings are attacked by a limited number of bacterial species. An organic polymer may be broken down partially by several species in pure culture and completely broken down in mixed culture. Similarly, when influent characteristics change, we might expect different microbial species to have competitive advantages. For example, the predominant bacterial species present may change from summer to winter in municipal treatment plants or with the products being produced in an industrial operation. Maintaining mixed cultures allows the most competitive species for a given set of conditions to predominate in a matter of days.
6.5 Effects of Environmental Variables Environmental variables affecting bacterial growth include temperature, pH, chemicals in the water, and the presence of surfaces on which microbial films can be established. Temperature strongly affects the reaction rates; has minor, if any, effects on reaction stoichiometry (Flegal and Schroeder, 1976); and, under conditions where molecular diffusion is important, mildly affects the rate of transport of nutrients to the cells (Kehrberger et al., 1964). The relationship between temperature and reaction rate differs with microbial species, and, in mixed cultures, temperature changes may result in significant changes in predominant species. Temperature-growth rate relationships are discussed in greater detail in the following section. Bacteria typically function well within broad pH ranges. Most of the bacterial species commonly found in wastewater treatment processes grow between pH 6 and 9, but activity falls off rapidly outside of that range. Nitrifying bacteria grow best under slightly alkaline conditions, such as pH 8 to 9.5. The sulfur-oxidizing bacteria
Biology of Fixed-Growth Process
are quite acid-tolerant, and some species grow well at pH values as low as 1, but do not grow at all at pH values above 6. When the growth rate is reported as a function of pH (Grady et al., 1999), clear optimum pH values are rarely apparent, and the curves are virtually flat within the region of growth. The role of dissolved organic and inorganic chemicals as nutrients, energy sources, and toxins is discussed elsewhere in this chapter. Total solute concentration, typically approximated satisfactorily by the salinity of water, is of some significance, because many industrial wastes are quite saline, and because the discharge of treated wastewater to saline environments is common. Biological treatment of saline wastewaters may be affected by differing microbial sensitivities to salinity; that is, species predominance may be affected. In discharging treated wastewater to saline environments, bacterial survival may be related to salinity, but the available data are not adequate to suggest modification of discharge requirements for ocean versus freshwater conditions. Most bacterial species are able to adjust to broad ranges in salinity (e.g., from less than 100 mg/L to 30 000 to 34 000 mg/L in seawater) and osmotic pressures if changes are made incrementally. A few species of bacteria, known as halophiles, grow at concentrations as high as 360 000 mg/L (Niedhardt et al., 1990). Large, rapid decreases in salinity may result in the leaching of water from cells and damage to the cytoplasmic membrane. Large, rapid increases in salinity may result in the accumulation of water within cells to the point that they burst. In general, salinity is not a major factor in the design or operation of biological treatment processes.
7.0 BACTERIAL GROWTH KINETICS IN BIOFILMS Bacterial growth is the result of a large number of extra- and intracellular reactions that provide energy, carbon, and nutrients for the biosynthesis process. Design and operation of biological treatment processes are related strongly to the rate of bacterial growth, and, for this reason, discussion of bacterial growth kinetics is essential. However, bacterial growth theory was developed from experiments with pure cultures, while biological treatment processes are almost always composed of mixed cultures. Rates of growth and nutrient removal typically are measured using lumped or surrogate parameters, such as suspended or volatile solids, BOD, COD, or TOC, and global rates observed cannot be extrapolated or generalized. Even with this limitation, the application of rate process concepts is an essential component of biological process design and operation.
55
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7.1 Rate of Bacterial Processes Rate expressions used to describe the removal of organic material, growth of the microbial population, and use of oxygen, the three terms of greatest importance in eq 2.1, typically are based on the modified Monod model given in eqs 2.14 through 2.16.
ro = −
kC X K +C
kC − kd X rg = − Yro − kd X = Y K +C
rO2 = ro + γ rg = ro − γ (Yro + kdX)
= ro(1 − γY) − γkdX
(2.14) (2.15)
(2.16)
Where ro = rate of removal of organic material (g/m3·d); rg = rate of production of new microbial mass (g/m3·d); rO2 = rate of oxygen use (g/m3·d); k = rate of removal coefficient (d–1); C = organic concentration, often measured as UBOD or COD (g/m3); X = microbial mass concentration measured as suspended solids or volatile suspended solids (g/m3); K = saturation coefficient (g UBOD/m3); Y = mass yield coefficient (g solids produced/g UBOD removed); kd = maintenance energy rate coefficient (d–1); and γ = ratio of UBOD removed to cell mass produced (g/g).
7.2 Note
C X, where µm is the maximum K +C specific growth rate or the maximum mass of cells produced per unit mass present per unit of time. This nomenclature is awkward from the process analysis viewpoint because of the way that the maintenance energy must be introduced. The solids production rate, rg, and oxygen uptake rate (OUR), rO2, are shown as functions of the organic removal rate, and more specifically the organic (UBOD) and microbial mass concentrations. These expressions (eqs 2.14 through 2.16) will be used to describe overall metabolism in all types of biological processes. The growth rate often is written as rg = µm
Biology of Fixed-Growth Process
Often, the specific rates, or rates per unit mass of microbial mass present, are reported rather than the overall rate. For example, the specific removal and growth rates are stated as follows:
Ro = Rg =
ro kC =− X K +C
rg X
=Y
kC − kd K +C
(2.17) (2.18)
In some commonly used biodegradation and growth models, rate expressions are written for two or more components of the organic mixture. For example, the UBOD may be divided into soluble and particulate portions, and separate expressions may be written for each component (Henze et al., 1995; Metcalf and Eddy, 2003). The ratio of UBOD removed to cell mass produced, γ, is based on chemical stoichiometry and empirical cell formulations. The value also can be referred to the theoretical BOD of the cells produced. The most commonly used empirical formula for cells is C5H7NO2, but a number of other formulas are in use (McCarty, 1965; Porges et al., 1953). A value of γ is estimated by calculating the stoichiometry of the oxidation of the empirical cell formulation, as shown in eqs 2.19 and 2.20.
C5H7NO2 + 5 O2 → 5 CO2 + NH3 + 2H2O γ=
5(32) = 1.42 g/g 113
(2.19) (2.20)
Other suggested empirical formulas result in similar values of γ. For example, the formula C60H87O23N12P has a theoretical UBOD of 1.39 g/g. It is important to remember that eqs 2.14, 2.15, and 2.16 are global, nonstructured, empirical models, which sum up the results of a large number of coupled reactions, including interphase transport from the liquid to the solid cell surface. The assumption that growth of new cell mass and uptake of oxygen are coupled directly to nutrient removal is not completely true. Organic material may be removed from the liquid by adsorption onto/into biofilms without immediate production of new cell material and without oxygen uptake (Dobbs et al., 1989). If the sorbed material is biodegradable, a fraction of the sorbed mass will show up later in time, as new cell mass and a corresponding oxygen uptake would result. These events are illustrated in Figure 2.7, and descriptions of these more complex phenomena are termed structured models.
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Concentration, mg/L and Rate, mg/L-s
58
Total Organic Carbon mg/L
Cell Mass, mg/L
Oxygen Uptake Rate, mg/L-s
Time, min
Figure 2.7 Relationships between organic removal, cell mass production, and OUR.
As shown in Figure 2.7, organic carbon removal occurs more rapidly than the actual breakdown or degradation of the compounds. As a result, oxygen uptake remains above background levels for a substantial period following carbon removal. Following organic removal, the cell mass concentration may continue to increase somewhat, may be approximately constant (as shown in Figure 2.7), or may decrease slowly, as conversion of the externally stored organic material occurs. When stored organics are used completely, the OUR decreases to the “basal” or “endogenous” value, and cell mass begins to decrease, as cell components, such as ribosomes and unneeded enzymes, are broken down. The use of lumped parameters as measures of the organic concentration (UBOD and COD) and the microbial mass (suspended solids or VSS) in eqs 2.14, 2.15, and 2.16 generally is required because of the complexity of the organic mixtures and microbial populations present in most situations. Compounds are taken up in parallel rather than sequence, and a single microbial species can metabolize more than one type of compound simultaneously. A problem exists where a specific compound must be removed, particularly if the target compound is present at a low concentration. In such cases, removal of the BOD or COD to satisfactory levels may not result in satisfactory degradation of the target compound. In addition, the equations, in the
Biology of Fixed-Growth Process
form presented, do not include either mass transport limitations or consideration of the reactions or reaction sequences that are involved in the biological treatment process. Lumping the organic species present into a single concentration term generally is satisfactory. In most cases, it would be impossible to monitor individual organic contaminants, and, perhaps more importantly, the use of overall concentration parameters based on oxygen consumption (UBOD and COD) provide direct relationships to other parameters of importance, such as growth. Finally, in a given source, the contaminant makeup typically is relatively consistent, and surrogate parameters, such as UBOD, COD, and TOC, provide the necessary rate information. An argument similar to that for the use of overall parameters for organic concentration can be made for the use of an overall mass concentration for the microbial mass. The assumption is made that the actual cell mass concentration is proportional to the particulate filtered solids mass concentration (suspended solids or VSS) measured. This assumption would mean that the EPS-mass-to-cell-mass ratio was constant, for example. Although this and other similar assumptions are somewhat shaky, the models have had considerable utility, and laboratory and field measurements typically follow the patterns shown in Figure 2.7.
7.3 Physical and Chemical Changes in Biofilms Resulting from Growth The growth of bacterial cultures is defined by increases in cell number rather than cell size. In biofilms, daughter cells separate as EPS is produced, keeping the density approximately constant. Because growth is constrained latterly and by the support surface, depth increases as cells divide. Cells at the liquid or atmospheric surface have greater access to nutrients and oxygen. Deeper in the biofilm, cells may starve and die, or anoxic or anaerobic conditions may develop. Thus, gradients in oxygen and nutrient concentration and pH develop as biofilms become thicker (Lewandowski and Beyenal, 2003). In thick biofilms, gas formation at the support surface may result in detachment of the biofilm (Atkinson, 1974), and, where a liquid is flowing over the biofilm, as in trickling filters, weakened portions may be sloughed off.
7.4 Structured Models Inclusion of reaction sequences or reaction process sequences in reaction models is termed structuring. Flow diagrams for two structured models are shown in Figure 2.8. In both cases, a storage term is included, which accounts for non-stoichiometric
59
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Organic materital
r1
→
Storage r2 w ithin → cell
B iodegradation and synthesis of new cell mass
a) Simple structured model of organic metabolism based on intracellular storage Organic material
r1
Intracellular metabolite pool
r
r2
3
Biodegradation and synthesis of new cell material
r4
Intracellular Storage
b) Structured model of organic metabolism based on maximum rates and intracellular storage. An assumption is made that the rate of transport into the cell can exceed the rate of conversion of organics. Materials not reacted are stored, presumably as glycogen or PHB. The stored materials are used during periods when the intracellular metabolite pool is low.
Figure 2.8 Structured models of metabolism include reactions or steps in the metabolic process that can be identified or separated: (a) the rate of metabolism is separated from the rate of organic removal from the liquid and (b) a somewhat more complex storage mechanism is shown, which could be used to account for several observed rates of growth and oxygen uptake.
differences in organic removal and OURs that have been observed. More complicated structured models have been proposed to describe the rates of individual cellular processes, such as protein synthesis. Because each rate term requires knowledge of additional rate coefficients, the use of structured models is difficult, even in research. Application of structured models in biological treatment processes composed of a wide range of microbial species and degrading a mixture of organics currently is not feasible.
7.5 Temperature Effects Removal and growth rates are affected strongly by temperature. Rate–temperature relationships typically follow the van’t Hoff-Arrhenius relationship up to approximately 40 °C (Flegal and Schroeder, 1976). In most cases, a modified form of this relationship is used to relate rate constants at a temperature, T, to those at a reference temperature, TR (typically 20 °C).
Biology of Fixed-Growth Process
kT = kTRθT − TR
(2.21)
Values of the temperature coefficient, θ, typically are in the range 1.02 to 1.06. Including temperature effects in biological treatment process design is critical. Most laboratory studies reported in the literature are run at 20 °C, but winter temperatures may drop to 5 °C in cold climates. Summer operating temperatures in municipal plants may reach 30 °C in hot climates, and some industrial wastewaters may approach 40 °C. Achieving satisfactory removals is a problem at low temperatures, unless the decrease in the value of the rate constant has been considered, while oxygen transfer may be limiting at high temperatures because of the decreased solubility and the increased removal rates. For a θ of 1.04, a 10 °C temperature differential from the reference condition results in an increase in the rate constant of 48%.
7.6 Example 3—Effect of Temperature on Organic Removal A batch biological reactor has an initial cell concentration of 30 mg/L and an initial organic concentration of 400 mg UBOD/L. Oxygen and inorganic nutrients are present in excess throughout the reaction period. The rate and stoichiometric coefficients for the system are given below. Compare the performance of the system at 10, 20, and 30 °C, assuming that the maintenance energy requirement (characterized by -kdX) is negligible. Use a temperature coefficient, θ, of 1.02.
k20 °C = 4.0 d–1, K = 20 mg/L Y = 0.40
7.7 Solution The solution can be obtained as follows:
(1) Determine a relationship between the cell concentration and the change in BODU. Because the maintenance energy requirement is negligible, in this case,
∆X = –Y∆C Xt – X0 = –Y (Ct – C0) Xt = –Y (Ct – C0) + X0 = 0.4(400 mg/L – Ct) + 30 mg/L
(2) Write the mass-balance equation for the BODU of the batch reaction system. In − Out + Generation = Accumulation
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0 - 0 + Vr0 = V
∫
C C0
t K +C = – k ∫ dt 0 C(Xo + YCo − YC )
CXo K Xo + YCo − YC 1 In In = − kt Co (Xo + YCo − YC ) Y Xo + YCo Xo
dC dt
dC kC =− (Y − (C0 − C )+ X0 ) dt K +C
(3) Determine the rate constant values at 10 and 30 °C
kT = K20 °CθT − 20 = 4.0 (1.02T − 20)
k10 °C = 3.28
k30 °C = 4.88
(4) The equation can be solved and tabulated and plotted for the three temperature values (Figure 2.9 and Table 2.8). 500
200
400 300 200 10°C
100 0 0.0
30°C
150 SS, mg/L
BOD, mg/L
62
30°C
0.2
0.4
0.6
0.8
Time, days
20°C
1.0
20°C 10°C
100
50 0
1.2
0.0
0.2
0.4
0.6
0.8
1.0
1.2
Time, days
Figure 2.9 Effect of temperature on organic removal rate and microbial growth in example 3. Initial BODU and suspended solids concentrations were 400 mg/L and 30 mg/L, respectively. The initial conditions provide for near-exponential growth because K << Co.
Biology of Fixed-Growth Process
Table 2.8 BODU and suspended solids concentrations for the three temperature values. Time (days)
BODU (mg/L)
Suspended solids (mg/L)
10°C
20°C
30°C
10°C
20°C
30°C
0
400
400
400
30
30
30
0.2
377
372
364
39
41
44
0.4
348
333
310
51
57
66
0.6
310
280
234
66
78
96
0.8
260
205
120
86
108
142
0.9
232
160
43
97
126
173
1.0
197
100
111
150
1.1
158
40
127
174
1.15
138
135
190
0.1
7.8 Inhibition and Toxicity Inhibition of microbial metabolism occurs in the presence of materials toxic to the microorganisms or when non-metabolizable materials present adhere to cell surfaces and block the attachment of substrates. Note that metals are toxic in the ionic form and that the oxidation state can be expected to affect the toxicity. Ammonia in the dissolved gas form is toxic to fish at very low concentrations (<0.5 mg/L, in most cases) and to microorganisms at concentrations that rarely occur in WWTPs. The effects of inhibitory substances on growth and removal rates typically are modeled by adding a term to the rate expressions, which results in a maximum removal rate occurring at a critical inhibitant concentration. The Haldane expression (Yu and Semprini, 2004), originally developed as a model for enzyme inhibition by the reaction substrate, often is used to describe rate inhibition that occurs when otherwise biodegradable compounds are present at high concentrations (Figure 2.10). kC 2 ro = − K + C + C Ki
where
i
X
= Haldane inhibition coefficient (mg/L).
(2.22)
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3.5 3.0 Specific Removal Rate, d–1
64
Nonmetabolizable inhibitant k = 4 d–1 K = 30 mg/L K'i = 50 mg/L CSubstrate = 100 mg/L
2.5 2.0 1.5
Substrate Inhibition k = 4 d–1 K = 30 mg/L Ki = 50 mg/L
1.0 0.5 0
0
20
40
60
80
100
Concentration, mg/L
Figure 2.10 Effect of inhibitant concentration on specific removal rate for two conditions: (a) substrate inhibition by a metabolizable inhibitant, represented by eq 2.22, and (b) a nonmetabolizable inhibitant, such as a heavy metal, represented by eq 2.23.
The application of eq 2.22 results in modified forms of eqs 2.17 and 2.18. In cases where inhibition results from the presence of non-metabolizable compounds, such as heavy metals, the inhibition relationship can be expected to be inversely proportional to the inhibitant concentration.
kC K'i ro = − X K + C K'i + CT
Where
K’i = inhibition coefficient (mg/L) and
CT = inhibitant concentration (mg/L).
(2.23)
7.9 Mass-Transfer-Rate Limitations Mass-transfer-rate limitations have been studied by a number of researchers (Atkinson, 1974; de Beer et al., 2004; Logan, 1993; Schroeder and Tchobanoglous, 1976; Swilley
Biology of Fixed-Growth Process
et al., 1964; Williamson and McCarty, 1976). Suggestions have been made that various rates are limited in both suspended- and attached-growth systems and that the limitations affect both process performance and microbial species makeup. Oxygen or one or more nutrients may be mass-transport-limited in the biofilms of fixed-growth processes. Such limitations result in changes in the type of metabolism (e.g., from O2 to NO3– as the terminal electron acceptor) and species predominance and removal and growth rates. A biofilm that is aerobic near the air interface and anaerobic near the supporting surface may have a significant pH gradient with depth.
8.0 KEY TRANSFORMATIONS IN BIOFILMS Chemical transformations in biofilms, as in all microbial processes, are associated with microbial growth. Sources of cell constituents are extracted from the liquid phase. Particulate matter is broken into soluble subunits through reactions catalyzed by extracellular enzymes, and soluble matter is transported across the cell membrane and converted into the necessary cell components in a complex series of enzymecatalyzed reactions. A parallel process exists for generating the energy required for cell functions. In nearly all fixed-growth processes used for wastewater treatment, chemical rather than photosynthetic energy is used, and the energy is released through redox reactions. Organic compound removal in biological processes occurs in a series of steps (Figure 2.11). The mechanism of nutrient transport to the biofilm–liquid interface differs somewhat with the type of unit. The first step, transport of the organic material to the biofilm surface, is brought about by turbulent transport in most reactor types. Only trickling filters are characterized by laminar flow in the liquid film. The mechanism of large molecule and particle attachment is not fully understood and probably involves sequential sorption and desorption. Biofilm surfaces are not smooth on the microscale, and the mechanisms of particle removal, such as interception and impaction, also may be significant. Smaller molecules pass into the spongelike biofilm structure until they come in contact with a bound enzyme or cell surface. A number of bacterial species typically are present, and the predominant species change with environmental conditions or the type of contaminants present. Changes in species predominance would be expected to take some time because of the low average growth rate imposed in most biological treatment processes and because the position of the bacteria in the biofilm would tend to make direct competition for substrate a slow process to develop.
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Transport of particles to biofilm surface Organic particle
Attachment of particles to biofilm
Biofilm
Extracellular breakdown and intrabiofilm transport
Biofilm
Transport into cells and microbial degradation
Biofilm
Figure 2.11 Steps in the removal and breakdown of an organic particle in a biofilm.
The biofilm serves as a storage unit for extracellular enzymes and enzymes released from dead cells and as a concentrator for nutrients. Biological treatment processes do not provide rich growth conditions for bacteria and in fact are operated with just the opposite objective—to produce an environment with minimal concentrations of necessary growth nutrients. Bacteria growing under such conditions need to minimize energy expenditures made in gathering nutrients and maximize the energy channeled toward reproduction. Evidently, the biofilm structure is one response to these constraints.
8.1 Chemoheterotrophic Processes As noted above, in chemoheterotrophic processes, organic compounds serve as both energy and carbon sources for growth. Aerobic processes are of principal interest, because a broader range of compounds can be used, and the majority of compounds
Biology of Fixed-Growth Process
can be oxidized to CO2 and, therefore, completely stabilized. In all cases, new microbial cells are a product, and only a portion of the organics metabolized are oxidized.
8.2 Chemoautotrophic Processes When carbon dioxide or bicarbonate ion, HCO3–, serves as the carbon source and an inorganic ion (e.g., NH3, NO2–, H2S, S, and H2) serves as the energy source for growth, the metabolism is termed chemoautotrophic or simply autotrophic. The oxidation of nitrogen is important in wastewater treatment, because the reaction provides a method of eliminating ammonia toxicity to fish in receiving waters. Sulfide oxidation presents a serious problem in collection systems and treatment plants, because the products (H2SO3 and H2SO4) are extremely corrosive. Hydrogen serves as an energy source in some of the methane-producing reactions that produce methane in anaerobic sludge treatment processes.
8.3 Biology of Nitrogen Transformations Nitrogen exists in oxidation states from −3 to +5, with the most common forms found in wastewater being nitrite (NO2–) and nitrate (NO3–). Cellular nitrogen is entirely in the organic (the majority of compounds can be oxidized to CO2 and, therefore, completely stabilized −3 oxidation state) form. Ammonium is used directly in the growth process, while oxidized forms of nitrogen must be reduced from the +3 or +5 oxidation states using energy supplied by heterotrophic or autotrophic redox reactions in a process referred to as assimilatory nitrate reduction. Assimilatory nitrate reduction is virtually ubiquitous among microorganisms and, beyond resulting in small decreases in cell yield, because of the energy requirement, has little effect on processes.
8.4 Denitrification Under oxygen-limited or anaerobic conditions, many species of bacteria are able to use NO3– and NO2– as terminal electron acceptors, a process referred to as dissimilatory metabolism. Nitrate is reduced to N2 via NO2–, nitric oxide (NO), and nitrous oxide (N2O). This is a beneficial process in wastewater treatment, because it results in an innocuous end product, whereas both NO and N2O are greenhouse gases, and their accumulation is undesirable. The reductase enzymes are repressed at oxygen concentrations greater than 2.5 mg/L, yet, at low concentrations of oxygen (tenths of mg/L), they are only inhibited (Rittman and McCarty, 2001). As a result, it is possible for
67
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Biofilm Reactors
enitrification to occur, even when residual oxygen is present. Biofilms particularly are suited to supporting multiple types of respiration because of their heterogenous structure.
8.5 Aerobic Nitrification The most important autotrophic bacteria in biological treatment systems are the nitrifiers—those organisms that oxidize NH3 and NO2–. Reduced forms of nitrogen are a source of oxygen demand in water, as shown in eqs 2.24 and 2.25.
NH4+ +
3 O2 → NO2– + 2 H+ + H2O 2
(2.24)
NO2– +
1 O2 → NO3– 2
(2.25)
Equations 2.24 and 2.25 are not completely correct, because carbon dioxide consumption and cell production are not included. Because very little energy is produced from the oxidation of NH3 and NO2–, and production of new cell material from CO2 requires a relatively large amount of energy, the cell mass yield per unit mass of nitrogen oxidized is small (typically approximately 0.05 g cells/g N). Estimates of oxygen requirements typically are based on the above equations. Thus, when nitrification occurs, approximately 4.57 g O2 must be supplied per gram of nitrogen oxidized. Domestic wastewaters typically contain 30 to 50 g/m3 of total Kjeldahl nitrogen (TKN), which includes organic nitrogen, NH3, and NH4+. Biodegradation of nitrogen-containing organic compounds results in release of the nitrogen as NH3. Thus, for the purpose of estimating oxygen demand, organic nitrogen can be treated as ammonia-nitrogen. Therefore, complete oxidation of TKN (nitrogen in the −3 oxidation state) in domestic wastewater will result in an oxygen demand of 130 to 190 g/m3. Estimates of the complete stoichiometry of the nitrification reactions have been made using “reasonable” values for the cell yield coefficients (Metcalf and Eddy, 2003); however, supporting data are difficult to obtain. Typical expressions are given in eqs 2.26 and 2.27. 55 NH4+ + 76 O2 + 109 HCO3– → C5H7NO2 + 54 NO2– + 57 H2O + 104 H2CO3
(2.26)
400 NO2– + 195 O2 + NH4+ + 4 H2CO3 + HCO3– → C5H7NO2 + 3 H2O + 400 NO3–
(2.27)
An important feature to note is that, for every mole of ammonia oxidized, approximately 2 mol H+ are produced. This means that 2 mol alkalinity are destroyed in
Biology of Fixed-Growth Process
the process, and, in many cases, nitrification results in significant decreases in pH. Because most biological processes work best in the pH range 6 to 8.5, nitrification can result in decreased efficiency or even process failure, if care is not taken to control alkalinity. The yields predicted from eqs 2.23 and 2.24 are 0.15 and 0.02 g cells/g N, respectively. The nitrifying bacteria are divided into two groups—those that oxidize NH3 and those that oxidize NO3–. The most widely studied NH3-oxidizing bacteria are of the genus Nitrosomonas, but several other genera exist. Nitrosococcus spp. and Nitrosomonas spp. commonly are found in wastewater, and members of the genus Nitrosospira may be the most common in soil (MacDonald, 1978; Stephen et al., 1996). Nitrosomonas also has an anaerobic metabolism and can use dinitrogen tetroxide (N2O4), a dimer of NO2–, as a terminal electron acceptor. Activity is low under anoxic conditions. N. eutropha also is capable of simultaneous denitrification; that is, it uses NO 2– as a terminal electron acceptor at O2 concentrations below 0.8 mg/L. It is estimated to denitrify at least 10% of the converted ammonia (Lipschultz et al., 1981). During this process, NO and N2O are produced. Nonetheless, all of the NH3-oxidizing bacteria are classified as obligate aerobes, because the oxidation of ammonia via hydroxylamine appears to be the only source of reducing power for biosynthesis in these organisms. Nitrobacter and Nitrospira are believed to be the most common genera among the – NO2 -oxidizing autotrophs. At least some of the Nitrobacter species also can denitrify under anoxic conditions, by reducing NO3– to NO2–. In addition, species of Nitrobacter are capable of growing mixotrophically, using organic carbon as a carbon source and NO2– as an energy source and yielding higher growth rates than when they grow lithoautotrophically. Even heterotrophic metabolism is possible under both oxic and anoxic conditions. Such growth is slower than lithoautotrophic growth (Bock, 1976), but can yield higher cell densities (Schmidt et al., 2003). The importance of heterotrophic metabolism by the nitrifying bacteria for fixed-growth systems has not been established, but the rates of nitrification reported in biological treatment are considerably greater than those reported for Nitrobacter growing heterotrophically or mixotrophically in pure cultures. It is entirely feasible that Nitrobacter uses a range of different strategies in biofilms. Nitrification generally is modeled using the modified Monod expressions.
8.6 Anoxic Nitrification/Denitrification Nitrification with concomitant denitrification is carried out under anoxic conditions by bacteria unknown to exist until the late 1980s. The bacteria belong to several genera
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of the order Planctomycetales and are referred to as anaerobic ammonium oxidizers, or anammox bacteria, for short. They have been found in over 30 freshwater and marine environments around the world (Op den Camp et al., 2006) Anammox bacteria use NH4+ as an energy source and CO2 as a carbon source; thus, the process is autotrophic. Both NO2– and NO3– can be used as electron acceptors, with N2 as a product. Anammox bacteria are unable to use oxygen and do not function at oxygen concentrations above approximately 0.5 g/m3. Growth rates and yields of anammox bacteria are very low, even compared with aerobic nitrifiers. However, the saturation (affinity) coefficients, KSNH4+, are less than 0.1 g/m3, resulting in operation at maximum rates at very low NH4+ concentrations. Anammox bacteria are believed to be responsible for over 50% of the nitrogen transformations in the oceans and for most of the N2 in the Earth’s atmosphere. The late discovery of the anammox bacteria is explained partially by the fact that the organisms do not grow under conventional laboratory culturing conditions. Nitrogen transformations that could not be explained by the concentrations of known denitrifiers led to the search for unidentified groups using molecular biological techniques. Anoxic nitrification/denitrification can be carried out in wastewater treatment, in conjunction with organic oxidation, by providing aerobic and anoxic zones spatially (plug-flow) or temporally (sequencing batch operation) (Strous et al., 1997). Aerobic nitrification occurs in the aerobic zones, providing the NO2– necessary for anammox in the anoxic zones. A stoichiometric expression for the anammox nitrification/denitrification process is given below (Strous et al., 1998).
NH +4 + 1.32 NO 2− + 0.066 HCO3− + 0.13 H + →
1.02 N 2 + 0.26 NO3− + 0.066 CH 2 O0.5 N 0.15 + 2.03 H 2 0
(2.28)
The term CH2O0.5N0.15 is an empirical formula for cell products. Note that the predicted cell yield is 0.11 g cells/g N oxidized. Domestic wastewaters have relatively low concentrations of ammonium and organic nitrogen, and maintaining an anammox population in a conventional treatment process is difficult. Application of anammox to industrial wastewaters containing high concentrations of ammonia or urine streams separated from domestic wastewater is promising.
8.7 Biological Phosphorus Removal Biological phosphorus removal provides an excellent example of the relationships that can be developed in microbial consortia treating wastewater. Several uncultured
Biology of Fixed-Growth Process
heterotrophic bacteria are capable of producing storage materials in the form of polyphosphates when volatile fatty acids (VFAs) are present in the feed and maintaining 2 to 5 times higher levels of phosphorus in the cell than typically are present in the biomass. Candidatus Accumulibacter phosphatis or Rhodocyclus-related bacteria have been found in several enriched laboratory-scale reactors and also in full-scale plants (Oehmen et al., 2007). Other organisms appear to be involved in enhanced biological phosphorus removal (EBPR), including Actinobacteria (Kong et al., 2005). Polyphosphateaccumulating organisms use this material as an energy source when exogenous terminal electron acceptors are unavailable. Polyphosphate use does not occur unless the fluid is devoid of oxygen, nitrate, and nitrite. The presence of sulfate probably makes little difference, because sulfate-reducing organisms do not store polyphosphate and do not seem to compete with polyphosphate-storing bacteria under the conditions that occur in biological treatment. When full anaerobic conditions exist and bacteria with stored polyphosphate are present, the stored polyphosphate is hydrolyzed, and the energy released is used by the cells to store organic material, principally as PHB (Chiesa, 1982). The orthophosphate produced is released into the liquid. If the cells are then transferred into an aerobic or anoxic environment, the stored PHB is used as an energy and carbon source for growth and polyphosphate storage. Because a net increase in cell mass occurs as a result of the process, a net uptake of phosphorus also occurs. Removal of phosphorus from the system then can be achieved by wasting the cells produced. A competing group of organisms, the glycogen-accumulating organisms, also is capable of taking up VFA under anaerobic conditions, but without accumulating polyphosphate under aerobic conditions. The EBPR is used worldwide, and there are several processes that successfully have integrated EBPR (Oehmen at al., 2007).
8.8 Sulfide and Sulfur Oxidation Chemolithoautotrophic sulfur-oxidizing bacteria exist in many different groups of prokaryotes and include acidophiles (optimum pH <6) and neutrophiles (optimum pH 7 to 8). Acidophiles play an important role in acidic metal leaching processes, and neutrophiles tend to dominate in wastewater treatment processes. In addition, there are sulfur-oxidizing bacteria, which thrive at pH values above 9 (Sorokin et al., 2006).
8.9 Hydrogen Oxidation Hydrogen is a source of energy, and reductant and hydrogen-oxidizing bacteria are a fairly diverse group (Schwartz and Friedrich, 2006). Nearly all of the hydrogen
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oxidizers also are able to metabolize simple organic compounds, and their taxonomic characteristics would place them in a variety of chemoheterotrophic genera. Molecular hydrogen is not a common product in aerobic environments; therefore, the hydrogen oxidizers are not major players in aerobic biological treatment. Autotrophic denitrification of drinking water can be accomplished with biofilms using hydrogen as an electron donor, either NO3– or NO2– as an electron acceptor, and CO2 as a carbon source. Hydrogen also is a key component of anaerobic metabolism in methanegenerating Archaea called methanogens and in anoxygenic phototrophs. Finally, hydrogen production allows cells to release excess reductant produced during fermentative metabolism.
9.0 FEATURES OF MICROBIAL COMMUNITIES IN BIOFILMS Bacteria most often are studied in pure cultures, but biological treatment processes use bacterial communities where species work together. This situation presents problems in translation of information gathered in bacterial research into applications in biological treatment problems, because it must be assumed that individual species behave differently as members of a consortium than in pure cultures. For example, the micro-environment surrounding a bacterium in a biofilm is affected by the activities of neighboring cells. The local oxygen concentration or tension may be low, although the average concentration in the medium is relatively high. In many cases, low oxygen concentrations result in the production of partially oxidized metabolites, such as organic acids, which are used by other species of bacteria or allow establishment of micro-aerophilic species, which would not be competitive at a higher oxygen concentration. Similar situations occur for local pH concentration effects. A number of important biological treatment processes are dependent on the activities of bacterial consortia consisting of different species or strains. Among these processes are nitrification and phosphorus removal (where acid-fermenting organisms provide substrate for the polyphosphate-storing organisms) and methane fermentation (in which acid and hydrogen fermenters produce substrate for the methane fermenters, and the methane fermenters remove compounds that would inhibit growth of the acid fermenters). In addition, many organic removal processes involve multiple bacterial strains or species where each may only be able to execute certain enzymatic steps in the degradation or transformation of a specific substrate.
Biology of Fixed-Growth Process
Microbial communities in fixed films (biofilms) and suspended systems used to be considered a black box, with respect to process engineering. In this line of thinking, there is sufficient microbial functional redundancy in a given community to lead to stable process performance (Briones and Raskin, 2003; Curtis and Sloan, 2006; Rittman and McCarty, 2001), such as COD removal, nitrification, and others. Curtis et al. (2006) further illustrated this concept by relating functional redundancy directly to diversity, where a more specialized function (e.g., nitrification) involves a less diverse pool of microbes or metacommunity than that associated with a less specialized task (e.g., simple carbohydrate breakdown). However, a long-term study revealed that stable biological treatment performance does not necessarily correlate to stable community dynamics (Fernandez et al., 1999). The relationship between microbial community structure and performance characteristics in fixed-film systems is not well-understood (Wuertz et al., 2004). What is known is that microbial community dynamics are reproducible in laboratory-scale reactors for at least 3 months, when operational parameters are kept unchanged (Falk et al., 2009). Advances in molecular microbial ecology and environmental biotechnology have opened the door for a more systems-based approach to the study of microbial consortia (Raes and Bork, 2008). It is known that many bacterial species can communicate via synthesis, release, and detection of diffusible signal molecules. They do so to monitor their own population density, in a process referred to as quorum sensing (Ryan and Dow, 2008), and thus modulate the level of expression of subsets of genes, such as those controlling polymer secretion (Nadell et al., 2008). Using pure culture biofilms, it has been shown that quorum sensing can influence the development and physical structure of biofilms. Bacteria also are capable of eavesdropping on signaling that occurs in their immediate environment. Such interspecies signaling is likely to play a role in fixed-film applications. Microbial consortia now can be studied using metagenomic approaches, even when it is not possible to cultivate individual members (Steele and Streit, 2005). Metagenomics refers to the study of mixed microbial populations from a specific habitat at the DNA level (Riesenfeld et al., 2004) and is made possible by the ability of researchers to clone large DNA fragments from an environmental sample and rapidly sequence the DNA. Increased knowledge about metabolic capacities at the genetic level, about the way intracellular communication organizes function in microbial consortia, and about natural fluctuations in microbial communities is expected to open the door to engineering microbial consortia for the optimization of fixed-growth
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processes. In the future, biofilms may be engineered as synthetic consortia, with specific metabolic properties, and monitored with high-throughput molecular screening methods using inexpensive gene-chip-based technology.
10.0 REFERENCES Alvarez-Cohen, L.; McCarty, P. L. (1991) Effects of Toxicity, Aeration, and Reductant Supply on Trichloroethylene Transformation by a Mixed Methanotrophic Culture. Appl. Environ. Microbiol., 57, 228–235. Atkinson, B. (1974) Biochemical Reaction Engineering; Pion Press: London, United Kingdom. Aust, S.; Fernando, T.; Brock, B.; Tuisel, H.; Bumpus, J. (1988) Biological Treatment of Hazardous Wastes by Phanerochaete Chrysoporium. Proceedings of the Conference on Biotechnology Applications in Hazardous Waste Treatment, Longboat Key Florida, Oct 30–Nov 4, Lewandowski, G., Baltzis, B., Armenante, P. (Eds.); Engineering Foundation: New York. Bock, E. (1976) Growth of Nitrobacter in Presence of Organic Matter: 2. Chemoorganic Growth of Nitrobacter agilis. Arch. Microbiol., 108, 305–312. Briones, A.; Raskin, L. (2003) Diversity and Dynamics of Microbial Communities in Engineered Environments and Their Implications for Process Stability. Curr. Opin. Biotechnol., 14 (3), 270–276. Chiesa, S. C. (1982) Growth and Control of Filamentous Microbes in Activated Sludge. Ph.D. Dissertation, University of Notre Dame, Notre Dame, Indiana. Curtis, T. P.; Head, I. M.; Lunn, M.; Woodcock, S.; Schloss, P. D.; Sloan, W. T. (2006) What is the Extent of Prokaryotic Diversity? Phil. Trans. R. Soc. B, 361, 2023–2037. Curtis, T. P.; Sloan, W. T. (2006) Towards the Design of Diversity: Stochastic Models for Community Assembly in Wastewater Treatment Plants. Water Sci. Technol., 54 (1), 227–236. de Beer, D.; Stoudley, P.; Roe, F.; Lewandowski, Z. (2004) Effects of Biofilm Structure on Oxygen Distribution and Mass Transport. Biotechnol. Bioeng., 43, 1132–1138.
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Dobbs, R. A.; Wang, L. P.; Govind, R. (1989) Sorption of Toxic Organic Compounds on Waste Water Solids—Correlation with Fundamental Properties. Environ. Sci. Technol., 23, 1092–1097. Enright, A. M.; Collins, G.; O’Flaherty, V. (2007) Low-Temperature Anaerobic Biological Treatment of Toluene-Containing Wastewater. Water Res.: 41, 1465–1472. Evans, W. C.; Fuchs, G. (1988) Anaerobic Degradation of Aromatic Compounds. Annual Review of Microbiology, Ornston, L. N., Balows, A., Baumann, P. (Eds.); Annual Reviews, Inc.: Palo Alto, California. Falk, M. W.; Song, K. G.; Matiasek, M. G.; Wuertz, S. (2009) Microbial Community Dynamics in Replicate Membrane Bioreactors—Natural Reproducible Fluctuations. Water Res., 43, 842–852. Fan, S.; Scow, K. M. (1993) Biodegradation of Trichloroethylene and Toluene by Indigenous Microbial Populations in Soil. Appl. Environ. Micriobiol., 59, 1911–1918. Fernandez, A.; Huang, S. Y.; Seston, S.; Xing, J.; Hickey, R.; Criddle, C.; Tiedje, J. (1999) How Stable is Stable? Function Versus Community Composition. Appl. Environ. Microbiol., 65 (8), 3697–3704. Flegal, T. M.; Schroeder, E. D. (1976) Temperature Effects on BOD Stoichiometry and Oxygen Uptake Rate. J. Water Pollut. Control Fed., 49, 2700–2707. Fuhs, G. W.; Chen, M. (1975) Microbiological Basis of Phosphate Removal in the Activated Sludge Process for the Treatment of Wastewater. Microb. Ecol., 2, 119–138. Grady, C. P. L., Jr.; Daigger, G. T.; Lim, H. C. (1999) Biological Wastewater Treatment, 2nd ed.; Marcel Dekker, Inc.: New York. Hackett, W. F.; Connors, W. J.; Kirk, T. K.; Zeikus, J. G. (1977) Microbial Decomposition of Synthetic 14C-Labeled Lignins in Nature: Lignin Biodegradation in a Variety of Natural Materials. Appl. Environ. Microbiol., 33, 43–51. Hall-Stoodley, L.; Costerton, J. W.; Stoodley, P. (2004) Bacterial Biofilms: From the Natural Environment to Infectious Diseases. Nat. Rev. Microbiol., 2, 95–108.
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Henze, M.; Gujer, W.; Mino, T.; Matsuo, T.; Wentzel, M. C.; Marais, G. v. R. (1995) Activated Sludge Model No. 2, IAWQ Scientific and Technical Reports, No. 3; International Association for Water Quality: London, United Kingdom. Hyman, M. R.; Murton, I. B.; Arp, D. J. (1988) Interaction of Ammonia Monooxygenase from Nitrosomonas europea with Alkanes, Alkenes, and Alkynes. Appl. Environ. Microbiol., 54, 3187–3188. Ingraham, J. L.; Maaloe, O.; Neidhardt, F. C. (1983) Growth of the Bacterial Cell; Sinauer and Associates, Inc.: Sunderland, Massachusetts. Kehrberger, G. J.; Norman, J. D.; Schroeder, E. D.; Busch, A. W. (1964) BOD Progression in Soluble Substrates. VII. Temperature Effects. Proceedings of the 19th Purdue Industrial Waste Conference, West Lafayette, Indiana, May 5–7; Purdue University: West Lafayette, Indiana. Kong, Y. H.; Nielsen, J. L.; Nielsen, P. H. (2005) Identity and Ecophysiology of Uncultured Actinobacterial Polyphosphate-Accumulating Organisms in Full-Scale Enhanced Biological Phosphorus Removal Plants. Appl. Environ. Microbiol., 71, 4076–4085. Lengeler, J. W. ; Drews, G. ; Schlegel, H. G. (Eds.) Biology of the Prokaryotes (1999) Thieme Verlag: Stuttgart, Germany. Levin, G. V.; Shapiro, J. (1967) Metabolic Uptake of Phosphorus by Wastewater Organisms. J. Water Pollut. Control Fed., 37, 800–821. Lewandowski, Z.; Beyenal, H. (2003) Mass Transport in Heterogeneous Biofilms. In Biofilms in Wastewater Treatment: An Interdisciplinary Approach, Wuertz, S., Bishop, P. L., Wilderer, P. A. (Eds.); IWA Press: London, United Kingdom. Lipschultz, F.; Zafiriou, O. C.; Wofsky, S. C.; McElroy, M. B.; Valois, F.W.; Watson, S. W. (1981) Production of NO and N2O by Soil Nitrifying Bacteria. Nature, 294, 641–643. Logan, B. E. (1993) Oxygen Transfer in Trickling Filters. J. Environ. Eng., 119, 1059–1076. MacDonald, R. M. (1978) Population Dynamics of the Nitrifying Bacterium Nitrosolubus in Soil. J. Appl. Ecol., 16, 529–535. Madigan, M. T.; Martinko, J. M.; Parker, J. (2003) Brock Biology of Microorganisms, 9th ed.; Prentice-Hall: Englewood Cliffs, New Jersey.
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Mattick, J. S. (2002) Type IV Pili and Twitching Motility. Annu. Rev, Microbiol., 56, 289–314. McCarty, P. L. (1965) Thermodynamics of Biological Synthesis and Growth. Proceedings of the 2nd International Conference on Water Pollution Research; New York, 169. Metcalf and Eddy, Inc. (2003) Wastewater Engineering: Treatment and Reuse, Tchobanoglous, G., Burton, F. L., Stensel, H. D. (Eds.); McGraw-Hill: New York. Nadell, C. D.; Xavier, J. B.; Levin, S. A.; Foster, K. R. (2008) The Evolution of Quorum Sensing in Bacterial Biofilms. PLoS Biol., 6, 0171–0179. Neidhardt, F. C.; Ingraham, J. L.; Schaechter, M. (1990) Physiology of the Bacterial Cell: A Molecular Approach; Sinauer Associates, Inc.: Sunderland, Massachusetts. Nielsen. P. H.; Jahn, A. (1999) Extraction of EPS. In Microbial Extracellular Polymeric Substances, Wingender, J., Neu, T. R., Flemming, H.-C. (Eds.); Springer: Berlin, Germany, 49–72. Oehmen, A.; Lemos, P. C.; Carvalho, G.; Yuan, Z. G.; Keller, J.; Blackall, L. L.; Reis, M. A. M. (2007) Advances in Enhanced Biological Phosphorus Removal: From Micro to Macro Scale. Water Res., 41, 2271–2300. Op den Camp, H. J. M.; Kartal, B.; Guven, D.; van Niftrik, L.; Haaijer, S. C. M.; van der Star, W. R. L.; van de Pas-Schoonen, K. T.; Cabezas, A.; Ying, Z.; Schmid, M. C. (2006) Global Impact and Application of the Anaerobic AmmoniumOxidizing (Anammox) Bacteria. Biochem. Soc. Trans., 34, 174–178. Porges, N.; Jaiswicz, L.; Hoover, S. R. (1953) Biological Oxidation of Dairy Waste, VII. Proceedings of the 24th Purdue Industrial Waste Conference, West Lafayette, Indiana, May 6–8; Purdue University: West Lafayette, Indiana. Raes, J.; Bork, P. (2008) Molecular Eco-Systems Biology: Towards an Understanding of Community Function. Nat. Rev. Microbiol., 6, 693–699. Rajal, V. B.; McSwain, B. S.; Thompson, D. E.; Leutenegger, C. M.; Kildare, B.; Wuertz, S. (2007) Validation of Hollow Fiber Ultrafiltration and Real Time PCR Using Bacteriophage PP7 as Surrogate for the Quantification of Viruses from Water Samples. Water Res., 41, 1411–1422.
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Rajal, V. B.; McSwain, B. S.; Thompson, D. E.; Leutenegger, C. M.; Wuertz, S. (2007) Molecular Quantitative Analysis of Human Viruses in California Storm Water. Water Res., 41, 4287–4298. Reineke, W.; Knackmuss, H. J. (1988) Microbial Degradation of Haloaromatics. Annu. Rev. Microbiol., 42, 263–287. Riesenfeld, C. S.; Schloss, P. D.; Handelsman, J. (2004) Metagenomics: Genomic Analysis of Microbial Communities. Ann. Rev. Genet., 38, 525–552. Rittman, B. E.; McCarty, P. L. (2001) Environmental Biotechnology: Principles and Applications; McGraw-Hill: New York. Ryan, R. P.; Dow, J. M. (2008) Diffusible Signals and Interspecies Communication in Bacteria. Microbiology, 154, 1845–1858. Schmidt, I.; Sliekers, O.; Schmid, M.; Bock, E.; Fuerst, J.; Kuenen, J. G.; Jetten, M. S. M.; Strous, M. (2003) New Concepts of Microbial Treatment Processes for the Nitrogen Removal in Wastewater. FEMS Microbiol. Rev., 27, 481–492. Schroeder, E. D. (2002) Trends in Application of Gas-Phase Bioreactors. Rev. Environ Sci. Biotechnol., 1, 65–74. Schroeder, E. D.; Eweis, J. B.; Chang, D. P. Y.; Veir, J. K. (2000) Biodegradation of Recalcitrant Compounds. Water Air Soil Pollut., 123, 133–146. Schroeder, E. D.; Tchobanoglous, G. (1976) Mass Transfer Limitations in Trickling Filters. J. Water Pollut. Control Fed., 48, 771–775. Schroeder, E. D.; Wuertz, S. (2003) Bacteria. In The Handbook of Water and Wastewater Microbiology, Mara, D., Horan, N. (Eds.); Elsevier: Amsterdam, Netherlands, 57–68. Serafim, L. S.; Lemos, P. C.; Levantesi, C.; Tandoi V.; Santos H.; Reis M. A. M. (2002) Methods for Detection and Visualization of Intracellular Polymers Stored by Polyphosphate-Accumulating Microorganisms. J. Microbiol. Meth., 51, 1–18. Schwartz, E.; Friedrich, B. (2006) The H 2-Metabolizing Prokaryotes. In The Prokaryotes: An Evolving Electronic Resource for the Microbiological Community, 3rd ed.; Springer: New York, 496–563. Shapiro, J.; Levin, G. V.; Zea, H. G. (1967) Anoxically Induced Release of Phosphate in Wastewater Treatment. J. Water Pollut. Control Fed., 39, 1810–1818.
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Singer, P. C.; Stumm, W. (1970) The Solubility of Ferrous Iron in CarbonateBearing Waters. J. Am. Water Works Assoc., 62, 198–202. Skraber, S.; Helmi, K.; Ferreol, M.; Gantzer, C.; Hoffmann, L.; Cauchie, H. M. (2007) Occurrence and Persistance of Bacterial and Viral Faecal Indicators in Wastewater Biofilms. Water Sci. Technol., 55 (8–9), 377–385. Sorokin, D. Y.; Banciu, H.; Robertson, L. A.; Kuenen, J. G. (2006) Haloalkaliphilic Sulfur-Oxidizing Bacteria. In The Prokaryotes: An Evolving Electronic Resource for the Microbiological Community, 3rd ed.; Springer: New York, 969–984. Spaeth, R.; Wuertz, S. (2000). Extraction and Quantification of Extracellular Polymeric Substances from Wastewaters. In Biofilms. Investigative Methods & Applications, Flemming, H.-C., Szewzyk, U., Griebe. T. (Eds.), Technomic Publishers: Lancaster, Pennsylvania, 51–68. Stanier, R. Y.; Ingraham, J. L.; Wheelis, M. L.; Painter, P. R. (1986) The Microbial World; Prentice-Hall: Englewood Cliffs, New Jersey. Steele, H. L.; Streit, W. R. (2005) Metagenomics: Advances in Ecology and Biotechnology. FEMS Microbiol. Lett., 247, 105–111. Stephen, J. R.; McCaig, A. E.; Smith, Z.; Prosser, J. L.; Embley, T. M. (1996) Molecular Diversity of Soil and Marine 16S rRNA Gene Sequences Related to β-Subgroup Ammonia-Oxidizing Bacteria. Appl. Environ. Microbiol., 62, 4147–4154. Steward, G. F.; Smith, D. C.; Azam, F. (1996) Abundance and Production of Bacteria and Viruses in the Bering and Chukchi Seas. Mar. Ecol. Prog. Ser., 131, 287–300. Strous, M.; Heijnen, J. J.; Kuenen, J. G.; Jetten, M. S. M. (1998) The Sequencing Batch Reactor as a Powerful Tool for the Study of Slowly Growing Anaerobic Ammonium-Oxidizing Microorganisms. Appl. Microbiol. Biotechnol., 50, 589–596. Strous, M.; Van Gerven, E.; Ping, Z.; Kuenen, J. G.; Jetten, M. S. M. (1997) Ammonium Removal from Concentrated Waste Streams with the Anaerobic Ammonium Oxidation (Anammox) Process in Different Reactor Configurations. Water Res., 31, 1955–1962. Swilley, E. L.; Bryant, J. O.; Busch, A. W. (1964) The Significance of Transport Phenomena in Biological Oxidation Processes. Proceedings of the 19th Purdue
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Industrial Waste Conference, West Lafayette, Indiana, May 5–7; Purdue University: West Lafayette, Indiana. Szewzyk, U.; Szewzyk, R. ; Manz, W. ; Schleifer, K.-H. (2000) Microbiological Safety of Drinking Water. Annu. Rev. Microbiol., 54, 81–127. Venugopalan, V. P.; Kuehn, M.; Hausner, M.; Springael, D.; Wilderer, P. A.; Wuertz, S. (2005) Architecture of a Nascent Sphingomonas sp. Biofilm under Varied Hydrodynamic Conditions. Appl. Environ. Microbiol., 71, 2677–2686. Wilderer, P. A.; Bungartz, H.-J.; Lemmer, H.; Wagner, M.; Keller, J.; Wuertz, S. (2002) Modern Scientific Methods and Their Potential in Wastewater Science and Technology. Water Res., 36, 370–393. Williamson, K.; McCarty, P. L. (1976) A Model of Substrate Uptake by Bacterial Films. J. Water Pollut. Control Fed., 48, 9–24. Woertz, J. R.; Kinney, K.A.; McIntosh, N. D. P.; Szaniszlo, P. J. (2001) Removal of Toluene in a Vapor-Phase Bioreactor Containing a Strain of the Dimorphic Black Yeast Exophiala lecanii-corni. Biotechnol. Bioeng., 75, 550–558. Woertz, J. R.; Kinney, K. A.; Szaniszlo, P. J. (2001) A Fungal Vapor-Phase Bioreactor for the Removal of Nitric Oxide from Waste Gas Streams. J. Air Waste Manag. Assoc., 51, 895–902. Wuertz, S.; Okabe, S.; Hausner, M. (2004) Microbial Communities and Their Interactions in Biofilm Systems: An Overview. Water Sci. Technol., 49 (11–12), 327–336. Yu, S.; Semprini, L. (2004) Kinetics and Modeling of Reductive Dechlorination at High PCE and TCE Concentrations. Biotechnol. Bioeng., 88, 451–464.
Chapter 3
Trickling Filter and Combined Trickling Filter Suspended-Growth Process Design and Operation 1.0 INTRODUCTION
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2.0 General Description
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4.0 Ventilation and Air Supply Alternatives
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2.1 Distribution System
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4.1 Natural Draft
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2.2 Biofilm Carriers
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4.2 Mechanical Ventilation
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2.3 Containment Structure 90 2.4 Underdrain System and Ventilation
5.0 Trickling Filter Process Models 103
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2.5 Trickling Filter Pumping Stations: Influent and Recirculation 91
5.1 National Research Council
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2.6 Hydraulic and Contaminant Loading
5.2 Galler and Gotaas
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5.3 Kincannon and Stover 106
3.0 Process Flow Sheets and Bioreactor Configuration 3.1 Standard Process Flow Diagrams
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5.4 Velz
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5.5 Schulze
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5.6 Germain
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5.7 Eckenfelder
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5.8 Chartered Institution of Water and Environmental Management 110
3.2 Bioreactor Classification 96 3.3 Hydraulic Application: Effect on Media Wetting, Flow Distribution, and Control 98
5.9 Logan Trickling Filter Model 81
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(continued)
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5.10 Selecting a Trickling Filter Model 112 5.11 Method for Combining Trickling Filter and Suspended-Growth Models 113 6.0 Process Design
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6.1 Combined Carbon Oxidation and Nitrification
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6.2 Nitrifying Trickling Filters
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6.2.1 Gujer and Boller Nitrifying Trickling Filter Model 122 6.2.2 Okey and Albertson Nitrifying Trickling Filter Model 124 6.2.2.1 Application of the Gujer and Boller Model 126 6.2.2.2 Application of the Albertson and Okey Model 128 6.3 Temperature and Hydraulic Application Effects 131 7.0 Design Considerations 132 7.1 Distribution System
133
7.1.1 Hydraulic Drive Rotary Distributors 134
7.1.2 Electronic or Mechanical Drive Rotary Distributors 136 7.1.3 Optimizing Rotary Distributor Operation 138 7.2 Construction of Rotary Distributors 138 7.3 Trickling Filter Media Selection 139 7.3.1 Depth 141 7.3.2 Structural Integrity 142 7.4 Trickling Filter Pumping Station or Dosing Siphon 144 7.5 Control Mechanisms for Trickling Filter Macro Fauna 144 7.5.1 Operational Strategies and Facility Improvements for Macro Fauna Control 145 7.5.2 Spülkraft 147 7.5.3 Flooding 148 7.5.4 Chemical Treatment 149 7.5.5 Physical Control 152 7.6 Trickling Filter Startup 153 7.7 Combined Trickling Filter and Suspended-Growth Processes 155 7.7.1 Activated Biofilter 156 7.7.2 Trickling Filter/Solids Contact 158 (continued)
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7.7.3 Roughing Filter/ Activated Sludge 162 7.7.4 Biofilter/Activated Sludge 162
7.7.5 Trickling Filter/ Activated Sludge 8.0 REFERENCES
162 163
1.0 INTRODUCTION Until the 1950s, trickling filter (TF) design protocol was scattered and empirical in nature. Then, during the 1950s and 1960s, the Dow Chemical Company (Midland, Michigan) began experimentation with modular plastic media (Bryan, 1955). Numerous trickling filter process studies were conducted during the same period (Eckenfelder, 1961; Galler and Gotaas, 1964; Germain, 1966; Schulze, 1960), which led to the development of generally accepted design criteria. After the U.S. Environmental Protection Agency (Washington, D.C.) issued its definition of secondary treatment standards in the early 1970s, the trickling filter process was regarded as being unable to consistently produce effluent water quality that met the published standards, in part, as a result of poor secondary clarifier design (Parker, 1999). Norris et al. (1982) described development of the trickling filter/solids contact process (TF/ SC) in response. The first full-scale TF/SC process included a rock-media trickling filter followed by a small aeration basin (receiving return sludge) and flocculator clarifier. The researchers demonstrated that wastewater treatment plant (WWTP) effluent water quality could be improved greatly by bioflocculation in the solids contact basin and improved secondary clarifier design. Combined trickling filter/suspended-growth (TF/SG) processes preceding the TF/SC process were designed with the suspended-growth reactor (SGR) primarily for oxidation. The modern trickling filter typically includes the following major components:
(1) Rotary distributors, with speed control; (2) Modular plastic biofilm carriers, or media (typically cross-flow [XF] media, unless the bioreactor is treating high-strength wastewater, which warrants the use of vertical-flow [VF] media); (3) A mechanical aeration system (which consists of air distribution piping and low-pressure fans); (4) Trickling filter effluent recirculation pumping station; and (5) Covers that aid with the uniform distribution of process air and foul air containment (for odor control).
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Covers may be equipped with sprinklers that discharge in-plant recycle to cool plastic media during emergency shutdown. Trickling filter mechanics are poorly understood. Consequently, there is a general lack of mechanistic mathematical models and design approaches, and the design and operation of trickling filter and TF/SG processes is empirical or semi-empirical. Benefits inherent to the trickling filter process (when compared with activated sludge processes) include operational simplicity, resistance to toxic and shock loads, and low energy requirements. However, trickling filters are susceptible to nuisance conditions that are primarily caused by macrofauna. Process mechanical components dedicated to minimizing the accumulation of macrofauna, such as filter flies, worms, and snail (shells), are now standard. Unfortunately, information on the selection and design of these process components is fragmented and has been poorly documented. The TF/SC process is the most common TF/SG process. This chapter describes state-of-the art trickling filter and TF/SC process design and operation.
2.0 General Description The trickling filter is a three-phase biofilm reactor, with fixed carriers. Wastewater enters the bioreactor through a distribution system, trickles downward over the biofilm surface, and air moves upward or downward in the third phase, where it diffuses through the flowing liquid and into the biofilm. Trickling filter components generally include an influent water distribution system, containment structure, rock or plastic media, underdrain, and ventilation system. Figure 3.1 illustrates a trickling filter crosssection and these typical bioreactor components. Wastewater treatment using the trickling filter results in a net production of total suspended solids (TSS). Therefore, liquid–solid separation is required and typically is achieved with circular or rectangular secondary clarifiers. The trickling filter process generally includes an influent/ recirculation pumping station, trickling filter, and liquid–solid separation unit.
2.1 Distribution System Primary effluent (or screened and degritted wastewater—fine screens with a maximum 3-mm bar spacing or perforated plate orifice diameter—should be applied for preliminary treatment, if primary clarification is excluded from the process flow sheet) is either pumped or flows by gravity to the trickling filter distribution system. Essentially, there are two types of trickling filter distribution systems—fixed-nozzle and rotary distributors. Because their efficiency is poor, distribution with fixed nozzles should not be used (Harrison and Timpany, 1988). Rotary distributors may be
ROTARY DISTRIBUTOR DISTRIBUTOR ARM CONTAINMENT STRUCTURE
BIOFILM SUPPORT MEDIUM
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FRP GRATING UNDERDRAIN
AIR DISTRIBUTION HEADER
INFLUENT
Figure 3.1 Typical trickling filter components and cross-section.
EFFLUENT
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driven hydraulically or electrically. A properly designed rotary distribution system allows for effective media-wetting and the intermittent application of wastewater to biofilm carriers. The intermittent application of influent wastewater (i.e., typically primary effluent diluted with recirculated trickling filter effluent) allows the biofilm to have periods of “resting,” which primarily serves as a process aeration mechanism. Poor media-wetting may lead to dry pockets, ineffective treatment zones, and odor. Hydraulically driven rotary distributors have discharge orifices along the front and back of each arm. The distributor is driven forward by energy imparted by the influent stream. To the extent possible, distributor rotational speed is controlled by flow passing through the orifices, which are situated along the distributor arm face that points to the direction of rotation. These back-spray orifices, or reverse-thrusting jets, slow rotational speed, but the designer must determine the extent of speed control based on site-specific operating conditions and process hydraulics. An electrically driven or modern hydraulically driven rotary distributor may be required if reverse-thrusting jets are incapable of achieving the degree of speed control required to maintain the desired trickling filter dosing rate(s). A modern hydraulically driven rotary distributor is depicted in Figure 3.2. This system uses gates (controlled by variable-frequency drive [VFD]) to open or close distributor orifices and adjust
Figure 3.2 (Left) Hydraulically driven rotary distributors use variable frequency drive controlled gates that either open or close distributor orifices, which adjust with varying pumped flowrates to maintain a constant preset rotational speed; and (right) electrically driven rotary distributor (Photographs courtesy WesTech, Inc., Salt Lake City, Utah).
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
rotational speed. Figure 3.2 also depicts an electrically driven rotary distributor. The use of a VFD allows for the more precise control of distributor-arm location when stopping rotation. Electrically driven rotary distributors have motorized drive units that control distributor speed independent of the wastewater pumped flow.
2.2 Biofilm Carriers Ideal trickling filter media provides a high specific surface area, low cost, high durability, and porosity that is high enough to avoid clogging and promote ventilation (Metcalf & Eddy, 2003). Trickling filter media types include rock, random (synthetic), vertical-flow (synthetic), and cross-flow (synthetic). Both vertical- and cross-flow media are constructed with smooth and/or corrugated plastic sheets. Another commercially available synthetic media, although not commonly used, is vertically hanging plastic strips. Horizontal redwood or treated wooden slats also have been used, but generally are no longer considered viable because of high cost or limited supply. Modules of plastic sheets (i.e., self-supporting vertical- or cross-flow modules) are used almost exclusively for new and improved trickling filters, but several trickling filters with rock media exist and have proven capable of meeting treatment objectives, when properly designed and operated. Table 3.1 compares the characteristics of some trickling filter media types. The higher specific surface area and void space in modular synthetic media allows for higher hydraulic loading, enhanced oxygen transfer, and biofilm thickness control, compared with rock media. Rock media ideally has a 50-mm diameter, but may range in size. Rounded (river) rock helps mitigate some issues associated with rigid rock (slag) media. The slag rock contains crevices that can retain water and accumulate biomass. As a result of structural requirements associated with the large unit weight of the rock media, rock media are shallow compared with synthetic-media trickling filters and are more susceptible to excessive cooling. Trickling filter performance aside, the excessive cooling can subject media to freeze–thaw cycles. Water retained inside slag rock crevices may expand and sever rock fragments. This can result in fine material accumulation which, with retained biomass, is a primary contributor to rock-media trickling filter clogging (Grady et al., 1999). Generally, rock media is considered to have a low specific surface area, void space, and high unit weight. Although recirculation is common, the low void ratio in rock-media trickling filters limits hydraulic application rates. Excessive hydraulic application can result in ponding, limited oxygen transfer, and poor bioreactor performance. Performance of existing rock-media trickling filters sometimes may be improved by providing mechanical ventilation, solids contact channels, and/or deepened secondary clarifiers, which include energy-dissipating inlets
87
Table 3.1 Properties of some trickling filter media. Nominal size (m [ft])
Bulk density (kg/m3 [lb/cu ft])a
Specific surface area (m2/m3 [sq ft/cu ft])b
River
0.024 to 0.076 (0.08 to 0.25)
1442 (90)
62 (19)
50
Slag
0.076 to 0.128 (0.25 to 0.42)
1600 (100)
46 (14)
60
0.61 × 0.61 × 1.22 (2 × 2 × 4)
24 to 45 (1.5 to 2.8)
100, 138, and 223 (30, 42, and 68)
95
0.61 × 0.61 × 1.22 (2 × 2 × 4)
24 to 45 (1.5 to 2.8)
102 and 131 (31 and 40)
95
Media type
Material
Void space (%)
Rock
Plasticc
Cross-flow PVC
Vertical flow PVC
0.185 ø × 0.051 H (7.3 in. ø × 2-in. H)
Randomd
27 (1.7)
98 (30)
95
Polypropylene a
lb/cu ft × 16.02 = kg/m3.
b
sq ft/cu ft × 3.281 = m2/m3.
Manufacturers of modular plastic media: (formerly) BF Goodrich (Greenville, South Carolina), American Surfpac Corporation (Downingtown, Pennsylvania), NSW (Nordenham, Germany), Munters (Aachen, Germany), (currently) Brentwood Industries (Reading, Pennsylvania), Jaeger Environmental (Houston, Texas), and SPX Cooling (Overland Park, Kansas).
c
d
Manufacturers of random plastic media: (formerly) NSW and (currently) Jaeger Environmental.
88
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
and flocculator-type feed wells. Replacement of the rock media (with synthetic media) may be required in cases where rock quality is poor. If WWTP expansion is expected, existing rock-media trickling filters may be deepened if the media is replaced with modular synthetic media. Rock-media trickling filter containment structures typically are capable of providing structural support for the deeper modular synthetic media, because the rock media is substantially heavier than synthetic media. A well-designed and -operated rock-media trickling filter can provide high-quality effluent. Grady et al. (1999) suggested that, under low organic loading (i.e., less than 1 kg BOD5/m3∙d), rock- and synthetic-media trickling filters are capable of equivalent performance. However, as organic loading increases, synthetic-media trickling filters are less susceptible to operational problems and have reduced potential for plugging. Synthetic trickling filter media has a higher specific surface area and void space and lower unit weight than rock media. Because of the reduced unit weight, synthetic-media trickling filters can be constructed at depths exceeding 3 times that of a comparably sized rock-media trickling filter. Modular synthetic media generally is manufactured with the following specific surface areas: 223 m2/m3 (68 sq ft/cu ft) as high-density, 138 m2/m3 (42 sq ft/cu ft) as medium-density, and 100 m2/m3 (30 sq ft/cu ft) as low-density. Both vertical- and cross-flow media are reported to remove 5-day biochemical oxygen demand (BOD5) and ammonia-nitrogen (NH3-N) (Aryan and Johnson, 1987; Harrison and Daigger, 1987), but sufficient scientific evidence exists to surmise that there is a difference in the treatment efficiency of trickling filters constructed with cross- and vertical-flow media, even when manufactured with virtually identical specific surface areas. Plastic modules with a specific surface area in the range of 89 to 102 m2/m3 are wellsuited for carbon oxidation and combined carbon oxidation and nitrification. Parker et al. (1989) recommended medium-density cross-flow media and recommended against the use of high-density cross-flow media in nitrifying trickling filters. This is supported by observations from a pilot-scale nitrifying trickling filter application data and the conclusions of Boller and Gujer (1986) and Gujer and Boller (1984), which show lower nitrification (flux) rates for lower density modular synthetic media. The researchers claim that lower rates occur with high-density media, as a result of the development of dry spots below the flow interruption points (i.e., higher-density media has more flow interruptions and therefore less effective wetting). Using medium-density media also reduces plugging potential. Vertically oriented modular synthetic (verticalflow) media generally is accepted as being ideally suited for high-strength wastewater (perhaps industrial) and high organic loadings, such as with a roughing trickling filter.
89
90
Biofilm Reactors
In some cases, cross-flow media has been placed in the top layer to enhance wastewater distribution, and vertical-flow media comprises the remainder of the trickling filter media. Typically, the top layer of a trickling filter’s modular plastic media is covered with fiberglass-reinforced plastic (FRP) or high-density polyethylene (HDPE) grating. The grating is skid-resistant, thereby reducing the potential for slip and fall occupational hazards. In addition, the grating protects modular plastic media from deterioration by UV light and potential structural damage, which may result from water-induced load exerted during periods of high-intensity dosing.
2.3 Containment Structure Rock and random synthetic media are not self-supporting and require structural support to contain the media within the bioreactor. These containment structures are typically precast or panel-type concrete tanks. When self-supporting media, such as plastic modules, is used, other materials, such as wood, fiberglass, and coated steel have been used as containment structures. The containment structure serves to avoid wastewater splashing and provide media support, wind protection, and flood containment. In some cases, trickling filter containment structures have been designed to allow flooding of the media, which increases operator flexibility in controlling macrofauna accumulation. A complete list and description of macrofauna control alternatives is presented later in this chapter.
2.4 Underdrain System and Ventilation The trickling filter underdrain system is designed to meet two objectives—collect treated wastewater for conveyance to downstream unit processes and create a plenum that allows for the transfer of air throughout the trickling filter media (Grady et al., 1999). Clay or concrete underdrain blocks commonly are used for rock-media trickling filters because of the required structural support. A variety of support systems, including concrete piers and FRP grating are used for other media types. Figure 3.3 illustrates field-adjustable plastic stanchions and FRP grating on the concrete floor of a trickling filter containment structure. The volume created between the concrete and media bottom creates the underdrain. The vertical flow of air through the media can be induced mechanically or by natural air draft. Natural air ventilation results from a difference in ambient air temperature outside and inside the trickling filter. The temperature causes air to expand when warmed or contract when cooled. The net result is an air-density gradient throughout the trickling filter, and an air front either rises or sinks, depending on the differential
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Figure 3.3 Adjustable plastic stanchions (top right) and FRP grating (bottom left) on the concrete floor of a bolted steel containment structure, and an HDPE mat used to support random synthetic media (top and bottom right).
condition. This rising or sinking action results in a continuous airflow through the bioreactor. Natural ventilation may become unreliable or inadequate in meeting process air requirements, when neutral temperature gradients do not produce air movement. Such conditions may be daily or seasonal and can lead to odorous anaerobic biofilm conditions and poor trickling filter performance. Therefore, mechanical ventilation generally is included to enhance and control airflow. Mechanical ventilation is accomplished with low-pressure fans to continuously circulate air throughout the trickling filter. When using mechanical ventilation, the designer must ensure that air is distributed uniformly across the trickling filter cross-section.
2.5 Trickling Filter Pumping Stations: Influent and Recirculation A critical unit in the trickling filter process is the pumping station that lifts primary effluent (or screened raw wastewater) and recirculates unsettled trickling effluent (referred to here as underflow) to the influent stream. Generally, trickling filter underflow is recirculated to the distribution system to achieve the hydraulic load (influent + recirculation) required for proper media-wetting and biofilm thickness control and decouple hydraulic and organic loading. Although effluent from the secondary
91
92
Biofilm Reactors
clarifier can be recirculated, this not common practice, because it may lead to the hydraulic overloading of secondary clarifiers. Trickling filter influent generally is pumped to allow trickling filter underflow to flow by gravity to the SGR (or solids contact basin), secondary clarifier, or other downstream of the trickling filter. When fit with weirs, a single pumping station can be used to convey both influent and recirculation streams.
2.6 Hydraulic and Contaminant Loading Trickling filters typically are classified by the intended mode of pollutant degradation and pollutant loading, namely carbon oxidation, combined carbon oxidation and nitrification, or nitrification. Organic loading typically is expressed as kilograms per day per cubic meter (kg/d·m3) of filter media as carbonaceous 5-day biochemical oxygen demand (cBOD5). General practice is to ignore the organic load imparted by recirculation streams, but the designer should be aware of the effects that recycle flow and pollutant loading (specifically ammonia-nitrogen) may have on treatment efficiency. Total organic load (TOL) may be calculated by eq 3.1.
Q in ⋅ S S,in BOD 5 applied = TOl = media volume VM
(3.1)
where VM = trickling filter media volume (m3). Nitrifying trickling filter data typically are expressed in terms of the surface-based ammonia-nitrogen loading rate. Dividing eq 3.1 by the media specific surface area (m2/m3) produces the surface loading rate. The trickling filter hydraulic loading rate can be calculated with and without recirculation. The wastewater hydraulic load, WHL (in cubic meters per square meters per day), excludes recirculation and can be calculated mathematically by eq 3.2, which has the units cubic meters per square meters per day. Q in (3.2) a The total hydraulic load (THL) is relevant for media-wetting and biofilm thickness control and considers trickling filter influent, Qin, and recirculation stream, QR. The total hydraulic load can be expressed mathematically by eq 3.3. wHl =
THl =
Q in + Q R a
(3.3)
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
3.0 Process Flow Sheets and Bioreactor Configuration The wastewater treatment train of a trickling filter process typically consists of preliminary treatment (including screening and grit removal), primary clarification, trickling filter influent/recirculation pumping station, secondary clarification, and disinfection unit processes. The TF/SG processes also include an aeration basin and return activated sludge pumping station. This section describes standard process flow sheets, different trickling filter classifications, and hydraulics.
3.1 Standard Process Flow Diagrams The trickling filter recirculation stream influences process flow. Generally, there are two types of trickling filter recirculation. The first allows for direct recirculation to the trickling filter, and the second passes flow through a primary clarifier before entering the trickling filter. Four trickling filter process flow diagrams, including both singleand two-stage trickling filters, are shown schematically in Figure 3.4. Combined TG/ SG process flow sheets are illustrated in Figure 3.5. Recirculation of trickling filter underflow or settled effluent dilutes influent wastewater and dampens the influent organic loading variability, as a result of diurnal fluctuations. Clarifying trickling filter effluent may enhance the performance of a subsequent trickling filter in twostage operation, but the designer must ensure that the recirculation flow required for trickling filter wetting and biofilm thickness control does not exceed the limiting hydraulic loading rate for the sedimentation tank. Intermediate clarification may be beneficial for systems treating wastewaters containing high concentrations of soluble cBOD5 (likely from an industrial source) because of excessive biomass production. These solids can adversely affect the second-stage trickling filter, if not removed by intermediate settling. The design of settling tanks in two-stage trickling filter systems also is affected by the recirculation pattern. The practice of alternating the lead trickling filter in a primary–secondary trickling system is referred to as an alternating double-filtration (ADF) system. This concept is beneficial when applied to modular synthetic-media trickling filters. Gujer and Boller (1986) and Parker et al. (1989) observed patchy biofilm growth in the lower section of pilot-scale nitrifying trickling filters (NTFs). The researchers attributed the patchy growth to dry spots. Boller and Gujer (1986) recommended the use of ADF to enhance pilot-scale NTF performance. Aspegren (1992) also observed improved nitrification using ADF compared with a single-stage NTF and reported that the
93
TFINF
RS
PE
TF
PC TFRCY
PE
TF
PC PS
TFEFF
WS
TFINF-2
TF
TF
(c)
TFEFF
PC
SE SC
TFRCY
PS
WS
TFINF-1
PE
TF
PC PS
(b)
SE SC
PE
RS
(a)
WS
TFRCY
TFINF-1
RS
SE SC
PS
RS
TFEFF
TFINF-2
TFEFF
ICE
TF
IC TFRCY-1
WS
(d)
TFEFF
SE SC
TFRCY-1
WS
Figure 3.4 Typical flow diagrams for the trickling filter process: (a) and (b) singlestage trickling filter process, (c) two-stage trickling filter process, and (d) two-stage trickling filter process with intermediate clarification. RS = raw sewage, PC = primary clarifier, PS = primary sludge, PE = primary effluent, TFINF = trickling filter influent, TF = trickling filter, TFEFF = trickling filter effluent, TFRCY = trickling filter recycle, SC = secondary clarifier, WS = waste sludge, SE = secondary effluent, IC = intermediate clarifier, and ICE = intermediate clarifier effluent. 94
TFINF
RS
PE
ABF
TFEFF
PC
SE
TFRCY
PS
RAS
RS
PE
TF
PC
WS
TFEFF
MLSS
TFRCY
PE
ABF
PC
RAS
TFEFF
(b)
WS
MLSS
SE
(c)
SecC
SC/AS
TFRCY
PS
SE SecC
SC/AS
PS
RS
(a)
SecC
RAS
WS
TFINF-1 (d) RS
PE
TF
PC PS
TFRCY-1
TFEFF
MLSS
ICE IC WS
SE SecC
AS RAS
WS
Figure 3.5 Typical flow diagrams for trickling filter-suspended growth processes: (a) activated biofilter, (b) trickling filter/solids contact or trickling filter/activated sludge, (c) activated biofilter/solids contact or activated biofilter/activated sludge, and (d) trickling filter and activated sludge. ABF = activated biofilter, AS = activated sludge, SC = solids contact, RAS = return activated sludge, and SecC = secondary clarifier; other variables as defined in Figure 3.4. 95
96
Biofilm Reactors
improved performance resulted from the reduced biofilm patchiness. Use of the ADF approach with two trickling filters in series encourages full-depth biofilm development in both bioreactors. The lead trickling filter should be switched every 3 to 7 days, to ensure that healthy biofilms, which maximize media coverage, are developed along the entire bioreactor depth. The primary drawback of ADF is an increased power requirement, which may be in excess of 50%, as a result of double pumping. In addition to increased operating cost, capital costs associated with pipes and valves also will increase. Sludge handling also affects the trickling filter process. Each of the process-flow diagrams illustrated in Figures 3.4 and 3.5 implies that waste biological solids are removed by co-settling the biological sludge with the primary sludge before withdrawal from the system. Many facilities exist that separately handle the primary and secondary sludge. The benefits of each must be evaluated by the designer. This practice does require that operators conscientiously withdraw the solids from the process and that designers provide equipment and means to maintain a near-zero sludge blanket, if necessary. A common operational problem resulting from improper maintenance of biosolids is “rising sludge.” Any trickling filter application that results in nitrification produces nitrate, which is reduced to nitrogen gas (N2) in an anoxic sludge blanket macroenvironment (and the presence of an electron donor). The N2 (g) can become entrained in the sludge blanket and float biomass to the sedimentation basin water surface. This biomass can float over the weirs and degrade secondary effluent water quality. Another issue that designers must consider is the improper maintenance of a primary clarifier sludge blanket. When combined with waste biological sludge, the soluble cBOD5 that exists in primary sludge may generate odor. Alternatively, particulate BOD5 may hydrolyze and cause soluble BOD5 to enter the bulk liquid. This can result in an increased TOL (on the trickling filter) and diminish bioreactor performance.
3.2 Bioreactor Classification Trickling filters can be categorized by four modes of operation or application—(1) roughing, (2) carbon oxidation, (3) carbon oxidation and nitrification, and (4) nitrification. Table 3.2 summarizes generally accepted criteria defining each operational mode. Roughing filters receive high hydraulic and high organic loadings and generally are constructed with vertical-flow media to minimize the potential for excess biofilm accumulation. Although they may provide a high quantity of organic load removal per unit volume, their settled effluent still contains substantial BOD5. Roughing filters provide approximately 50 to 75% soluble BOD5 conversion and may receive total
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
97
Table 3.2 Trickling filter classification. Carbon oxidizing cBOD5 and (cBOD5 removal) nitrification
Design parametera
Roughing
Media typically used
Vertical flow
Rock, cross-flow, or vertical flow
Rock, cross-flow, or vertical flow
Cross-flow
Wastewater source
Primary effluent
Primary effluent
Primary effluent
Secondary effluent
Hydraulic loading
52.8 to 178.2 (0.9 to 2.9)
14.7 to 88.0 (0.25b to 1.5)
14.7 to 88.0 (0.25b to 1.5)
35.2 to 88.0 (0.6 to 1.5)
1.6 to 3.52 (100 to 220)
0.32 to 0.96 (20 to 60)
0.08 to 0.24 (5 to 15)
NAc
NAc
NAc
0.2 to 1.0 (0.04 to 0.2)
0.5 to 2.4 (0.1 to 0.5)
Effluent quality (mg/L) (unless noted)
50 to 75% filtered cBOD5 conversion
15 to 30 mg/L cBOD5 and TSSd
<10 cBOD5; <3 NH3-Nd
0.5 to 3 NH3-Nb
Predation
No appreciable growth
Beneficial
Detrimental Detrimental (nitrifying biofilm)
Filter flies
No appreciable growth
No appreciable growth
No appreciable growth
No appreciable growth
Depth (m [ft])
0.91 to 6.10 (3 to 20)
≤12.2 (40)
≤12.2 (40)
≤12.2 (40)
m 3 (gpm/sq ft) d.m 2
Nitrification
BOD5 and NH3-N load kg m 3 .d g m 2 .d
(lb BOD5/d·1000 cu ft) (lb NH3-N/d·1000 sq ft)
gpm/sq ft × 58.674 = m3/m2·d (cubic meter per day per square meter of trickling filter plan area); lb BOD5/ d/1000 cu ft × 0.0160 = kg/m3∙d (kilograms per day per cubic meter of media); and lb NH3-N/d/1000 sq ft × 4.88 = g/m2∙d (grams per day per square meter of media). a
b Applicable to shallow trickling filters. gpm/sq ft = gallons per minute per square foot of trickling filter plan area. c
Not applicable.
d
Concentration remaining in the clarifier effluent stream.
BOD5 loadings of 1.5 to 3.5 kg BOD5/m3∙d. Carbon-oxidizing trickling filters may provide settled effluent concentrations of 15 to 30 mg/L for BOD5 and TSS, respectively, and may receive BOD5 loadings of 0.7 to 1.5 kg/m3∙d. Combined carbon oxidation and nitrification trickling filters may provide an effluent BOD5 concentration less than 10 mg/L and NH3-N less than or equal to 0.5 to 3 mg/L (after solids separation). These trickling filters may receive BOD5 loadings less than 0.2 kg/m3∙d and TKN loadings
98
Biofilm Reactors
of 0.2 to 1.0 g/m2∙d. The NTFs provide 0.5 to 3 mg/L effluent NH3-N when receiving a clarified secondary effluent and NH3-N loadings of 0.5 to 2.5 g N/m2∙d.
3.3 Hydraulic Application: Effect on Media Wetting, Flow Distribution, and Control Recirculation and distributor operation are important to good trickling filter performance and may be used to achieve proper media-wetting, flow distribution, biofilm thickness control, and prevent macrofauna accumulation. Albertson and Eckenfelder (1984) postulated that the active biofilm surface area in a trickling filter is dependent on biofilm thickness and media configuration and that the active biofilm surface area decreases with increasing biofilm thickness. The researchers stated that, for mediumdensity cross-flow media with 98 m2/m3 specific surface area, a 4-mm increase in biofilm thickness would cause a 12% reduction of active biofilm area (assuming that the entire media has been appropriately wetted). Poor trickling filter media-wetting results in reduced effluent water quality. In a study of rotary distributor efficiency, Crine et al. (1990) found that the wetted-area-to-specific-surface-area ratio ranged from 0.2 to 0.6, with the lowest values for high-density random-pack media. Figure 3.6 shows a threedimensional representation of liquid-flow distribution in cross-flow media, both with and without biofilm (Lekhlif et al., 1994). Many of the later mentioned design formulations incorporate a term that allows for specific surface area reduction, as a result of
(a)
(b)
Figure 3.6 Three-dimensional representation of liquid flow distribution in crossflow media, both (a) with and (b) without biofilm (Lekhlif et al., 1994).
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
distributor inefficiency in trickling filter media-wetting. The interrelationship of liquid residence time, dosing, and media configuration on BOD5 removal kinetics has not been addressed, and additional research is required. Increasing the average hydraulic application rate reduces the liquid residence time, but has been proven to increase wetting efficiency. The recirculation ratio (Q/QR) typically is in the range 0.5 to 4.0. Data obtained by Dow Chemical Company (Bryan, 1955, 1962; Bryan and Moeller, 1960) demonstrate that vertical-flow media require an average application rate exceeding 1.8 m3/m2·h to maximize the BOD5 removal efficiency. Shallow towers using crossflow media have used hydraulic rates in the range 0.4 to 1.1 m3/m2·h. Slowed distributor operation has benefited trickling filter facilities because of interrupted flow (periodicity of dosing), increasing wetting efficiency (percent of media wetted), and biofilm thickness control. The designer should consider recirculation capabilities, the effect of reverse-thrusting jets, or the use of speed control on the distributor, to enhance performance or improve operation. A German-developed (ATV, 1983) process control parameter, referred to as Spülkraft, allows for the calculation of a dosing rate (mm/pass) by eq 3.4. mm m Sk = mi N a ⋅ ω d ⋅ 60 hour THl ⋅ 1000
Where SK = Spülkraft (mm/pass), Na = number of arms on the distributor, and ωa = distributor rotational speed (rpm).
(3.4)
The typical hydraulically driven distributor in North America operates in the range 2 to 10 mm/pass. Table 3.3 lists recommended dosing rates for rotary distributors. Higher dosing rates are recommended for higher organic loading rates, to enhance biofilm thickness control. Besides a normal operating dosing rate, it may be beneficial to periodically use a higher flushing dosing rate for 5 to 10% of the 24-hour operating period. Albertson (1989a, 1989b, 1995) and Parker and coworkers (Parker, 1999; Parker et al., 1995, 1997) demonstrated that biofilm thickness control enhanced nitrification in the trickling filter process, reduced odors, reduced power use for recycling, reduced the accumulation of nuisance organisms, and eliminated heavy sloughing cycles. Parker and coworkers (Parker, 1999; Parker et al., 1995, 1997) recommended the use of VFD-controlled recirculation pumps and distributor speed control
99
100
Biofilm Reactors
Table 3.3 Operating and flushing dosing rates for distributors.* Total organic load (kg/m3·d [lb BOD5/d·1000 cu ft])
Operating dosing rate (mm/pass [in./pass])
Flushing dosing rate (mm/pass [in./pass])
<0.4 (<25)
25 to 75 (1 to 3)
100 (4)
0.8 (50)
50 to 150 (2 to 6)
150 (6)
1.2 (75)
75 to 225 (3 to 9)
225 (9)
1.6 (100)
100 to 300 (4 to 12)
300 (12)
2.4 (150)
150 to 450 (6 to 18)
450 (18)
3.2 (200)
200 to 600 (8 to 24)
600 (24)
*Actual values are site-specific and vary with media type.
(if necessary) to maintain constant trickling filter hydraulic application. Parker et al. (1995) conducted pilot studies, which demonstrated that mechanically driven distributor dosing alone did not improve the performance of an NTF. Parker (1999) has pointed out that there is little research describing the effect of hydraulic transients on synthetic trickling filter media and their effect on media life.
4.0 Ventilation and Air Supply Alternatives Trickling filters require oxygen to sustain aerobic biochemical transformation processes. Several researchers have demonstrated that at least some portion (if not the entire bioreactor) of roughing, carbon-oxidizing, combined carbon-oxidizing and nitrification, and NTFs operates under oxygen-limited conditions (Kuenen et al., 1986; Okey and Albertson, 1989a; Schroeder and Tchobanoglous, 1976). Ventilation is essential to maintain aerobic conditions in a trickling filter. The vertical flow of air through trickling filter biofilm carriers can be induced by natural air draft or mechanical ventilation. Mechanical ventilation enhances and controls airflow with low-pressure fans, which continuously circulate air throughout the trickling filter. Current design practice requires the provision of adequate underdrain and effluent channel sizing to permit free airflow. Passive devices for ventilation include vent stacks on the trickling filter periphery, extensions of underdrains through trickling filter side walls, ventilating manholes, louvers on the side wall of the tower near the underdrain, and discharge of trickling effluent to the subsequent settling basin in an open channel or partially filled pipes.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
4.1 Natural Draft Naturally occurring airflow results from a difference in ambient air temperature and humidity outside and inside the trickling filter. The temperature causes air to expand when warmed or contract when cooled. The result is an air-density gradient throughout the trickling filter, and an air front rises or sinks, depending on the differential condition. This rising or sinking action results in a continuous airflow through the bioreactor. If the air inside the trickling filter is colder than the ambient air, the air will flow downward. Alternatively, if the ambient air is colder than the air inside trickling filter, air will flow upward. Schroeder and Tchobanoglous (1976) state that upward airflow is the worst case from a mass-transfer perspective, because the dissolved oxygen driving force is lowest in the region of highest oxygen demand (i.e., the top of the trickling filter). Natural ventilation may become unreliable or inadequate in meeting process air requirements when neutral temperatures do not produce air movement. Such conditions may be daily or seasonal and can lead to the development of anaerobic biofilms and poor trickling filter performance. Modular synthetic-media trickling filters that rely on natural draft to provide process oxygen for municipal wastewater treatment should include the following design features: • Drains, channels, and pipes should be sized sufficiently to prevent submergence greater than 50% of their cross-sectional area under design hydraulic loading. • Ventilating access ports with open grating covers should be installed at both ends of the central collection channel. • Large-diameter trickling filters typically have branch channels (to collect the treated wastewater). • These branches also should include ventilating manholes or vent stacks installed at the trickling filter periphery. • The open area of the slots in the top of the underdrain blocks should not be less than 15% of the trickling filter area. • Approximately a 1-m2 gross area of open grating in ventilating manholes and vent stacks should be provided for each 23 m2 of trickling filter area. Typically, 0.1 m2 of ventilating area is provided for every 3 to 4.6 m of trickling filter periphery, and 1 to 2 m2 of ventilation area is provided in the underdrain area per 1000 m3 of trickling filter media. • Another criterion for rock-media trickling filters is the provision of a vent area at least equal to 15% of the trickling filter cross-sectional area.
101
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Biofilm Reactors
4.2 Mechanical Ventilation A majority of new and improved trickling filters use low-pressure fans to mechanically induce airflow. The airflow resulting from natural draft will distribute itself; this will not occur with mechanical ventilation. Pressure loss through synthetic trickling filter media typically is low—often less than 1-mm water (H2O) per meter of trickling filter depth (Grady et al., 1999). The low pressure drop typically results in low fan power requirements (approximately 3 to 5 kW). The head on the fan typically is less than 1500-mm H2O. Unfortunately, the low pressure drop allows air to rise upward through the trickling filter media without distributing itself across the bioreactor section. Therefore, fans typically are connected to distribution pipes. The airflow distribution piping has openings sized such that airflow through each is equal and airflow distribution is uniform. The pipes typically have a velocity in the range 1100 to 2200 m/h, to further promote uniform airflow distribution. Airflow requirements are calculated based on process oxygen requirements and characteristic oxygen transfer efficiency, which is typically in the range 2 to 10%. The mechanical air stream may flow upward to downward. Downflow systems can be designed without covers; however, covers are required for systems that do not have air distribution through a network of pipes under the media. Covering trickling filters offers a benefit in winter of limiting cold airflow and minimizing wastewater cooling. Mechanical ventilation and covered trickling filters may be used to destroy odorous compounds. A trickling filter mechanical aeration system (distribution pipes and fans) is pictured in Figure 3.7. (a)
(b)
Figure 3.7 (a) Trickling filter underdrain and media supports with air distribution piping throughout the underdrain, and (b) typical low-pressure fan applied to provide process air.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
5.0 Trickling Filter Process Models Numerous investigators have attempted to delineate the fundamentals of the trickling filter process, by developing relationships among variables that affect bioreactor operation. Existing trickling filter process models range in complexity, from simplistic empirical formulations to complex mathematical models. Analyses of operating data have been made to establish equations or curves to fit available data. Results of these data analyses have led to the development of various empirical formulas that describe trickling filter performance. Unfortunately, numerous models exist, and there is lack of industry standard defining a method for the selection of an appropriate process model. Trickling filter models may be classified as dissolved organic loading-based models, particulate organic loading-based models, hydraulic loading-based models, and mass-transfer models. Although these formulas represent attempts to include key variables that have been observed to affect trickling filter operation, none of the formulas adequately describe the complex processes that are characteristic of trickling filter bioreactors. Designers need to assess which equation best fits a particular situation when selecting which model to apply in a particular design, especially with regard to the confidence level necessary to meet project-specific effluent water quality (permit) requirements. The following models are discussed in this section:
(1) National Research Council (1946); (2) Galler and Gotaas (1964); (3) Kincannon and Stover (1982); (4) Velz (1948); (5) Schulze (1960); (6) Germain (1966); (7) Eckenfelder and Barnhart (1963); (8) Chartered Institution of Water and Environmental Management (1988); and (9) Logan trickling filter model (TRIFL) (Logan et al., 1987a, 1987b).
5.1 National Research Council The National Research Council (Washington, D.C.) (NRC) formula (1946) resulted from an extensive analysis of operational records from full-scale rock‑media trickling filters serving military installations. The NRC data analysis demonstrated that cBOD5 removal in the rock-media trickling filters studied was dependent on contact time between the biofilm and bulk liquid, which is dependent on bioreactor dimensions and the number of passes. The removal efficiency of cBOD5 increases with contact
103
104
Biofilm Reactors
time. Because treatment efficiency decreases with organic load, the primary determinant of treatment efficiency in a rock-media trickling filter, according to the NRC formulas, is increased contact (with the biofilm surface) and applied load. For the 34 rock‑media trickling filter installations evaluated during the NRC study, a curvefitting effort resulted in eqs 3.5 and 3.6 for single-stage and second-stage trickling filters, respectively.
E1 =
E2 =
100 W 1 + 0.0085 ⋅ V ⋅ F
0.5
(3.5)
100 W2 1 + 0.0085 ⋅ E V⋅ F 1− 1 100
(3.6)
Where E1 = BOD5 removal efficiency through the first‑stage trickling filter and settling tank (%); W1 = BOD5 loading to the first‑ or single‑stage trickling filter, not including recycle (kg/d [lb/d]); V = volume of the trickling filter stage (cross-sectional area × media depth) (m3 [ac-ft]); 1+ F = number of organic material passes =
QR Q
QR 1 + ( 1 − P ) ⋅ Q
2
;
QR = dimensionless recirculation ratio; Q P = a weighing factor, for military trickling filters, with rock media = 0.9; E2 = BOD5 removal efficiency through second‑stage trickling filter and settling tank (%); and W2 = BOD5 loading to the second‑stage trickling filter, not including recycle (kg/d [lb/d]). Equations 3.5 and 3.6 are empirical formulas based on data obtained from rock‑ media trickling filters with and without recirculation. Because of the nature of their development, the NRC formulas include the following limitations and conditions:
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
(1) Characteristically, military wastewater has a higher strength (250 to 400 mg BOD5/L) than average domestic wastewater. (2) The effect of temperature on trickling filter performance is not considered (most of the studies were conducted in the Midwest and southern latitudes of the United States). (3) The clarifier design practice, at the time the formulas were developed, favored shallow units that were operated with a higher hydraulic load than recommended in current practice. The shallower clarifiers may have resulted in BOD5 and TSS losses. (4) Applicability may be limited to stronger-than-normal domestic wastewater, because no factor is included to account for differing treatability rates for lower-strength wastewater.
The formula for second‑stage trickling filters is based on the existence of intermediate settling tanks following the first‑stage trickling. Results from the NRC formula may deviate substantially from actual trickling filter performance. Scattered data for loadings less than 0.3 kg/m3·d (20 lb/d/1000 cu ft) may be biased by a lack of BOD5 test protocol requiring the inhibition of nitrification during the BOD5 test. Inadequate hydraulic application rate, poor ventilation, and inefficient clarifier design could have contributed to poor trickling filter process performance. Thus, the foregoing should be considered when designing the trickling filter process using the NRC formula(s).
5.2 Galler and Gotaas Galler and Gotaas (1964) attempted to describe the performance of rock-media trickling filters through a multiple regression analysis of data obtained from existing full-scale trickling filters. Equation 3.7 was developed based on an analysis of 322 observations.
Se =
Where
K ⋅ ( Q ⋅ Si + QR ⋅ Se ) 1.19
( Q + QR ) 0.78 ⋅ ( 1+ D ) 0.67 ⋅ ra0.25
43 , 560 0.464 ⋅ π K = coefficient = 0.28 0.15 Q ⋅T
0.13
,
Q = flowrate (ML/d [mgd]), QR = recirculation flowrate (ML/d [mgd]),
(3.7)
105
106
Biofilm Reactors
D = trickling filter depth (m [ft]), Se = settled trickling filter effluent BOD5 concentration at 20°C (mg/L), Si = trickling filter influent BOD5 concentration at 20°C (mg/L), and ra = trickling filter radius (m [ft]). Some key variables of the Galler and Gotaas (1964) formula include recirculation, hydraulic loading, trickling filter depth, and wastewater temperature. According to their analysis, deeper trickling filters increased BOD5 removal. The researchers further indicated that recirculation improved trickling filter performance, but established a 4:1 ratio as a practical upper limit for recirculation. The hydraulic flowrate through the trickling filter was found to be statistically unimportant in determining the bioreactor treatment efficiency. The statistical analysis performed by Galler and Gotaas (1964) suggested that BOD5 loading was, statistically, the most significant parameter for describing bioreactor performance.
5.3 Kincannon and Stover Kincannon and Stover (1982) developed a mathematical model based on a relationship between the specific substrate utilization rate and total organic loading, which followed a Monod plot to determine the required biofilm area. The relationship is described mathematically by eq 3.8.
8.345 ⋅ q ⋅ Si µ ⋅ S max i − Kb As = Si − Se
(3.8)
Where Si = influent BOD5 concentration (mg/L), Se = effluent BOD5 concentration (mg/L), and Kb = proportionality constant of specific surface area (m2). Biokinetic parameters, namely the maximum specific substrate utilization rate and Monod-type half-saturation constant (or μmax and Ki, respectively), are reported based on pilot‑plant tests, full‑scale observations, or a summary of previous experiences. When extracting these parameters from test data, these parameters may be determined graphically by plotting BOD5 loading versus the inverse of BOD5 removed. The y intercept is μmax-1, and the slope is Kb. According to the Kincannon and Stover (1982) model, BOD5 removal is most sensitive to volumetric loading and treatability. The BOD5 removal is not influenced by trickling filter depth.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
5.4 Velz Velz (1948) proposed the first formulation based on a fundamental law, in contrast with previous empirical attempts, which were based on data analyses. The Velz equation (eq 3.9) relates BOD5 remaining in the bulk phase as follows: Se = 10− kV ⋅ t Si
(3.9)
Where Si = influent BOD5 concentration (mg/L), SD = BOD5 concentration remaining at filter depth D (mg/L), t = residence time (days), and kV = Velz first-order rate constant (d-1). Equation 3.9 suggests that kV is constant for all hydraulic rates; however, Albertson and Davies (1984) presented evidence that kV varies with the hydraulic rate. The Velz equation is presented herein because of its foundation of later-developed design formulations, which are applied in present day design, namely the Eckenfelder and Barnhart (1963) and Schulze (1960) equations.
5.5 Schulze Schulze (1960) postulated that the tome liquid that is in contact with the biofilm is directly proportional to the trickling filter depth and inversely proportional to the hydraulic loading rate. This is expressed by eq 3.10.
Where
tc =
c⋅ D THln
(3.10)
tc = liquid contact time (days), c = constant (dimensionless), D = trickling filter depth (m), THL = hydraulic loading rate (m3/m2·d), and n = exponent on hydraulic loading (dimensionless). Combining the time of contact with the first‑order equation for BOD5 removal in an adaptation of the Velz (1948) equation, Schulze derived the following formula (eq 3.11):
Se =e Si
− kS ⋅ D THln
(3.11)
107
108
Biofilm Reactors
Where Se = soluble BOD5 in trickling filter effluent stream (mg/L), Si = soluble BOD5 in trickling filter influent stream (mg/L), kS = Schulze coefficient (d−1 when n = 1), D = trickling filter depth (m), N = exponent on hydraulic loading (dimensionless), and THL = hydraulic loading rate (m3/m2·d). Equation 3.11 is similar to that proposed by Velz (1948). However, Velz’s constant, kV, was not formulated to consider hydraulic load. For a given wastewater strength, the hydraulic rate is proportional to the loading rate. Thus, volumetric organic loading still may be the controlling process variable. The value of kS published by Schulze (based on United States customary units) for a rock-media trickling filter with a 1.8‑m (6‑ft) depth at 20°C was 0.69 day-1. The dimensionless constant characteristic of rockmedia trickling filter, n, was found to be 0.67. The common temperature correction value of θ = 1.035 could be applied to determine kt, as follows: kt = k20 × 1.035( T − 20 ) for ks
(3.12)
Where kt = temperature-corrected coefficient value (d-1 when n = 1), kS = Schulze coefficient (d-1 when n = 1), and As = clean surface area of the media (m2).
5.6 Germain Germain (1966) applied the Schulze (1960) formulation to a synthetic-media trickling filter (eq 3.13). − k ⋅D
Where
G Se n = e THL Si
(3.13)
Si = soluble BOD5 in trickling filter influent stream (typically primary effluent excluding recirculation) (mg/L), THL = hydraulic loading rate (m3/m2·d), and kG = Germain coefficient (d-1 when n = 1). The values of kG and n are related to media configuration, secondary clarifier efficiency, dosing, and hydraulic application rate; k G is a function of wastewater
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
characteristics, trickling filter depth, media specific surface area, and configuration. Therefore, because a high degree of interdependency exists between kG and n, this must be considered in data comparisons. Germain (1966) reported the value of kG of 0.24 (L/s)n∙m2 (0.088 gpmn/sq ft) for a synthetic-media trickling filter having a 6.6 m (21.5 ft) depth and treating domestic wastewater with a value of 0.5 for n. This verticalflow media had a clean surface area of 89 m2/m3. Correction of kG for the high BOD5 concentration represented by the Chartered Institution of Water and Environmental 150 Management (London, United Kingdom) (CIWEM) model, kG 360
0.5
, resulted
in similar predictive values from these two models for plastic media operating in the loading range of 0.2 to 1.5 kg/m3·d at 20°C. In tests designed to determine the effects of recirculation on BOD 5 removal, Germain (1966) found no statistically significant difference. However, the relatively tall (6.6 m) tower resulted in high influent application rates, which ensured adequate media-wetting. This observation is consistent with the practice of using recirculation for shallow trickling filters subject to low influent hydraulic rates and likely inefficient wetting. Equation 3.13 has been used extensively for synthetic-media trickling filter analysis and design. The kG data were developed from more than 140 pilot studies performed by Dow Chemical Company and others. Most of these tests used a trickling filter media depth of 6 to 7 m.
5.7 Eckenfelder Eckenfelder and Barnhart (1963) expanded upon previously developed trickling filter formulas, which account for trickling filter media specific surface area. The formula proposed for soluble BOD5 removal can be expressed mathematically by eq 3.14. Se =e Si
− kS’ ⋅ a( 1+ b ) ⋅ D THLn
(3.14)
Where kS’ = overall treatability coefficient based on soluble BOD5 [(m3/d)0.5·m2], D = depth (m), THL = hydraulic loading rate (m3/m2·d), and b = surface area modifier for surface loss with increasing area. With recirculation, eq 3.14 can be extended and expressed mathematically by eq 3.15.
109
110
Biofilm Reactors
− ks’ ⋅ D
n
e WHL Se = Si QR 1 + Q − e
(3.15)
− k ’s ⋅ D WHLn
Using the Eckenfelder formula and kS’ = a × ks, eq 3.15 can be rewritten as eq 3.16; this is known as the modified Velz equation. Si
Se =
kS ⋅ a⋅ D⋅θ ( T − 20 )
QR + 1 Q
QR WHL⋅ Q + 1 ⋅ e
(3.16)
n
−
QR Q
5.8 Chartered Institution of Water and Environmental Management The CIWEM developed a formula describing the BOD5 in trickling filters having rock, random-packed synthetic media, or modular synthetic-plastic media. Equation 3.17, resulting from a multiple regression analysis, follows:
Si = Se
1
(3.17)
a 1 + kC IWEM ⋅ θ ( T − 15 )⋅ VLRn m
Where Si = influent BOD5 concentration (mg/L), Se = effluent BOD5 concentration (mg/L), kCIWEM = kinetic coefficient (mm-1·d n-1), θ = temperature coefficient, a = media specific surface area (m2/m3), m = reduction factor for surface loss with increasing area, VLR = volumetric hydraulic loading rate (m3/d·m3) of trickling filter media, and n = hydraulic rate coefficient. Equation 3.17 has reported coefficients that account for 90% of data variability.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
• kCIWEM = 0.0204 (rock and random), 0.40 (modular plastic); • θ = 1.111 (rock and random), 1.089 (modular plastic); • m = 1.407 (rock and random), 0.732 (modular plastic); and • n
= 1.249 (rock and random), 1.396 (modular plastic).
The model was developed using data collected from tests performed on a strong domestic wastewater with primary effluent concentrations of 360 mg/L BOD5, 240 mg/L TSS, and 52 mg/L NH3-N. The model predicts a continuous performance curve from low‑ to high‑rate loadings. The trickling filter depths, from which the samples were collected, ranged from 1.74 to 2.10 m; biofilm growth areas ranged from 1.0 to 5.0 m2; and loadings were 0.3 to 16 kg/m3·d. The CIWEM model is temperaturesensitive, which may be caused by site‑specific wastewater characteristics and datareduction procedures. The NRC equations agree with the CIWEM’s projection based on an influent strength of 360 mg/L BOD5 at loadings up to 1.0 kg/m3·d.
5.9 Logan Trickling Filter Model The TRIFL model, which commonly is referred to as the Logan Trickling Filter (LTF) model, is based on characterizing modular synthetic trickling filter media as a series of inclined plates covered with a thick, partially penetrated, biofilm. The rate of soluble chemical oxygen demand (SCOD) removal is determined using a numerical model to solve transport equations that describe biochemical transformation rates resulting from diffusion through a thin liquid film and into the biofilm. Although the model was calibrated using a single data set for only one type of plastic trickling filter media, a variety of laboratory, pilot-plant, and full-scale trickling filter studies have been conducted, which suggest that the LTF model accurately predicts SCOD removal (Bratby et al., 1999; Logan and Wagenseller, 2000; Logan et al., 1987a, 1987b). Unlike kinetic models, the LTF model cannot be described with a single equation; thus, a computer program is required. The computer model of Logan et al. (1987a, 1987b) was developed to predict soluble BOD5 removal in plastic media trickling filters as a function of plastic media geometry. A disadvantage of kinetic models, such as the Velz equation, is that new kinetic (k20) and hydraulic (n) constants may require determination for each type of trickling filter media. The LTF model requires only that the media geometry be measured and input. Consequently, there was no need to recalibrate the model for new synthetic media types. The actual, computer-code-based model (written in Fortran [IBM,
111
112
Biofilm Reactors
San Jose, California]) was given the name TRIFIL2; the LTF model computer program uses tabulated values for a range of conditions for specific media. Dissolved organics that compose SCOD in wastewater are assumed to be equally distributed into a fivecomponent molecular size. As the wastewater flows over the biofilm, the dissolved organics diffuse into the biofilm. Small molecules diffuse faster than larger ones and are predicted to be removed more efficiently. Temperature affects water viscosity (μ), which affects fluid film thickness and thus retention time in the trickling filter media. Changes in chemical diffusion coefficients (D) with temperature (T) are adjusted by D⋅ µ is constant (Welty et al., 1976). The model is available the assumption that T free-of-charge at http://www.engr.psu.edu/ce/enve/logan/bioremediation/ trickling_filter/model.htm (Penn State University, University Park, Pennsylvania). Additional information on the model can be obtained from the original publications cited below and a chapter in Logan (1999).
5.10 Selecting a Trickling Filter Model Design engineers may use various empirical criteria and design formulations for sizing trickling filters. The NRC (1946) (eqs 3.5 and 3.6) or Galler and Gotaas (1964) (eq 3.7) formulas commonly are used for rock-media trickling filter design. The Schulze equation (eq 3.11), Eckenfelder equation (eq 3.14), and CIWEM equation (eq 3.17) are commonly used formulas for both rock- and synthetic-media trickling filter design over a wide range of media-specific surface areas and trickling filter depths. The coefficients k and n vary, however. The word coefficient is used to describe k (or K) and n, because they are not constants or treatability factors. Thus, the value of k in hydraulically based equations can be modified as a function of media depth (D) by eq 3.18, as follows:
D k2 = k1 ⋅ 1 D2
0.5
(3.18)
The variation of k with trickling filter depth is an important consideration. A k-value developed for a specific trickling filter depth should not be used for a different depth without modification. Using data from a number of installations and simultaneous tests, Albertson and Davies (1984) showed that k could be used effectively for any trickling filter configuration, if corrected for depth. However, this
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
research also indicated that inadequate wetting might have produced lower k values than optimum in many studies. The Eckenfelder, Germain, Schulze, and Velz equations are fundamentally the same and have similar limitations. Because the coefficients k (or K) and n are derived empirically, background data are influenced by a host of variables, such as hydraulic loading rate, dosing mechanism, temperature, wastewater characterization, media configuration and depth, and ventilation (i.e., air supply). The above equations have been applied to accurately describe trickling filter performance, but also have been proven to deviate significantly from observed performance. When modifying the trickling filter configuration, the coefficient k (or K) value typically changes, even when considering the same trickling filter media and wastewater. The NRC, Germain, and Eckenfelder equations may be used for rock-media trickling filter design, although it is recognized that the results may be highly variable. In many cases, designers use each of these design models to bracket the design variables before making a design decision. Synthetic-media trickling filters may be designed using the Eckenfelder, Germain, and/or TRIFL models. However, the designer must be aware that much of the historical pilot‑plant and full‑scale data are impaired by a lack of proper dosing (i.e., poor media-wetting and biofilm thickness control). In addition, many pilot plants used to generate coefficients typically used in the design models summarized here were equipped with continuous-flow nozzles. Therefore, the pilotscale units failed to capture the effect of periodic dosing that is typical of full-scale operating trickling filters. The Eckenfelder equation often is used to define soluble organic matter removal efficiency. The beneficial effect of recirculation is reflected in this formula; the benefits were derived from low application rates typical of carbon-oxidizing rock-media trickling filters. The literature values of n were derived from continuous-flow studies. To compare the k values, the use of 0.5 for n is suggested.
5.11 Method for Combining Trickling Filter and Suspended-Growth Models A previous section presented several design models that commonly are used for the design and evaluation of trickling filters, but there is generally a lack of models describing TF/SG systems. Daigger et al. (1993) and Takács et al. (1996) have developed TF/SG process models.
113
114
Biofilm Reactors
The model of Daigger et al. (1993) was established based on performance observations at the Garland WWTP, Garland, Texas. The TF/SG model accounts for SGR autotrophic nitrifier seeding by the trickling filter. The mass of autotrophic nitrifiers propagating toward the SGR is calculated as a function of the ammonia-nitrogen oxidized in the trickling filter. Model inputs include wastewater temperature, SGR solids retention time (SRT), and trickling filter influent and effluent ammonia-nitrogen. The TF/SG process effluent then is calculated with eq 3.19. 2 1 µmax − MCRT + kd ⋅ NH 3eFF + 1 + kd ⋅ NH 3TFe ⋅ K s NH 3eFF + MCRT
(
)
(
)
1 + kd ⋅ NH 3TFe − K s − µmax ⋅ NH 3Pe ⋅ MCRT = 0 (3.19)
(
)
Where
µmax = maximum nitrifier growth rate (d-1), MCRT = mean cell residence time (days), kd = specific decay rate (m/d), Ks = ammonia-nitrogen half-saturation constant (mg/L), NH3,EFF = TF/SG process effluent ammonia-nitrogen (mg/L), NH3,TFE = trickling filter process effluent ammonia-nitrogen (mg/L), and NH3,PE = trickling filter process influent ammonia-nitrogen (mg/L). The performance relationship developed by eq 3.19 is illustrated in Figure 3.8, where the effluent ammonia-nitrogen concentration is plotted as a function of the SGR SRT divided by the minimum nitrifier SRT. There are several curves that correspond to different nitrification efficiencies inside the trickling filter. Detachment of autotrophic nitrifiers from biofilm growing inside the trickling filter allows the SGR to maintain nitrification, even when operating at a nitrification design factor that ordinarily would result in autotrophic nitrifier washout. The model of Daigger et al. (1993) has been evaluated independently by Biesterfeld et al. (2005). The researchers applied the model to performance observations at the Boulder 75th Street WWTP, Boulder, Colorado, and subsequently used the model for process optimization. The researchers noted that the model of Daigger et al. (1993) is primarily dependent on the rate of nitrification inside the trickling filter and SGR SRT.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
20
Influent Ammonia-N Conc. = 20 mg N/L Temp. = 20°C
Effluent Ammonia - N Conc., mg N/L
18 16 14 12 10
0 20
8 6 4 2 0
Trickling Filter Nitrification Efficiency
40 60 80 0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.0 Suspended Growth SRT SRTmin for Nitrification
Figure 3.8 Effect of SRT in a suspended-growth reactor and nitrification efficiency in an upstream trickling filter on effluent ammonia-nitrogen concentration in a combined TF/SG reactor (Daigger et al., 1993).
6.0 Process Design Basic criteria defining various trickling filter types have been presented. This section describes, in greater detail, process design criteria for combined carbon oxidation and nitrification trickling filter and NTF processes.
6.1 Combined Carbon Oxidation and Nitrification Combined carbon-oxidizing (i.e., cBOD5 removal) and nitrification trickling filters may be used with synthetic or rock media. The effect of combined carbon oxidation and nitrification in rock-media trickling filters is a result of reduced soluble cBOD5 loading. The design of combined carbon oxidation and nitrification in synthetic media trickling filters is empirical (Parker, 1998). The U.S. Environmental Protection Agency (Washington, D.C.) (U.S EPA, 1991) conducted a survey of 10 combined carbon oxidation and nitrification
115
Biofilm Reactors
facilities. Six of the facilities included the TF/SC process. The survey was used to create empirical guides for achieving nitrification in trickling filters based on fullscale observations. The manual for nitrogen control (U.S. EPA, 1993) presented recommended BOD 5 loadings, which also are presented later, to achieve both carbon oxidation and nitrification in a single-stage trickling filter. The kinetics of BOD5 and NH3-N removal in a single-stage trickling filter are complex, and there is a general lack of research describing combined carbon oxidation and nitrification trickling filter mechanics. Therefore, empirical design criteria are described in this section. U.S. EPA (1975) summarized full‑ and pilot‑scale rock-media trickling filter performance data from Lakefield, Minnesota; Allentown, Pennsylvania; Gainesville, Florida; Corvallis, Oregon; Fitchburg, Massachusetts; Ft. Benjamin Harrison, Indiana; Johannesburg, South Africa; and Salford, United Kingdom. Figure 3.9 (U.S. EPA, 1975) was created using the pilot- and full-scale rock-media combined carbon oxidation and nitrification trickling filter data and illustrates a relationship between BOD5 volumetric loading and nitrification efficiency. These observations
100 No Recirculation
80 Nitrification Efficiency, %
116
Recirculation
60
40
20
0
0
10
20
30
40
50
60
BOD5 Load, 1b/1000 cu ft/day
Figure 3.9 Nitrification efficiency versus organic loading in rock-media trickling filters (U.S. EPA, 1975).
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
indicate that an organic loading rate of 0.08 kg BOD5/m3·d (5 lb BOD5/d/1000 cu ft) is required for rock media to achieve approximately 90% nitrification. Recirculation typically improved nitrification, particularly for nitrification efficiencies greater than 50%. Stenquist et al. (1974) related organic loading to nitrification efficiency in a pilotscale combined carbon oxidation and nitrification synthetic-media trickling filter. The researchers observed 89% ammonia-nitrogen removal at an organic loading of 0.36 kg/m3∙d (22 lb/d/1000 cu ft). The nitrification capacity of the trickling filter was reported a function of surface BOD5 loading (kg BOD5/ m2∙d of trickling filter media). Figure 3.10 illustrates results from the studies at Stockton, California (fulland pilot-scale). Daigger et al. (1994) presented an evaluation of three full-scale cross-flow media trickling filters (each having a 100-m2/m3 specific surface area), which were dosed with rotary distributors and designed for combined carbon oxidation and nitrification. The data collected from these installations are presented in Figure 3.11.
100
++ ++ ++ + ++ + + + + + + + + +
Nitrification, %
80
60
+
40
+
+ + + + +
+
+ +
20
0
0
10 20 Effluent Soluble BODsi mg/L
30
Figure 3.10 Nitrification efficiency as a function of filtered, or soluble, BOD5 in the effluent of a vertical-flow media trickling filter at Stockton, California (U.S. EPA, 1975).
117
Biofilm Reactors
100 Rowlett Creek Duck Creek Missouri River
90 80 Nitritication Efficiency (%)
118
Literature Curve for 100 m2/m3 Media (EPA, 1975)
70 60 50 40 30 20 10
0
0.6 0.8 0.2 0.4 Organic Loading (Kg BOD5/m3-day)
1.0
Figure 3.11 Nitrification efficiency versus organic loading in trickling filters (Daigger et al., 1994).
The data suggest that an organic loading less than 0.24 kg BOD5/m3·d (15 lb BOD5/ d/1000 cu ft) is required to achieve 90% ammonia-nitrogen conversion in crossflow media trickling filters. Similar to the observations of Stenquist et al. (1974), synthetic-media trickling filters (specifically vertical- or cross-flow media) are able to meet ammonia-nitrogen conversions in excess of 89% when the BOD5 load is 3 to 4 times greater than that acceptable for rock-media combined carbon oxidation and nitrification trickling filters. Parker and Richards (1986) described test results obtained from Garland, Texas, and Atlanta, Georgia. The results, reported on the basis of average surface loadings, are shown in Figure 3.12. Parker and Richards (1986) compared combined carbon oxidation and nitrification in rock- and synthetic-media trickling filters. When compared on a surface-area-loading basis, their performance is similar.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Nitrification efficiency, %
A
100
0
kg/1000 m2·d 10 15
5
Key Media
80
Cross-Flow - 60° Cross-Flow - 45° Vertical
60
20
Depth, ft 10 20
Curve for rock media 40
20
0
0
1.0
2.0
3.0
4.0
5.0
Organic loading, ib BOD5/1000 sq ft/day
kg/1000 m2·d B
Nitrification efficiency, %
100
0
5
10
20
Cross-fow media, garland pilot study Vertical media, garland pilot study
80 60 40 20 0
15
Key
Curve for rock media and garland cross-flow media
Curve for garland vertical media 0
1.0
2.0
3.0
4.0
Organic loading, ib BOD5/1000 sq ft/day
Figure 3.12 Nitrification efficiency versus organic loading in trickling filters (Parker and Richards, 1986). To produce the lowest NH3-N concentrations, biofilm thickness control is recommended to prevent sloughing cycles and macrofauna accumulation by
(1) Constant hydraulic application (dosing), (2) Periodic high-intensity hydraulic application (flushing), (3) Flooding (if the containment structure has been designed to retain water), or (4) A combination of these practices.
119
120
Biofilm Reactors
Daigger et al. (1994) proposed eq 3.20 to describe BOD5 and NH3-N removal in combined carbon oxidation and nitrification trickling filters.
Q VOR = Si + 4.6 ⋅ SNOx -N ⋅ VM
(3.20)
Where VOR = volumetric oxidation rate (kg/m3·d); Si = BOD5 concentration in the influent stream (g/m3); SNOx-N = nitrate/nitrite-nitrogen concentration in the effluent stream (g/m3); Q = flowrate, including recirculation streams (m3/d); and VM = synthetic media volume (m3). Using eq 3.20, Daigger et al. (1994) reported the VOR for three combined carbon oxidation and nitrification trickling filter (with plastic media) processes in the range 0.4 to 1.3 kg/m3·d.
6.2 Nitrifying Trickling Filters Nitrifying trickling filters are a reliable and cost-effective method of NH 3-N conversion. The following design practices have been demonstrated in full-scale application:
(1) Use medium-density cross-flow media to optimize hydraulic distribution and oxygenation; (2) Use mechanical ventilation; (3) Periodically alternate the lead NTF to avoid patchy biofilm development in the lower reaches of the second-stage unit; (4) The influent should be secondary effluent, to minimize competition between bacterial species for substrates inside the biofilm; (5) Maximize wetting efficiency to avoid the formation of dry spots; (6) Dose the NTF at a rate that will minimize the accumulation of macrofauna; and (7) Equalize NH3-N-laden supernatant from solids processing operations to even out diurnal load variability (Parker et al., 1995, 1997).
Benefits to NTFs include low energy consumption, stability, operational simplicity, and reduced sludge yield. The reduced sludge yield and resulting low TSS concentration in the NTF effluent stream has led some units to be constructed without downstream liquid–solid separation units. This is dependent on site-specific treatment objectives and effluent water quality standards. An operational issue that
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
121
can be detrimental to process performance is the control of predatory macrofauna. Therefore, the designer must include means for managing solids and macrofauna laden water resulting from macrofauna control measures. Design and operational features dedicated to macrofauna control are presented in a subsequent section. Nitrifying trickling filters having 6‑ to 12.2‑m (20‑ to 40‑ft) modular plastic media depths has demonstrated improved performance. The NTFs have been constructed with depths up to 12.8 m (42 ft). Shallower units can operate as a two-stage system. Recirculation should be minimized to that required for biofilm thickness control to maximize the NH3-N concentration (i.e., maintain a high driving force). The practice of alternating the lead trickling filter in a two-stage NTF system is referred to as alternating double filtration (ADF). Gujer and Boller (1986) and Parker et al. (1989) observed patchy biofilm growth in the lower section of pilot-scale NTFs. The researchers attributed the patchy growth to dry spots. Aspegren (1992) observed improved nitrification and reduced biofilm patchiness when using ADF. Use of the ADF approach with trickling filters in series promotes full-depth biofilm development in both trickling filters. The lead trickling filter should be switched every 3 to 7 days. The primary drawback of ADF is an increase in power requirements, which may be in excess of 50%, as a result of double pumping. In addition to increased operating cost, capital costs associated with pipes and valves also will increase cost. Parker (1998, 1999) described nitrification efficiency differences in NTFs containing cross-flow and vertical-flow modular synthetic media. Table 3.4 summarizes the observations, which demonstrate that zero-order ammonia-nitrogen fluxes are Table 3.4 Reported zero-order nitrification rates for vertical- and cross-flow media (after Parker, 1998, 1999). Media type
0
Location
Reference
J N (g/m2·d) Temperature range (°C)
Central Valley, Utah
Parker et al. (1989)
XF 140
2.3–3.2
11 to 20
Malmö, Sweden
Parker et al. (1995)
XF 140
1.6–2.8
13 to 20
Littleton/Englewood, Colorado Parker et al. (1997)
XF 140
1.7–2.3
15 to 20
Midland, Michigan
Duddles et al. (1974)
VF 89*
0.9–1.2
7 to 13
Lima, Ohio
Okey and Albertson (1989b)
VF 89*
1.2–1.8
18 to 22
Bloom Township, Illinois
Baxter and Woodman (1973)
VF 89*
1.1–1.2
17 to 20
*Fully corrugated.
Biofilm Reactors
5 4.5 Nitrification rate, gN/m2*d
122
Influent SS > 15 mg/l Influent SS < 15 mg/l
4 3.5 3
SS < 15 mg/l
2.5 2
SS > 15 mg/l
1.5 1 0.5 0
Ammonia concentration, mg N/l 0
3
6
9
12
15
16
21
24
27
30
Figure 3.13 Effect of bulk liquid TSS concentration on nitrification in a pilot-scale NTF (Andersson et al., 1994). greater for cross-flow than vertical-flow media. A factor contributing to the enhanced performance may be improved oxygenation resulting from the increased number of interruption points in cross-flow media (Gujer and Boller, 1986; Parker et al., 1989). Autotrophic nitrifying biofilms are thin when compared with the heterotrophic biofilms, which are primarily responsible for BOD5 removal; therefore, denser cross-flow media can be used in a NTF. The concentration of TSS in the secondary effluent has been shown to have an effect on NTFs (Boller et al., 1990; Parker et al., 1989). Andersson et al. (1994) demonstrated (see Figure 3.13) that the maximum zero-order nitrification rate in a pilotscale NTF apparently decreased from approximately 2.6 g N/m2·d to approximately 1.8 g N/m2·d. Three NTF design models are described in this section. They include (1) the model of Gujer and Boller (1986), (2) the model of Gujer and Boller (1986) with the modification of Parker et al. (1989), and (3) the model of Okey and Albertson (1989b).
6.2.1 Gujer and Boller Nitrifying Trickling Filter Model Gujer and Boller (1986) developed a semi-empirical NTF model, as shown in eq 3.21. J N (S, T ) = J N ,max (T ) ⋅
SB,N K N + SB,N
(3.21)
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Where J N ( S, T ) = NH3-N flux at SB,N (g/m2·d) and T; J N, max (T ) = maximum NH3-N flux at temperature T (g/m2·d); = JO2 ,max ( T ) 4.3 ; SB,N = bulk liquid NH3-N concentration (g/m3); KN = NH3-N half-saturation concentration (g N/m3) (1.0 to 1.5, typical value 1.0); and T = temperature (°C). Based on a “line-fit” relationship, the flux at any depth in the NTF can be calculated, as J N ( z , T ) = J N ( 0 , T ) ⋅ e − k ⋅z . Two solutions were developed to account for a change in the rate of nitrification with NTF depth (k ≠ 0) (eq 3.22), and the second assumes no decrease in the rate of nitrification with NTF depth (k = 0) (eq 3.23). a ⋅ J N ,max ( T )
k ⋅ vh
Sin , N ⋅ ( 1 − e − k⋅ z ) = Sin , N − SB , N + K N ⋅ ln SB , N
(3.22)
And, when k = 0, z ⋅ a ⋅ J N ,max ( T )
vh
Sin , N = Sin , N − SB , N + K N ⋅ ln SB , N
(3.23)
Where a = specific surface area (m2/m3); k = empirical parameter rate decrease (1/m) (= 0 to 0.16, typical value 0.1); vh = hydraulic load (with or without recirculation) (m3/ m2·d); z = NTF depth (m); and Sin,N = NH3-N concentration in influent stream (g/m3). These equations can be solved directly to size an NTF for a desired SB,N. When recirculation is used, an iterative solution routine, which includes eq 3.24, is required because of the effect recirculation has on both vh and Sin,N. SN , i =
S0 , N + R ⋅ S B , N 1+ R
or R =
S0 , N − Sin , N Sin , N − SB , N
(3.24)
where S0,N = ammonia-nitrogen concentration (g/m3) in the influent stream before being mixed with the recirculation stream.
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Biofilm Reactors
The ammonia-nitrogen concentration in NTF influent stream, Sin,N, will be less than S0,N when recirculation is applied. Parker et al. (1989) proposed a modification of this model to account for oxygentransfer efficiency variability among modular plastic media types and operating conditions. The revised expression is as follows:
Where
J N ( z , T ) = EO2 ⋅
JO2 ,max ( T ) 4.3
⋅
SB , N K N + SB , N
(3.25)
EO2 = the dimensionless NTF media effectiveness factor and JO2 ,max ( T ) = maximum dissolved oxygen flux at temperature, T (g/m2·d). Gujer and Boller (1986) reported, based on their experience, an EO2 value in the range 0.93 to 0.96 for KS,O2 = 0.2 g O2 /m3 and the temperature range 5 to 25°C. Parker et al. (1989), on the other hand, observed lower EO2 values (in the range 0.7 to 1.0) and claimed that a departure from EO2 = 1.0 accounts for wetting inefficiency, biofilm grazing by macrofauna, or competition for dissolved oxygen between autotrophic nitrifying and heterotroph bacteria inside the biofilm. The researchers recommended that medium-density cross-flow media is used in NTF applications and that EO2 may range from 0.7 to 1.0 for such media. High-density cross-flow media had a corresponding EO2 approximately equal to 0.4 (Parker et al., 1995). Readers should note, however, that the basis for these observations was pilot-scale NTFs that continuously were fed influent wastewater through a nozzle-type system. No research exists describing the effect of rotary distributors and intermittent hydraulic application of influent wastewater on high-density media NTF performance. According to Parker et al. (1995), EO2 ⋅ JO2 ,max ( T ) 4.3 is the zero-order ammonianitrogen flux. The maximum dissolved oxygen flux reflects the oxygen-transfer efficiency of the selected modular plastic media, and was determined by the researchers with the TRIFL (Logan et al., 1987b). The coefficient KS,O2 determined for the Central Valley WWTP, Utah, was between 1 and 2 mg/L (Parker et al., 1989).
6.2.2 Okey and Albertson Nitrifying Trickling Filter Model Okey and Albertson (1989b) summarized nitrification rates observed from five different NTF facilities. These data were not corrected for temperature, and each of the trickling filters relied on natural draft ventilation; therefore, variability in the data may be the result of assignable causes also including wastewater characterization, media type, and hydraulic application rate. Okey and Albertson (1989b) suggest that
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
the NH3-N flux will approach 100% removal for loadings less than 1.2 g/m2·d. The empirical design procedure proposed by Albertson and Okey (1988) can be summarized as the sum of the medium-density NTF media for the zero- and first-order regions. The design procedure includes the following steps:
(1) Using eq 3.26, determine the trickling filter media volume based on zero‑ order kinetics using medium density (138‑m2/m3) media and an NH3-N flux (JN) of 1.2 g/m2·d over a temperature range of 10 to 30°C. Below 10°C, adjust the rate using Θ = 1.045(T-10). (2) Determine trickling filter media volume based on first-order kinetics using a rate (J’N), which equals the following formulation and does not have a temperature correction between 7 and 30°C: J = J ’ N
avg N
SN , e ⋅ SN , TRaN
0.75
g SN , e = 1.2 ⋅ d ⋅ m2 SN , TRaN
0.75
(3.26)
where SN,TRAN = a transition NH3-N concentration (mg/L), which can be determined from Figure 3.14. This concentration is dependent on the degree of oxygen saturation and temperature.
Ammonium-N Concentration, mg/L
10 9 8 7 6
Zero-order Region
5
Dissolved Oxygen 100% Saturation 75% Saturation 50% Saturation 25% Saturation
4 3 2 1 0
0
5
10
15
20
25
30
35
40
Temperature, °C
Figure 3.14 Transition NH3-N concentrations as a function of temperature (Okey and Albertson, 1989b).
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126
Biofilm Reactors
The designer can determine the total trickling filter media volume by adding the volume required for both the zero- and first-order kinetic realms. The above design procedure stipulates that the following conditions are met: • Ratio of influent BOD5 to TKN concentrations ≤1.0, • Filtered BOD5 concentration ≤12 mg/L, • Q(1 + R) ≥ 0.54 L/m2·s (0.8 gpm/sq ft), • Carbonaceous BOD5 and TSS concentrations ≤30 mg/L for medium-density (138 m2/m3 [42 sq ft/cu ft]) synthetic media, • Mechanical ventilation is used, and • Distributor speed control to provide dosing rate in the range 25 to 75 mm/ pass and flushing intensity ≥300 mm/pass. The following illustrates application of the Gujer and Boller (1986) and Okey and Albertson (1989a) model. The following parameters were applied to define the NTF height required to process 219 530 m3/d (58 mgd) with a 9.4-mg NH3-N/L concentration (at T = 16°C) to meet the treatment object of 1.0 mg NH3-N/L remaining in the effluent stream: • Specific surface area, a = 138 m2/m3 (medium-density cross-flow modular plastic media); • Maximum ammonia-nitrogen flux at temperature T, JN,max = 1.1 g N/m 2·d (Gujer and Boller, 1986); • Total hydraulic load, vh = 65 m/d (35.2 to 88.0 m/d according to WEF, 2009); • Empirical nitrification rate decrease, k = 0 m/d (0 to 0.16, 0.11 typical when k > 0); and • Ammonia-nitrogen half-saturation coefficient, KN = 1.0 g N/m3 (Gujer and Boller, 1986).
6.2.2.1 Application of the Gujer and Boller Model Because the ammonia-nitrogen concentration is not reduced by a recirculation stream, and the influent ammonia-nitrogen concentration indicates that virtually no nitrification occurred in the secondary process, a high k-value is not expected, because the increased driving force, or bulk-liquid ammonia-nitrogen concentration, will promote oxygen rate-limiting conditions. Consequently, a zero-order ammonia-nitrogen
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
flux is expected. This condition is commensurate with a constant ammonia-nitrogen flux (i.e., the rate does not change, or k = 0). Therefore, rearranging eq 3.23, the following can be used to describe a total NTF height required to meet the treatment objective: Sin,N ⋅ Sin,N − SB,N + K N ⋅ ln a ⋅ J N,max (T ) SB,N g m 9.4 3 65 g g g hr m = ⋅ 9.4 3 − 1.0 3 + 1.0 3 ⋅ ln g gN m2 m m m 1.0 138 3 ⋅ 1.1 2 m 3 m m ⋅ d
Z=
vh
= 11.4 m In the event that k ≠ 0, eq 3.22 would be rearranged as follows:
1 Z = − ⋅ ln 1 − k
Sin,N k ⋅ vh ⋅ Sin,N − SB,N + K N ⋅ ln SB,N a ⋅ J N,max (T )
The total NTF cross-sectional area can be calculated based on the desired total hydraulic load, which, in this case, is assumed to be 80 m/d. area =
Q vh
m3 d = m 80 d = 3427 m 2 22 785
(3.27)
Next, the process designer must determine the number of NTFs and their dimensions. Assuming that the treatment objective can be met with eight equally sized NTFs operating in four parallel trains, the process will function as a two-stage system (i.e., NTFs in series). Then, the area of each NTF can be calculated as follows:
127
128
Biofilm Reactors
area =
Atotal N trains
3 427 m 2 4 = 857 m 2
(3.28)
=
Consequently, the area for each of the eight NTFs is 857 m2. Each of the circular NTFs will have a diameter calculated as follows: d=
or d=
π ⋅ d2 4
a⋅ 4 π
857 m 2 ⋅ 4 = π = 33 m (109 ft)
(3.29)
Because each of the four trains consists of two NTFs, the height (H) of each NTF can be calculated as follows: H = 11.4 m ÷ 2 = 5.7 m (18.75 ft) The media volume can then be calculated as follows: 3427 m2 × 11.4 m = 39 000 m3
6.2.2.2 Application of the Albertson and Okey Model The following parameters were applied to the Okey and Albertson (1989a) model, to provide a comparison for the process requirements defined using the Gujer and Boller (1986) model: • Transitional ammonia-nitrogen concentration in NTF, S N,TRAN = 3.0 g/m 3 (Figure 3.14, T = 16°C, and 75% saturation); and • Maximum ammonia-nitrogen flux, JN,max = 1.1 g/m2·d (consistent with the assumptions applied to the analysis above). First, calculate the biofilm area required to account for the zero-order region of the NTFs.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Q ⋅ (SN,i − SN,TRaN )
area =
J N,max
g g m3 ⋅ (24.4 3 − 3.0 3 ) d m m = g 1.1 2 m ⋅d = 4 330 078 m 2 222 785
Then, applying eq 3.26, the first-order flux can be calculated as follows:
J = J ’ N
avg N
SN,e ⋅ SN,TRaN
0.75
g 1.0 3 g m = 1.1 2 ⋅ g m ⋅d 3.0 3 m g = 0.48 2 m ⋅d
0.75
The biofilm area required to account for the first-order region of the NTF can be calculated as follows:
area =
Q ⋅ (SN,i − SN,TRaN ) J N,max
g g m3 ⋅ (3.0 3 − 1.0 3 ) d m m = g 0.48 2 m ⋅d = 923 346 m 2 222 785
129
130
Biofilm Reactors
Finally, the total biofilm area required in the NTF can be calculated as follows: area = area0− ord + area1st − order = 4 330 078 m 2 + 923 346 m 2
= 5 253 4224 m 2
The total NTF media volume can be calculated as follows: 5 252 424 m2 ÷ 138 m2/m2 = 38 068 m3 (1 343 811 cu ft). The surface area is fixed at 3427 m2, to maintain an 80-m/h total hydraulic load. In addition, the eight proposed NTFs will be operated as a two-stage process. Therefore, by fixing the total NTF surface area and assuming the use of four two-stage NTF trains allows for the determination of a total height required to provide the abovecalculated media volume: Height =
=
Volume area 38 068 m 3 3 427 m 2
(3.30)
= 11.1 m The height (H) of each NTF can be calculated as follows: H = 11.1 m ÷ 2 = 5.55 m (18.3 ft) Wall et al. (2001) compared the design procedures of Gujer and Boller (1986) and Okey and Albertson (1989a). Wall et al. (2001) concluded that both models generally provide a good description of NTF performance under average NH3-N loading conditions. However, the models generally showed more significant peaks and troughs than the samples indicated. The Gujer and Boller (1986) model predicted peaks that were more exaggerated than the Okey and Albertson (1989b) model. No justification was presented for the models’ inability to account for peak NH3-N loading conditions. Parker et al. (1995) demonstrated that the modified Boller and Gujer (1986) model (eq 3.25) effectively predicted NTF effluent NH3-N loading concentrations under both average and peak conditions. Example results of the study performed by Parker et al. (1995) are illustrated in Figure 3.15.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
20
Ammonia - N, mg/l
Influent
10 Actual Effluent
0 25
31 October
5
10
15 November 1992
20
Predicted effluent
25
30
Figure 3.15 Actual and predicted effluent from an NTF (Parker et al., 1995). Predicted effluent was calculated using the modified Gujer and Boller (1986) model.
6.3 Temperature and Hydraulic Application Effects Two processes—mass transfer and biochemical conversion—are characteristic of all biofilm reactors and influence biofilm structure, function, and ultimately reactor performance. Mass transport of soluble substrates inside the biofilm is controlled by molecular diffusion, which is typically slow when compared with biochemical conversion. The resistance to mass transfer inside the biofilm results is strong soluble substrate concentration gradients (typically normal to the biofilm growth medium). The concentration gradients exist in the bulk liquid, near the biofilm surface, and inside the biofilm. Mass transport is the primary mechanistic difference between biofilm reactors, such as the trickling filter and suspended-growth reactors. Typically, full-scale operating suspended-growth systems are kinetically (i.e., biomass) limited, whereas biofilm reactors are diffusion (i.e., surface-area) -limited (Boltz and Daigger, 2010). Cold wastewater increases solubility (thereby increasing the rate of molecular diffusion), but a reduction in temperature also retards the rate of biochemical transformation. Temperature will have a substantial effect on trickling filter efficiency only if the biochemical transformation rate is less than the rate of molecular diffusion. Otherwise, soluble substrates simply will penetrate the biofilm deeper and support active biomass in a thicker (and/or denser) biofilm.
131
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Biofilm Reactors
Benzie et al. (1963) observed that excessive cooling has a negative effect on BOD5 removal efficiency in carbon-oxidizing trickling filters, and Parker et al. (1990) observed a temperature effect on NTF performance. Benzie et al. (1963) reported that there was a significant difference in trickling filter BOD5 removal efficiency between the summer (mean temperature −4 to −1°C) and winter (mean temperature 19 to 23°C) months. In the trickling filter processes studied, where recirculation was practiced, a total seasonal variation of 35 to 40% in trickling filter BOD5 removal efficiency was reported, while only a 10% reduction in BOD5 removal efficiency was reported in a facility that did not recirculate (secondary clarifier effluent). The researchers concluded that a primary contributor to the marked reduction in performance during the winter months was the substantial cooling of wastewater resulting from recirculation. Parker et al. (1990) reported a temperature effect on NH3-N flux in NTFs for temperatures in the range 10 to 22°C. A key to good trickling filter performance is limiting temperature changes to a range over which temperature has little effect and minimizing heat loss during operation. Methods for reducing heat loss include the following: • Construct deep trickling filters with reduced cross-sectional area, to increase the WHL and reduce the recirculation flow; • Extend the containment structure upward so that it is 1.5 to 2.0 m above the distributor; • Cover the trickling filter; and • Provide a mechanical aeration system to control airflow (Grady et al., 1999). If a mechanical aeration system is not feasible, natural airflow can be controlled during the winter months with adjustable dampers on the vents and other air inlets. Optimal hydraulic requirements for promoting maximum nitrification rates are unknown. Gullicks and Cleasby (1986, 1990) and Okey and Albertson (1989b) presented data from studies indicating that increasing the application rate (L/m2·s) increased the rate of NH3-N oxidation. According to the researchers, hydraulic application rates greater than 1 L/m2·s (1.5 gpm/sq ft) produced the best results.
7.0 Design Considerations The following sections are related to equipment selection, specification, and trickling filter construction. The information described here is a guide, but it is recommended
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
that the final design result from the collaboration of discipline-specific specialists (e.g., mechanical, electrical, and structural engineers) and manufacturers’ technological specialists to completely define and evaluate a system design, interfacing, material integrity, and structural integrity test methods/tolerances.
7.1 Distribution System Methods of supplying wastewater to the trickling filter distributor include gravity feed, dosing siphons, and pumping. The type of conveyance process selected depends on the hydraulic gradient available and the distributor. Trickling filter distributors require piping between these conveyance systems and the trickling filter distribution system. Where the trickling filter is not designed for continuous dosing, a pump or dosing tank and siphon may precede the distribution system. Any type of trickling filter distribution system requires only a single conduit to convey wastewater. The hydraulic design of this conduit conforms to that for other WWTP process piping. Design variables include headloss, provision for drainage, structural considerations, and protection against applicable climatic and corrosive conditions. Flow distribution is an important feature in a trickling filter system. Wastewater influent to the bioreactor must be applied evenly at a rate that keeps the trickling filter media wet and prevents ponding and clogging. Uneven wastewater distribution and/or a dosing rate that is insufficient for biofilm thickness control will result in poor trickling filter performance. Odors will emanate as solids accumulate in the trickling filter, and biomass near the biofilm–growth-medium interface becomes anaerobic. Furthermore, insufficient biofilm thickness control promotes the accumulation of macrofauna. It is recommended that new trickling filters are constructed as circular tanks to accommodate the use of a rotary distributor. A rotary distributor may be equipped with 2 to 6 arms, but units with 4 arms are typical. The distributed flow may be staggered for full coverage per arm. That is, each arm may provide 50 or 100% coverage per revolution. Nozzles and dosing siphons are not recommended for systems treating municipal wastewater. Rotary distributors are the focus of material presented in this section. If a rectangular trickling filter is upgraded with a rotary distributor, special provisions should be made for wetting media that is outside of the distributor diameter. Otherwise, this media should be removed, or it will provide a breeding area for macrofauna. The need for and benefits of providing rotary distributor speed control has been discussed in the Distribution System section.
133
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Biofilm Reactors
7.1.1 Hydraulic Drive Rotary Distributors Traditionally, hydraulically driven rotary distributor speed control is mediated by pumped water discharging from orifices situated along the trailing side of the distributor arms, as illustrated in Figure 3.16. Historically, little attention has been given (a)
Diameter
Thrust Bearing Stabilizing Bearing Support Cage
DC Drive 0.25 hp Center Mast Tie Rods
Barrel
Turnbuckle
Flush Gate
Tank wall
Square Arm Filter Media
Spreaders and Orifices
Braking Jet Spreader Locations
(b) Slip ring assembly Rotary distributor mechanism
Valve gate positioner typical (4) arms
Center cage
Air line drip leg typical (4) arms
Tie rods
Vent pipes Center wall
Fixed mast
Anchor bolts Center column
Influent feed pipe Influent flow
Tank wall Trickling filter media
Figure 3.16 (a) Typical hydraulically driven rotary distributor with reversethrusting jets and electric drive, and (b) modern hydraulically driven rotary distributor with variable-frequency controlled gates that close or open orifices (therefore, the flowrate contributes to reverse thrust) (Courtesy of WesTech HydroDoc™).
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
to the rotational speed of trickling filter distribution systems. As a result, typical dosing rates for these hydraulically driven rotary distributors were in the range 2 to 10 mm/pass at 0.2 to 1.5 rpm. These dosing rates are significantly less than those listed in Table 3.3. No minimum speed has been specified for a hydraulically driven rotary distributor. Trickling filters with hydraulically driven rotary distributors have reported stalling speeds varying from 4 to 20 min/rev. The modern hydraulically driven rotary distributor (e.g., HydroDoc [WesTech, Salt Lake City, Utah]) provides means for controlling the rotary distributor rotational speed using energy provided by the pumped influent wastewater. While traditional hydraulically driven distributors are rotated by the thrust of flow expelling flow from the back side of the distributor arms, modern hydraulically driven rotary distributors provide means to proportionally distribute flow between orifices in the front and rear of the distributor arm. Proportioned directional flow control provides rotational speed and, consequently, dosing rate control. Discharge orifices have spreaders that are placed along the front and rear walls of each arm. The orifices extend approximately 30% of the arm length. In front of each orifice row, a pivoting gate is placed against the wall inside the distributor arm. The front and rear gates are raised simultaneously and lowered in opposite directions (i.e., as the front gate closes, the rear gate opens). A gate linkage is designed to position the front and rear gates at their respective port centers at any point in time, which maintains consistent media coverage by maintaining the equivalent of full orifice area resulting from the sum of orifices along the front and back distributor arm. The gate is positioned by an air cylinder fit with a pneumatic controller that receives a 21- to 103-kPa (3- to 15-psi) signal, which represents a desired cylinder extension. The pneumatic position controller senses the degree of cylinder rod extension and directs air pressure to the appropriate end of the cylinder. The induced air pressure causes the cylinder to extend or retract, moving the sliding gates and controlling hydraulic thrust. When in position, pressure is applied to both sides of the cylinder to maintain the gate position indicated by the signaled air value. Control of the modern hydraulically driven rotary distributor consists of a programmable logic controller (PLC) or supervisory control and data acquisition (SCADA) system, which provides comparative monitoring between a feedback signal that describes the actual distributor arm rotational speed and an operator set point value defining the desired rotational speed. Based on the signal comparison results, a 4- to 20-mA signal is directed to the rotary distributor. This signal adjusts gate positioning, which either increases or decreases the rotational speed. The 4- to
135
136
Biofilm Reactors
20-mA signal from the PLC enters the system at the mast base and transmits through a conduit, which is positioned through the slip-ring assembly located at the top of the rotary distributor mast. The signal then passes downward through the air panel, where it is received by a current/power (I/P) transducer (which varies pressure and voltage). The I/P transducer converts the 4- to 20-mA current signal into a 21- to 103kPa (3- to 15-psi) pressure signal piped to the gate position controller located on each arm. Thereby, the I/P signal results in appropriate gate positioning. A feedback signal, indicating rotary distributor arm speed, originates from a rotary encoder, which is mounted below the slip-ring assembly (at the top of the mast). The encoder directs a signal that identifies (1) arm rotational speed and (2) rotational direction. The benefits to this mechanism include rotary distributor arm speed control and a zero-speed, which allows the operator to stop the mechanism for emergency shut down or maintenance (without waiting for the distributor arm to gradually decrease speed before stopping). Recommended operator-controlled PLC functions include the following:
(1) Local start/stop switch, (2) Speed display, (3) Zero-speed warning light and audio alert, and (4) Operator set point rotary distributor rotational speed.
Recommended algorithm-controlled functions include (1) PLC-based rotational speed settings predefined by the operator; or (2) programmed set points, which include the following: • Flushing start time = 0 to 23 hours, 0 to 59 minutes (hh:mm); • Flushing duration period = 0 to 23 hours, 0 to 59 minutes (hh:mm); • Normal operation rotary distributor rotational speed = 0 to 100% maximum speed; and • Flushing operation rotary distributor rotational speed = 0 to 100% maximum speed.
7.1.2 Electronic or Mechanical Drive Rotary Distributors Electrically driven rotary distributors may have a center or peripheral drive. Such an apparatus typically can be retrofitted easily at low cost. Like the PLC hydraulically driven rotary distributors, these units can be programmed to operate at varying dosing rates, as required, to optimize carbon oxidation, nitrification, and biofilm thickness control (which minimizes the accumulation of macrofauna and minimizes the odor
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
potential from anaerobic biofilms). The center-driven unit will be anchored to non-rotating parts of the influent structure, as shown in Figure 3.17. Where no upper steady bearing exists, bearing support must be installed with a stationary shaft, to provide a platform for the drive unit. This can be located in the mast support for the arm guy wires. Where an upper steady bearing does exist, the stationary shaft into this bearing assembly can be extended to support the drive assembly. A peripherally mounted electric drive can be used instead of a center drive. The traction drive can use the inside or
(a)
(c)
(b)
(d)
Figure 3.17 (a) Modern hydraulically driven rotary distributor, (b) gate controller, (c) orifices, and (d) full-length arm and distributor.
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top of the wall. By spring-loading the drive wheel, it can operate with wall irregularities. A rotary union is used to transmit power. The arrangement is similar to that of a traction drive clarifier. It should be emphasized that hydraulic motors with equally wide speed ranges may be used instead of electrically driven units. Electrically driven rotary distributors with remote variable‑speed controllers and timers can operate independently of influent flow. This is may be advantageous for WWTPs without recycle flow or with sufficient recycle flow to minimize the rotational speed. However, the use of speed-controlled rotary distributors almost never replaces the need for a recirculation pumping station. The designer must carefully evaluate minimum flow conditions. Recirculation is typically applied, and the rotary distributor speed control is provided to increase operator flexibility, to optimize the trickling filter process performance.
7.1.3 Optimizing Rotary Distributor Operation Trickling filter optimization can be determined by adjusting the dosing rate and recording effluent water quality (after the system has reached a quasi-steady-state, with respect to effluent water quality). Flushing should be performed during the optimization process. The flushing ideally is conducted during low flow and loading periods, which likely occur between 1:00 a.m. and 6:00 a.m. If flushing has not been practiced, it may take several weeks of daily flushing to achieve proper biofilm thickness control. The optimum dosing rate can be determined by simultaneously operating trickling filters at different operating dosing rates, recording total and soluble BOD5 removal (and ammonia-nitrogen, if optimizing a carbon-oxidizing and nitrification trickling filter), and adjusting the individual trickling filter distribution speed accordingly. Daily high‑intensity flushing routines also can be programmed for the units, to define optimum flushing dosing rate and durations, to maximize trickling filter process performance. Once the optimum dosing and flushing rates have been defined, macrofauna accumulation will be minimized, and BOD5 removal and ammonia-nitrogen (if applicable) removal will be maximized.
7.2 Construction of Rotary Distributors Rotary distributor arms may be tubular or rectangular. Galvanized steel and aluminum are the most common construction materials. However, stainless steel may be used in corrosive environments. Influent wastewater distribution is accomplished with a series of orifices that are positioned along the distributor arm and cover 30 to 100% of the trickling filter (plan) area per pass. These orifices are equipped with manually or PLC
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
controlled sliding gates, and each orifice typically is equipped with splash plates, to further promote wastewater distribution. In many cases, distributors may be equipped with four arms in a high–low flow arrangement. Essentially, two of the arms operate at flows up to and slightly above the average design flow, while the remaining two arms operate only during peak flows. This generally is achieved by baffles that are inside the distributor arm and are near the trickling influent feed pipe. This arrangement maximizes wastewater distribution and flushing intensity, if practiced. The hydraulic head required to drive a distributor and provide influent wastewater distribution typically is in the range 410 to 1000 mm (16 to 40 in.) of water column. The head required for minimum flow typically is in the range 300 to 610 mm (12 to 24 in.) above the center line of the rotary distributor arm orifices. Somewhat greater head is needed to accommodate wide flow ranges. For some distributors, this head requirement has been reduced by the use of overflow devices, which result in dosing through additional arms during periods of high flow. Forcing flow through two arms aids in maintaining the flow through the orifices during minimum flow periods, which is required to maintain adequate wastewater distribution over the trickling filter surface. Trickling filter designs that operate with rotary distributor rotational speeds in the range 8 to 50 min/rev will exert negligible centrifugal force. Nevertheless, rotary distributors require a seal between the fixed influent column and the rotating section. Older designs have various types of water traps, mercury seals, or packed mechanical seals to prevent water from leaking between fixed and rotating parts. Modern designs use one of two types of seals. The first is an overflow arrangement without a lower seal. This type offers the advantage of no seal friction imposed on the mechanism and inherently no seal maintenance; however, the hydraulic head requirement is higher than the modern mechanical seal. Modern mechanical seals have a double neoprene seal with a stainless-steel seal ring; this seal also requires no maintenance and requires less hydraulic head than the aforementioned seal-free design. When older units are upgraded, improvements often include one of these arrangements.
7.3 Trickling Filter Media Selection The design engineer must make an informed decision regarding the selection of media for specific trickling filter applications, including the construction of new facilities and the retrofit of existing facilities. Guidance has been provided earlier in this chapter, but from a process performance perspective. This section describes mechanical considerations.
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Drury et al. (1986) demonstrated that the increased specific surface area, ability to apply increased hydraulic loading, and capacity for improved biofilm thickness control make the replacement of rock media in existing trickling filters with modular plastic media advantageous. Because of the robust construction of containment structures that is required to support the structural load imposed by rock media, existing rock-media trickling filter containment structures may have their walls extended vertically. The taller trickling filter equipped with modular plastic media will result in expanded trickling filter capacity. The WWTP capacity is dependent on an evaluation of each unit process capacity, system hydraulics, and process mechanical appurtenances. Therefore, expanded bioreactor capacity does not necessarily imply that the WWTP capacity has been expanded. Replacing existing rock media with synthetic media may be warranted by a variety of conditions, which include odor generation, existing media deterioration, and treatment capacity expansion, while maximizing the use of existing assets in the existing footprint. Many trickling filter media types can be applied in a range of hydraulic and pollutant loading conditions and to a wide variety of wastewaters, including screened and degritted raw wastewater, primary effluent, secondary effluent, or industrial wastewater. Sufficient scientific evidence exists to surmise that cross-flow media produces better effluent water quality when compared with vertical-flow media, when treating wastewaters with low to medium total organic loading (Harrison and Daigger, 1987). However, if the total organic load becomes substantial, biofilm thickness control and other complicating factors lead to deteriorating effluent water quality. Therefore, vertical-flow media is recommended for roughing applications and the treatment of high-strength industrial wastewaters. Parker (1999) suggested that this efficiency change illustrates the switch-over effect (efficiency switch-over from cross-flow to vertical-flow media) at high total organic loads, which is not observed in the other studies. However, detailed studies do not exist that specify for which organic loading the cross-over originates. The previously described importance of adequate media-wetting and the relative inefficiency in media-wetting inherent to the trickling filter process has led some designers to combine modular plastic trickling filter media in a single bioreactor. The effect of combining trickling filter media is such that the cross-flow media, which is efficient in hydraulic distribution, is situated in the upper layers of the trickling filter, and vertical-flow media, which is less prone to excessive biofilm accumulation, constitutes the remainder of the trickling filter. The designer should note that the upper portions of trickling filters generally are subject to the highest organic load.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Numerous reports (Boller and Gujer, 1986; Crine et al., 1990; Gullicks and Cleasby, 1986, 1990; Onda et al., 1968; Parker et al., 1989) indicate that denser trickling filter media (e.g., random and cross-flow) are more prone to solids retention and fouling. Vertical-flow media is recommended for the treatment of high-strength (industrial) wastewaters, roughing applications, and wastewaters receiving only preliminary treatment (i.e., 3-mm fine screens and no primary clarifiers).
7.3.1 Depth Rock-media trickling filters in North America typically are 1 to 2 m (3 to 6 ft) deep, but may be as deep as 2.4 m (8 ft). The depth limitation of rock-media trickling filters typically is associated with inadequate ventilation (by natural draft) and propensity to pond. In Europe, deeper rock-media trickling filters are common; units in Arnheim, Netherlands, were constructed at 4.9 m (16 ft) deep, but were equipped with mechanical ventilation. There is a general absence of data comparing deep, mechanical ventilated rock-media trickling filters with shallow, natural draft rock-media trickling filters. Synthetic-media trickling filters commonly are constructed between 5 and 8 m (16 and 26 ft) deep, although units up to 12.8 m (42 ft) deep exist. Trickling filters in excess of approximately 9 m (30 ft) deep may require special provision to redistribute wastewater before reaching the bioreactor effluent plane. Factors limiting the depth of synthetic-media trickling filters include aesthetics, serviceability, pumping requirements, and structural integrity of the plastic media. Increased trickling filter depth has no implication on biological treatment efficiency. However, increasing trickling filter depth does reduce the minimum flow required to achieve optimal wetting efficiency (i.e., taller trickling filters generally have a smaller diameter, which is more susceptible to efficient wetting). The improved efficiency in wetting reduces recirculation pumping requirements. In taller filters that have high loadings, oxygen deficiency may occur in the uppermost layers. However, adequate ventilation and biofilm control measures can prevent problematic odors. The effect of trickling filter media depth on bioreactor performance has been treated as a matter of controversy in previous design manuals. Several investigators suggest that volume, regardless of depth, controls performance (Bruce and Merkens, 1970, 1973; Galler and Gotaas, 1964; Kincannon and Stover, 1982; National Research Council, 1946). Practical limitations placed on trickling filter performance generally are related to inefficiencies in media-wetting and physical constraints on trickling filter depth. Many of the investigators previously cited in this chapter (regarding the same topic) attributed the improved performance with trickling filter depth to
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improved media-wetting. The average hydraulic rate should exceed 0.5 L/m2·s (0.75 gpm/sq ft) to ensure maximum performance.
7.3.2 Structural Integrity The choice of rock trickling filter media often is governed by the materials available locally or the cost of transportation. Field stone, gravel, broken stone, blast-furnace slag, and anthracite coal have been used for this purpose. Whatever material is chosen, it typically is specified that it should be sound, hard, clean, free of dust, and insoluble in wastewater constituents. There is some difference in opinion as to the optimum size of rock trickling filter media. A common specification requirement is that 95% or more of the media pass 2600-mm2 (4-sq in.) mesh screens and be retained on 1600-mm2 (2.5-sq in.) mesh screens. The pieces typically are specified to be uniform in size, with all three dimensions as close to equal as possible. The material should not disintegrate under service conditions, such as breaking into smaller pieces or crumbling into fine material. Frequently, the material is specified to be substantially sound, as determined by the sodium sulfate soundness test. Specifications for placing rock media frequently include statements such as the following:
(1) When placing trickling filter media, breakage and segregation of differently sized particles must be prevented; (2) Trickling filter media will be screened and cleaned immediately before placing, to eliminate as many fine sediments, or stone fragments, as possible; (3) Trickling filter media will be placed by a method that does not require heavy traffic of any type on the top of those media already placed; and (4) Placing media by means of belt conveyor, wheelbarrow, or bucket crane will be acceptable.
Synthetic trickling filter media, specifically the bundle type (0.61 m × 0.61 m × 1.22 m [2 ft × 2 ft × 4 ft]), are the most common media used in new trickling-filter-based WWTPs. Proper specifying and testing criteria should be used to ensure that the trickling filter media will satisfy duty requirements. Bundle media are manufactured from polyvinyl chloride (PVC), while the random media typically are manufactured from polyethylene or polypropylene. The testing procedures herein apply to bundle media, but it is equally important that the random media have sufficient strength to resist subsidence resulting from the combined weight of the media, water, and biomass. Consideration should be given to the long-term (96-hour) and short-term (typically <2-hour) test and the ability of either test to predict media strength over what
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
will hopefully be a minimum 20-year life. The PVC is suitable as a structural material, as long as deformation or creep loading is not exceeded. The material fails by deformation, which can be a slow process that persists if loading is maintained. New trickling filter media is stronger than 10-year-old trickling filter media, because PVC weakens with time. In addition, the plasticizer dissipates, and the media also becomes brittle. Because of the high initial strength-to-weight ratio, exceptionally thin media can meet inadequately drafted specifications and place long-term useful life in question. Typically, this is a problem of not understanding the relationship between short-term testing results and long-term load capacity. Test temperature is important, because PVC loses strength as the temperature exceeds 18 to 21°C (65 to 70°F). The load testing should be conducted at the maximum water temperature. The database temperature from media suppliers is 23 ± 1°C (73 ± 2°F); however, this 23°C temperature may not satisfy specific duty requirements. Mabbott (1982) not only introduced the short-term compression test to assess media strength, but also reported that the modulus of elasticity (and corresponding media strength) dropped drastically (with increased temperature). When temperatures of the wastewater exceed 30°C (86°F), all structural testing should be conducted at the maximum operating temperature of the media. The designer must carefully consider heat build-up coinciding with trickling filter shutdown and its effect on modular-plasticmedia structural integrity. The issue is amplified when trickling filters are covered by a dome where air may not easily escape the biological reactor. Aerobic biochemical reactions generally proceed within the temperature range 35 to 40°C, which is the upper limit for growth of mesophilic bacteria (Grady et al., 1999). These temperatures may be observed in activated sludge systems, but are not common operational temperatures for trickling filters. The pore (internal biofilm) temperature approaches equilibrium with the air temperature inside the trickling filter, which may be in the range 10 to 30°C. In addition to ambient conditions, the amount of biomass present, biomass condition, and mode of ventilation affect trickling filter internal temperatures (Harrison, 2007), which are important considerations during emergency shutdown or even during the installation of filter media. When procuring media, a good design practice is to specify a service temperature that exceeds the actual water temperature, to provide adequate protection during unplanned conditions. A service temperature of 32 to 49°C (90 to 120°F) would not be unreasonable for warmer climatic conditions, high organic loading rates, or for filters where partial plugging or temperature concerns exist. An alternative may be to provide rotary sprinklers within domed trickling filters for heat dissipation during shutdown periods—emergency or otherwise.
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7.4 Trickling Filter Pumping Station or Dosing Siphon Most trickling filters use recirculation pumps, which typically are constant‑speed, low‑head centrifugal units designed to operate with a total head of the trickling filter media depth, 2.0 to 3.0 m of static head, and friction losses. The VFD-controlled motors currently are typical fixtures on process pumps. Both submerged and non-submerged (dry-pit) vertical pumps have been used extensively. Pump intake screens typically are unnecessary, because the recirculated flow typically is free of clogging solid materials. Hydraulic computations are necessary. Computations for minimum flow are necessary to ensure adequate head to drive hydraulically driven distributors; computations for maximum flow indicate the head required to ensure adequate discharge capacity. The net available head at the horizontal center line of the distributor’s arm and other points may be calculated by deducting the following applicable losses from the available static head: entrance loss; drop in level in the dosing tank, as distributor pipes are filled (only applicable to dosing siphons); friction losses in the piping to the distributor; proper allowance for minor headlosses; headloss through distributor riser and center port; friction loss in distributor arms; and velocity head of discharge through nozzles necessary to start the hydraulically driven rotary distributor. Trickling filter distribution head requirements are set by a system’s manufacturer. Despite the headloss resulting from the trickling filter commonly being the greatest in a given WWTP, power requirements for the trickling filter process (including recirculation pumping and auxiliary powered equipment) typically are less than those for the activated sludge process.
7.5 Control Mechanisms for Trickling Filter Macro Fauna Macrofauna and their larvae graze biofilm. Curds and Hawkes (1975) reported that grazing activity in full-scale trickling filters treating municipal and industrial wastewaters was affected by the nature of the biofilm and operational variables. Possible benefits associated with the presence of macrofauna include reduced sludge production, improved sludge settlability, and biofilm thickness control. Williams and Taylor (1968) conducted laboratory-scale experiments demonstrating that macrofauna-free systems transformed only 40% of the available organic carbon, with almost a complete absence of NH 3-N transformation. In contrast, macrofauna-containing systems converted approximately 90% of the organic carbon, and substantial nitrification occurred. Fauna respiration explains reduced biomass yield in the trickling filter process, suggesting that approximately 10% of the carbon dioxide produced in the trickling filter process results from macrofauna
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
(a)
(b)
Figure 3.18 (a) “Pouch” snail typically found in trickling filters treating m unicipal wastewater, and (b) a modular plastic media sheet with biofilm and s nail-grazing pathways. respiration (Curds and Hawkes, 1975). Solbe et al. (1967) found that the absence of macrofauna adversely affected trickling filter sludge settleability. During a 1-hour settling period, 34% of the solids settled compared with 68% when the macrofauna were present. On the other hand, trickling filter macrofauna may have the following detrimental effects:
(1) (2) (3) (4)
(5) (6)
(7) (8)
Grazing of nitrifying biofilm (Figure 3.18); Plugging of process piping; Damaging pumps; Damaging belts on gravity-belt sludge thickening and belt-press dewatering equipment; Organic snail bodies remaining in the effluent stream, which may exert BOD5; Shells remaining in the effluent stream, which may increase effluent fecal counts by shielding bacteria from disinfection processes; May exert additional solids loading on secondary sedimentation tanks; and May accumulate in aeration basins, thereby reducing aeration capacity and/ or efficiency in combined TF/SG processes (Figure 3.19).
7.5.1 Operational Strategies and Facility Improvements for Macro Fauna Control Several strategies have been applied to manage macrofauna accumulation and/or development in trickling filters, including physical, chemical, or a combination of
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(a)
(b)
Figure 3.19 Snail shells covering fine-bubble diffusers in suspended-growth reactors that are located in (a) Pueblo, Colorado; and (b) Garland, Texas.
physical and chemical applications. The key is application of a condition that is either toxic to the animals or creates an environment not conducive to their accumulation. Lee and Welander (1994) demonstrated increased nitrification after predator control using substances toxic to eukaryotic organisms. The toxic substance must have no effect on, or only temporarily inhibit, beneficial microorganisms (Parker et al., 1997). Operators have conducted site maintenance, which aids in reducing macrofauna presence in trickling-filter-based WWTPs. For instance, some operators have observed that the presence of filter flies may be reduced simply by maintaining a short stand of grass on the WWTP site. More specific strategies include periodic high-intensity hydraulic application, trickling filter flooding, pH adjustment with lime or sodium hydroxide, high-concentration ammonia dosing, trickling filter humus screening or accelerated gravity separation, gravity separation in low-velocity channels with a dedicated pumping circuit, eliminating dissolved oxygen from the trickling filter feed, adding salt, draining and freezing the infested unit, raising the temperature quickly, adding molluscicide (e.g., copper sulfate), and chlorinating the trickling filter’s influent stream. Many of these strategies have proven ineffective in some trickling filters, and others may be detrimental to trickling filter performance. Biochemical reactions are influenced, to some degree, by temperature, pH, and alkalinity; adjusting these parameters may inhibit the biochemical reactions and lower transformation rates. Chemicals, such as chlorine, are toxic to all organisms in the trickling filter and may result in the destruction of
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
sensitive biomass (Parker et al., 1989). Control mechanisms reviewed here include the following: • Periodic high-intensity hydraulic flushing (Spülkraft), • Trickling filter flooding and chemical application, • Chemical treatment (focus on high concentration NH 3-N dosage and pH adjustment with sodium hydroxide), • Trickling filter effluent or underflow (humus) screening or accelerated gravity separation using equipment typically associated with grit removal, and • Gravity separation in low-velocity channels and removal with a dedicated pumping circuit.
7.5.2 Spülkraft Spülkraft is the instantaneous dosing intensity as a function of distributor speed (ATV, 1983). Slowing distributor rotation increases the instantaneous dosing intensity and increases flushing, which removes excess biomass and helps to control macrofauna, while increased distributor rotation increases the wetting efficiency and improves the treatment efficiency. When mechanical distributor speed control—with an electrically driven rotary distribution mechanism, for example—is used, the definition of Spülkraft may be inverted to calculate the rotational speed required to achieve flushing intensity. The optimum Spülkraft is to be defined and, to a degree, may be siteand application-specific. If applied, the flushing Spülkraft period typically is 5 to 10% of the 24-hour period and will operate at 6 to 15 times the routine operating Spülkraft, which emphasizes the need for a significant range in rotational speed control. Those WWTPs retrofitted with distributor speed modulation initially may observe (2 to 10 weeks) higher settled trickling filter effluent BOD5 and TSS, as these sloughed solids are not readily settleable. This is similar to sloughing cycle effects that are common to many trickling filter facilities (Albertson, 1995). Hawkes (1955, 1963) demonstrated that high hydraulic loadings and instantaneous dosing rates can control filter fly development. As illustrated in Figure 3.20, increased hydraulic loading improves trickling filter media-wetting efficiency, reducing trickling filter media dry spots and eliminating ideal spawning areas for filter flies. Gujer and Boller (1984) reported that filter fly larvae were reduced, to quantities that did not have an effect on the NTF process, by hydraulic application. The THL must be sufficient to guarantee complete media-wetting. Pilot-scale NTFs, with a medium-
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Interval of Distributor 1–5 minutes 400
Rotation
30–55 minutes
Film
300 Organic Matter 200 g/cu ft 100 0
Anisopus (Fly)
40
20 0
Psychoda (Fly)
Organisms 30 103/cu ft 20 10 0
Lumbricilus (Worm)
10 0
Jan Jan Jan Jan Jan Jan 1952 1953 1954 1955 1952 1953 1954 1955
Figure 3.20 Biofilm and filter fly control with low-frequency dosing (cu ft × 0.028 32 = m3) (Hawkes, 1963).
density cross-flow media and fixed distributors, required a THL of 3 m/h. Grady et al. (1999) later reported that adequate media-wetting may be achieved at THLs of 1.8 to 2 m/h, with rotary distributors. In contrast, Andersson et al. (1994) tested three flushing intensities (SK values of 5, 40, and 80 mm/pass) and reported that the variable flushing intensity had no apparent effect on filter flies and worms in a pilot-scale NTF.
7.5.3 Flooding Trickling filter flooding requires adequate duty units to isolate a trickling filter for a period of 3 to 6 hours. The trickling filters must be designed as water-retaining structures, which is not typical and represents a small fraction of existing trickling filters.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Variations include (1) saline flooding, and (2) flooding and backwashing with an alkaline solution. Parker et al. (1997) reported the use of flooding to control filter flies and an alkaline backwash process to control other macrofauna in two 32-m diameter, 7.3m-deep, medium-density XFM NTFs at the Littleton-Englewood WWTP, Colorado. Online pH probes and a sodium hydroxide metering system allow for flood water pH adjustment by set point. The alkaline flood water is pumped in through the bottom of the NTFs, discharges into an overflow trough, and then is directed to the head of the WWTP for treatment. Alkaline treatment is reported to have removed 76% of the larvae at pH 9 and 99% at pH 10 (Parker et al., 1997). Subsequent research trials designed in response to full-grown snail development showed that flooding and backwash (4 hours at pH 9) reduced the snail quantity by two-thirds and returned the NTFs to high nitrification efficiency (Parker, 1998).
7.5.4 Chemical Treatment Everett et al. (1995) summarized several chemical treatment alternatives, including pH adjustment and chlorination, sodium chloride, and molluscicides (e.g., copper sulfate, metaldehyde, niclosamide, and trifenmorph). Factors such as pH, turbidity, and molluscicide dose are key in determining the chemicals’ application rate. Rotating biological contactors (RBCs) in Lafayette, Louisiana, applied sodium chloride at a dosing concentration of 10 mg/L for a 24-hour period, to effectively control the snail population. Calcium hypochlorite at 60 to 70 mg/L for a 2- to 3-day period effectively controlled snails in RBCs at the Deer Creek WWTP, Oklahoma City, Oklahoma. Copper sulfate at low concentrations (0.45 kg of copper sulfate per 3.785 m3) may effectively control snail accumulation. Ammonia is toxic to snails (Arthur et al., 1987). Lacan et al. (2000) conducted a laboratory-scale study and plant-scale application of undissociated aqueous ammonia (NH3-N(aq)) solutions with elevated pH to control snail growth (P. gyrina) in NTFs. Undissociated aqueous NH3-N(aq)—not the ammonium ion—is the snail P. gyrina toxophore. The concentration producing 100% mortality is a function of exposure time and the bulk-liquid NH3-N(aq) concentration. The laboratory-scale study demonstrated that an ammonium chloride (NH4Cl) solution at pH 9.2 (NH3-N(ag) = 150 mg N/L) resulted in 100% snail mortality. A much higher concentration of ammonia in a trickling filter’s influent stream (1000 to 1500 mg N/L) is necessary in practice to maintain the required NH3-N(ag) = 150 mg N/L because of the immediately reduced concentration resulting from axial dispersion, biofilm diffusion (both external and internal), and biochemical reaction (according to trickling filter bioreactor
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hydrodynamics). Lacan et al. (2000) estimated that an influent ammonia concentration of 1080 mg N/L resulted in an average concentration throughout the NTF of 185 mg N/L. Such a high-concentration NH3-N(aq) stream may be readily available in municipal WWTPs as solids processing recycle streams. In some cases, however, it may be necessary to purchase NH3-N(aq). The first full-scale application of this snail-control method was reported by Gray et al. (2000) at the Truckee Meadows WWTP, Reno Sparks, Nevada, which uses highdensity (215 m2/m3) media. Ammonia-rich anaerobic digester centrate was directed to an NTF recirculation pumping station. Sodium hydroxide was added to the recirculation stream to raise the pH to 9.05 (range between 9.0 and 9.5), which increased the NH3-N(aq) content of the centrate solution. Applied once per month, during application, an NTF is isolated, and the solution is recirculated through the trickling filter for approximately 2 hours. The first 20 to 50 minutes is dedicated to reaching hydrodynamic steady-state, and the remainder is the minimum recommended exposure time for 100% mortality of both adult snails and their larvae. The treatment solution is returned to the head of the WWTP after dosing is completed, and the NTFs then are flushed with secondary effluent in the “recirculation mode” for 10 hours. Figure 3.21 illustrates these operating modes. Periodic grab samples or online monitoring during the flushing cycle is necessary to ensure that the effluent NH3-N concentration is less than or equal to the influent concentration before the NTFs were returned to service. Lucero et al. (2002) observed that a similar procedure on carbon-oxidizing trickling filters required less than 8 hours. Lacan et al. (2000), who reported that the use of ammonia for snail control also has been applied to carbon-oxidizing trickling filters, suggested that the abovedescribed procedure will not produce deleterious effects in the carbon-oxidizing trickling filters because of the lower sensitivity of the heterotrophic bacteria to pH and NH3-N and more expedient regeneration of the heterotrophic biofilm. Rather than a reduced flushing period, it may be possible to eliminate the 10-hour backwashing step, if the carbon-oxidizing trickling-filter-based WWTP is not subject to NH3-N permit limitations and has typical pH discharge standards of “6 to 9” (Lacan et al., 2000). Lucero et al. (2002) applied this snail-control operation to combined carbon-oxidizing and nitrification trickling filters in a TF/SC process at the Duck Creek Wastewater Treatment Center, Garland, Texas. The TF/SC process returned to full nitrification capacity after a period of declining performance, subsequent to snail treatment. Diminishing nitrification performance because of an inhibitory substance in the raw wastewater was reported later.
Nitrifying TF Normal Operating Mode
C
To Headworks C C Centrate Sodium C Hydroxide
OO OO Secondary Effluent TF Effluent
O
C O
C Recirculation Pump Station
Nitrifying TF Centrate Treatment/Recirculation Mode
C OO
C
OO Secondary Effluent TF Effluent
C
O C
To Headworks O Centrate Sodium O Hydroxide
O Recirculation Pump Station
Nitrifying TF Flushing Mode
O OO OO Secondary Effluent TF Effluent
O
C C
C
To Headworks C Centrate Sodium C Hydroxide
O Recirculation Pump Station
Legend O = Open Valve C = Closed Valve = Operating Flow Path
Figure 3.21 Nitrifying trickling filter operating modes for high-concentration undissociated aqueous ammonia dosing (Lacan et al., 2000).
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7.5.5 Physical Control Physical removal techniques described here include the following:
(1) Trickling filter effluent or underflow (humus) screening, (2) Gravity separation in low-velocity channels and removal with a dedicated pumping circuit, and (3) Accelerated gravity separation using equipment typically associated with grit removal.
The Central WWTP, Baton Rouge, Louisiana, uses trickling filter secondary clarifier underflow screening to control snail accumulation in, or damage to, solidshandling equipment. Through a series of statistical analyses, Lin and Sansalone (2001) found that snail infestation in another trickling-filter-based WWTP in Baton Rouge, Louisiana, did not affect BOD5 removal efficiency in carbon-oxidizing trickling filters. However, snail shells filled gravity-thickening tanks, plugged sludge process piping, and damaged belts in sludge thickening and dewatering equipment. A structure containing parallel static screens subsequently was erected, with underflow pumps directing the trickling filter humus through them. Snail shells are retained by the screens and fall by gravity into a collection bin for disposal. Waste sludge flows through the screens to sludge thickening. Although effective in protecting biosolids handling equipment, the system is reported to be a source of odor. The City of Lawton, Oklahoma (49 000 m3/d) and both the South San Luis Obispo, California, County Sanitation District Oceana Regional Plant (19 000 m3/d) and the City of San Luis Obispo Water Reclamation Facility, San Luis Obispo, California (9000 m3/d), pump trickling filter secondary clarifier underflow to a free vortex classifier for snail shell removal. Each of these systems is reported to prevent excessive snail shell buildup in digesters. The units remove approximately 0.69, 0.076, and 0.23 m3 snails/d, respectively (Neumayer, 2002). The Econchate Water Pollution Control Plant in Montgomery, Alabama, removes snail shells remaining in the secondary effluent, which were shielding pathogens from chlorine disinfection. The chlorine contact basin was modified to a two-pass channel, to serve as a low-velocity sedimentation basin for snail shells escaping secondary clarification. The snail shells deposited in the low-velocity channel are collected in a sump and pumped through a static screen, where they fall, by gravity, into a collection bin. The wastewater is returned to the chlorine contact basin for disinfection and final disposal.
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Tekippe et al. (2006) reported the use of baffles, grit pumps, and classifiers to remove snails from the Ryder Street WWTP, Vallejo, California. The facility treats wastewater with a TF/SG process consisting of two 32-m-diameter and 7.3-m-deep XFM trickling filters. The trickling filters were reported to produce snail shells in quantities that hindered aeration basin diffuser performance, requiring both the aeration basins and the secondary clarifiers to be taken out of service for cleaning and maintenance. The Spülkraft was ineffective to control snail accumulation in the trickling filters. The properties of the snail shells were found to be similar to grit— essentially light enough to be pumped by conventional centrifugal pumps. A small portion of the snails have air entrained in their shells, and others may be disconnected from their shell, so a fraction of the snails in a trickling filter effluent may float or be neutrally buoyant. The influent of the Ryder Street WWTP’s aeration basins subsequently was improved to provide a zone for the majority of the shells to settle and an automatic mechanism to remove the settled shells (Tekippe et al., 2006). Redwood baffles were inserted to the basin entrance to distribute inlet velocity and others downstream to contain the snail shells in the first 6 m of the rectangular aeration basins. The basin floor was sloped to prevent snail shell accumulation in the corners and direct the snails to a sump, and the aeration system in this region of the basins was modified to promote settling, with a rolling flow pattern similar to that in an aerated grit removal chamber. The shells then are conveyed, via grit pump, to a grit cyclone and classifier system adjacent to the aeration basin influent pumping station. The classifier discharges to a bin, and the shells are transported to final disposal. Tekippe et al. (2006) reported that, while the system removed approximately 49 000 m3/d during the first few days of operations, since startup, the quantity of shells stabilized to approximately 1.53 m3/d.
7.6 Trickling Filter Startup Understanding the formation sequence and composition of the biofilms that grow in trickling filters will aid designers in making well-informed decisions regarding process optimization, design, startup, and emergency shutdown. The phylogenetic composition of mature trickling filter biofilms is defined by their development and life-cycle processes. On a time series, trickling filter biofilm development follows an S-shaped curve, or sigmoidal function, and can be divided into (1) initial activities, (2) exponential accumulation, and (3) steady-state (Characklis and Marshall, 1990). Biofilms exist in cycles that include transport and adsorption of organic molecules to the substratum (surface conditioning); transport and attachment of microbial cells
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to the surface of the trickling filter media (formation); bacterial transformation processes (biofilm growth and extracellular polymeric substance [EPS] excretion); and partial biofilm detachment, as a result of fluid shear stress (Trulear and Characklis, 1982). The initial process of organic layer adsorption occurs within minutes of initial exposure of the clean surface to the wastewater stream. Baier (1973) conducted studies demonstrating that this initial adsorption is a prerequisite to biofilm formation because of the substratum-conditioning effect. Bacterium transport to the substratum is dependent on the bulk liquid flow regime. Some of the bacteria are reversibly attached, and, as a result, some of these cells are detached as a function of trickling filter hydrodynamics. Others remain and become irreversibly attached. The irreversibly attached cells use bulk liquid substrate and nutrients. This is a continual process, in which its rate is best defined by reactions kinetics post-biofilm formation. The attachment process is promoted by the presence of EPSs. The irreversibly attached bacteria then produce EPSs, which bioflocculate other bacterial cells and particulate matter (both organic and inorganic), which accelerates biofilm formation—hence the presence of an exponential biofilm formation phase. Thörn et al. (1996) performed pilot-scale investigations on NTFs treating municipal wastewater. A portion of the researchers’ results are illustrated in Figure 3.22. Similarly, Biesterfeld and Figueroa (2002) investigated autotrophic nitrifying biofilm development on glass slides in fullscale NTFs treating municipal wastewater. Thörn et al. (1996) observed virtually identical nitrifying biofilm formation/maturation periods for two media densities. The investigation of Biesterfeld and Figueroa (2002) improved on the other study by measuring these parameters at different trickling filter depths, rather than apparent rates measured for the entire bioreactor. Although the rates are different in each zone, the trends are comparable. Nitrification may be observed after 1 month. However, these investigations suggest that the maximum transformation rate occurs between 2.5 and 3 months. Interestingly, there is an increase in observed biomass accumulation before obtaining the maximum transformation rate for the system or zone. Then, the biomass diminishes slightly when the maximum transformation rate is approached. The decrease is attributed to biofilmdegrading processes, such as grazing, sloughing, and bacterial hydrolysis. Nitrifying biofilms develop at a much slower rate than carbon-oxidizing biofilms. The heterotrophic bacteria that constitute the carbon-oxidizing biofilm will demonstrate removal after only days, but there is little comparative evidence that allows for the definitive assignment of a trickling filter carbon-oxidizing biofilm formation period. Stenquist et al. (1974) observed that a pilot-scale, synthetic media, combined carbon-oxidizing
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
Nitrification rate, g N m–2 d–1
2 157 m2 m–3 226 m2 m–3
1.5 1 0.5 0 –0.5
0
20
40 60 80 100 120 Time after start-up, days
Dry weight
Bacterial count 8 1011 7 1011
35 30
6 1011 5 1011 4 1011
25 20 15 10
3 1011 2 1011
5
1 1011
0
0
20
40 60 80 100 120 Time after start-up, days
Bacterial count
Dry weight, g m–2
140
0 140
Figure 3.22 Nitrification rate and biomass accumulation in pilot-scale NTFs after startup (Thörn et al., 1996).
and nitrification trickling filter treating high-strength cannery wastes achieved fairly steady BOD5 removal rates only 1.5 weeks after startup. Little evidence exists to support the concept of accelerated biofilm formation promoted by seeding trickling filters with activated sludge. In fact, the competition for substrate between the suspended flocs in the bulk liquid and the forming biofilm may retard the biofilm’s developmental process.
7.7 Combined Trickling Filter and Suspended-Growth Processes Biological processes, including both a trickling filter and suspended-growth reactor, build on the known performance and operating characteristics of the parent processes. Numerous combinations of TF/SG processes are possible and are influenced by the
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type of parent process used, loading to the treatment units, and the point at which biological or recycle sludge is reintroduced to the main flow stream (Harrison et al., 1984). Most combined TF/SG processes use a biofilm reactor in series with an SGR and a secondary clarifier. The systems can be classified by their suspended-growth component. When the SGR is used primarily as a bioflocculating system, with limited oxidation of soluble organic matter, the unit is referred to as the TF/SC process. All other TF/SG processes use the SGR as a substantial oxidizing unit. If a relatively small trickling filter is used, then a larger SGR is required (Grady et al., 1999). The capacity to which each of the units is designed is largely an economic tradeoff. Trickling filters are known for their energy efficiency, low maintenance requirements, and ability to resist shock loads. Designers have found that combining the shock resistance of a trickling filter with suspended-growth processes known for producing high-quality effluent often can yield better results than using only one of the two parent processes. Commonly used TF/SG processes typically are classified by those that have low to moderate organic loadings to the trickling filter versus those with high organic loads (e.g., roughing trickling filter). The typical TF/SG process schematic in Figure 3.5 illustrates common methods for returning secondary sludge, with or without reaeration. Designers use separate terms to differentiate process modes. When an intermediate clarifier is used, the process typically is considered to be a staged system; that is, the processes are not directly coupled. Common modes for dual processes are described below. Table 3.5 lists generally accepted design criteria for various TF/SG processes.
7.7.1 Activated Biofilter The activated biofilter (ABF) process uses a moderately loaded, with respect to organics, trickling filter. Modular plastic or redwood media must be used in the ABF trickling filter component (rather than rock), because return sludge is incorporated to the primary effluent and recycled over the trickling filter media. Some designers have concluded that improved sludge settleability occurs with the ABF. One theory as to why the benefit of a lower sludge volume index (SVI) occurs is that the high foodto-microorganism ratio and dispersed plug-flow of the trickling filter allows heterotrophic bacteria to be more competitive than filamentous bacteria. Bacterial selection occurs as a result. Although it performs well at low organic loads, the ABF process typically has been unsuccessful at consistently achieving effluent water quality, within tolerances (imposed by regulatory agencies) that would merit the selection of a trickling-filter-based process (i.e., concentrations less than 30 mg BOD5 or TSS/L) at organic loads of 0.9 to 1.6 kg BOD5/m3∙d (60 to 100 lb BOD5/d/1000 cu ft). The TF/ SC process allows for high-quality effluent with trickling filter total organic loading
Table 3.5 Design criteria for selected combined TF/SG processes (lb/d/1000 cu ft × 0.016 02 = kg/m3∙d; gpd/sq ft × 0.001 698 4 = m/h).* Design criteria Parameter
Range
Common
Activated biofilter (plastic media) Solids production (mg VSS in waste/mg BOD5 removed)
0.7 to 0.9
0.7
ABF hydraulic load (m /m ·d) (gpm/sq ft)
47 to 293 (0.8 to 5.0)
82 (1.4)
Trickling filter influent total organic load (TOL) (lb/d/1000 cu ft)
10 to 75
30
Sedimentation basin overflow rate at average day flow (m /m ·d) (gpd/sq ft)
33 to 49 (800 to 1200)
41 (1000)
Underflow concentration (% total solids)
0.6 to 1.2
1.0
Solids production (mg VSS in waste/mg BOD5 removed)
0.7 to 0.9
0.7
Trickling filter hydraulic load (m /m ·d) (gpm/sq ft)
5.9 to 117 (0.1 to 2.0)
59 (1.0)
Trickling filter influent total organic load (TOL) (lb/d/1000 cu ft)
20 to 130
80
Solids contact basin side water depth (m) (ft)
3.7 to 6.1 (12 to 20)
4.9 (16)
Solids contact basin HRT at average day flow (minutes)
10 to 60
45
Solids contact basin HRT at peak flow (minutes)
10 to 60
20
Solids contact basin SRT (days)
1.0 to 2.0
1.5
3
2
3
2
Trickling filter/solids contact (modular synthetic media) 3
2
Solids contact basin mixed-liquor suspended solids concentration (mg/L) 1.5k to 3000
2200
Sedimentation basin overflow rate at peak flow (gpd/sq ft)
1.2k to 1800
800
0.6 to 1.2
0.8
Solids production (mg VSS in waste/mg BOD5 removed)
0.8 to 1.2
1.0
Trickling filter hydraulic load (m /m ·d) (gpm/sq ft)
47 to 293 (0.8 to 5.0)
1.0
Trickling filter influent total organic load (TOL) (lb/1000/d/cu ft)
75 to 300
150
Aeration basin side water depth (m) (ft)
3.7 to 6.4 (12 to 21)
18
Aeration basin HRT at average day flow (minutes)
30 to 240
120
Aeration basin HRT at peak flow (minutes)
10 to 40
30
Aeration basin SRT (days)
1.0 to 8.0
3.0
1.5k to 6000
2500
Sedimentation basin overflow rate at average day flow (m /m ·d) (gpd/sq ft)
20.4 to 40.74 (500 to 1000)
32.6 (800)
Underflow concentration (% total solids)
0.6 to 1.2
0.8
Underflow concentration (% total solids) Trickling, roughing, or biofilter/activated sludge (plastic media) 3
2
Aeration basin MLSS concentration (mg/L) 3
2
* Values are valid for circular flocculator clarifiers with rapid sludge withdraw.
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of 150 to 200% of these values. Therefore, the ABF typically is dismissed as an inferior biotechnology. Furthermore, in cold-weather climates without short-term aeration, the ABF process has proven incapable of maintaining good performance. To overcome these problems, the ABF system later was modified to include a relatively small aeration basin, which collectively has been called the biofilter/activated sludge (BF/AS) section (Harrison and Timpany, 1988).
7.7.2 Trickling Filter/Solids Contact A majority of organic matter in municipal wastewater is colloidal or particulate material (Boltz and La Motta, 2007; Levine et al., 1985, 1991). Like most biofilm reactors, trickling filters are poor bioflocculating systems (Boltz et al., 2006). Therefore, the TF/SC process operates under the premise that trickling filter effluent contains a high concentration of not readily settleable colloidal and particulate organics. This material may be removed from the trickling filter effluent stream by bioflocculation in a solids contact basin (La Motta et al., 2004). The entrapped particulate and colloidal organics are hydrolyzed or removed from the biological process with the wasted biomass. The TF/SC process includes the following:
(1) A rock- or modular plastic-media trickling filter; (2) An aerated solids contact basin and/or return sludge or underflow, reaeration tank; and (3) A circular flocculator-clarifier capable of rapid sludge withdrawal.
The U.S. EPA (in a report prepared by Matasci et al., 1986) described the following distinguishing characteristics of the TF/SC process: • The majority of soluble BOD5 removal occurs in the trickling filter; • Return sludge is mixed with trickling filter effluent rather than primary effluent; • A solids contact basin functions to promote the bioflocculation of dispersed solids remaining in the trickling filter effluent stream (i.e., solids capture); • The solids contact basin SRT is less than 2 days; • The solids contact basin aerobic hydraulic retention time (HRT) is 1 hour or less based on total flow (including recycle); and • The solids contact basin is not designed for nitrification, although it may occur as a result of nitrifying biofilms growing in the trickling filter, detaching, and
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
essentially bioaugmenting the suspended biomass inventory that has accumulated in the solids contact basin. A solids contact basin typically is 5 to 20% of the volume that would be required with equivalent treatment by activated sludge alone. There are three modes of operating the TF/SC process—mode I, mode II, and mode III. Mode I relies exclusively on the solids contact basin for colloidal and particulate organic matter bioflocculation and the oxidation of residual soluble organic matter. Mode II relies exclusively on a return sludge aeration tank. After aeration, the aerated return sludge is mixed with trickling filter effluent for colloidal and particulate organic matter bioflocculation. Mode III makes use of both the solids contact basin and a return sludge aeration tank. A typical TF/SC process operates as mode I, but, as of 2001, more than one-half of the TF/SC-based WWTPs operated as mode III (or have the operational flexibility to operate as modes I or III). It should be noted that mode II seldom is used and typically is not recommended, as it does not have a solids contact basin and only a sludge reaeration tank. The TF/SC mode I typically is used when there is a high trickling filter TOL, which results in the potential for significant soluble BOD5 breakthrough to the solids contact basin. The use of a sludge reaeration tank in TF/SC mode III allows WWTP operators to sustain an increased biomass inventory, which accommodates a higher SRT, dampens the effects of inaccurate sludge wasting, and reduces the potential for suspended biomass washout during wet-weather events (Parker and Bratby, 2001). These operational modes are illustrated in Figure 3.23. Rock-media trickling filters in the TF/SC process have been demonstrated to be capable of processing TOLs up to 0.4 kg BOD5/m3·d (25 lb BOD5/d/1000 cu ft). Matasci et al. (1988) demonstrated that a TF/SC process incorporating a rock-media trickling filter produced average effluent suspended solids and BOD5 concentrations of 19 and 15 mg/L, respectively, at a 0.9-kg BOD5/m3·d (55-lb BOD5/d/1000 cu ft) TOL. The TF/SC processes, which include a modular plastic-media trickling filter, also have been demonstrated to be capable of processing a higher TOL than reported for the trickling filter process. Parker and Bratby (2001) stated that a TOL in the range 0.2 to 2.1 kg BOD5/m3·d (20 to 130 lb BOD5/d/1000 cu ft) may be applied to cross-flow media trickling filters in the TF/SC process. The lower end of this range refers to combined carbon oxidation and nitrification trickling filters. Suspended solids remaining in the effluent stream are expected to increase with increasing TOL on the cross-flow media trickling filter. Significant odor potential exists for TOLs on
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Biofilm Reactors
Trickling filter Aerated solids contact tank
Secondary clarifier
Flocculator center well
Mixed liquor
Treated effluent
Primary effluent Return sludge
Waste sludge
TF/SC Mode I – Solids Contact Trickling filter Secondary clarifier Mixed liquor Primary effluent Waste sludge
Flocculator center well Treated effluent
Return sludge
TF/SC Mode II – Reaeration Trickling filter Aerated solids contact tank
Secondary clarifier Mixed liquor
Primary effluent Waste sludge
Flocculator center well Treated effluent
Return sludge
TF/SC Mode III – Solids Contact/Reaeration
Figure 3.23 Three modes of TF/SC process operation (after Parker and Merrill, 1984). c ross-flow media trickling filters in the range 1.3 to 1.8 kg BOD5/m3·d (80 to 110 lb BOD5/d/1000 cu ft). In context of the TF/SC process, the TOL that results in substantial odor generation from a rock-media trickling filter must be evaluated on a site-specific basis. The threshold for substantial odor generation may establish the upper limit TOL for a specific application. Covers, ventilation, and captured air treatment are recommended for trickling filters receiving a high TOL. Inadequate research
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
exists to demonstrate that the TF/SC process is capable of withstanding shock loads or that the process exhibits characteristic reliability and robustness when receiving a greater TOL. Solids contact basin HRT typically is in the range 10 to 60 minutes. Parker and Bratby (2001) reported that a minimum 2.0-mg/L bulk-liquid dissolved oxygen concentration has served as a design guide. La Motta et al. (2003) demonstrated that the suspended solids concentration remaining in the TF/SC process effluent stream at the Marrero WWTP (Marrero, Louisiana) was reduced when the air-induced velocity gradient was minimized. The researchers replaced existing coarse-bubble diffusers with fine-bubble diffusers and controlled the bulk-liquid dissolved oxygen concentration in the range 1.0 to 1.4 mg/L. Air generally is delivered through fine-bubble diffusers, because the air-bubble diameter directly affects TF/SC process performance. Low SVI values were observed in full-scale TF/SC processes that operated the solids contact basin at an aerobic SRT in the range 1.0 to 1.5 days (Parker and Bratby, 2001). Flocs may be subjected to mechanical shear and oxygen deprivation during transport from the solids contact basin to the flocculator-clarifier. Even mild air-induced turbulence in the solids contact basin can disrupt floc structure. Distressing the structural integrity of the flocs or allowing the biomass to revert to an anaerobic state may be detrimental to TF/SC process performance. Therefore, effective biomass management is an important characteristic of well-operating TF/SC processes. A long distance between the solids contact basin and clarifiers may require aerated channels or other process features to promote the maintenance of aerobic flocs. Additional flocculation opportunity is desired and typically is provided with a flocculating clarifier center well. The rapid return of biomass to the solids contact basin also is important to prevent the recurrence of anaerobic biomass. Flocculator-clarifier unit process design has been described by Parker et al. (1996). These units have special features, including flocculating center wells, weir placement, increased sidewater depth, and provision for sludge inventory management. Wahlberg et al. (1994) presented data suggesting that a 20-minute detention time (at average dry-weather flow, with an additional allowance for 50% return sludge flow) in the flocculation well promoted 92% of the obtainable degree of floc formation. Typically, the flocculation center well has a diameter in the range 32 to 35% of the clarifier diameter. The desired design and operational objective is to maintain a minimum sludge blanket thickness (e.g., 0.15-m thickness or less) during normal operating conditions. During a peak wetweather flow, the biomass inventory is transferred from the solids contact basin to the clarifier by the surging flow. The biomass accumulates in the clarifier as a rising
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Biofilm Reactors
sludge blanket. Over time, return sludge pumps convey the accumulated biomass to the solids contact basin. Flocculator clarifiers typically are deeper than conventional units. The sidewater depth generally is in the range 5.5 to 6.0 m. The extra volume exists to accommodate the biomass inventory transferred during wet-weather flow events (Parker et al., 1996). The process designer should note that a scraping-sludgewithdrawal mechanism promotes prolonged biomass retention in the clarifier sludge blanket. Thus, an increased potential exists for the return of anaerobic biomass and elevated suspended solids remaining in effluent stream.
7.7.3 Roughing Filter/Activated Sludge A common method of upgrading existing activated sludge plants is to install a roughing trickling filter before the activated sludge basin. This mode of operation is called the roughing filter/activated sludge (RF/AS) process. The roughing filter is a highly loaded trickling filter that uses 10 to 40% of the media volume required, if treatment has been accomplished through the use of the trickling filter process alone. The aeration basin HRT typically is 30 to 50% of that required with the activated sludge process alone. Although the terms trickling filter/activated sludge (TF/AS) and RF/AS are used interchangeably by designers, the term RF/AS will be used in this discussion. The TF/SC and RF/AS processes have the same process schematic. However, with RF/AS, a smaller trickling filter is used primarily to reduce the organic loading applied to the activated sludge process. Therefore, the suspended growth is responsible for a majority of BOD5 removal.
7.7.4 Biofilter/Activated Sludge The BF/AS process is similar to that of RF/AS, except that RAS is recycled over the trickling filter in a similar fashion as the ABF process. Incorporating RAS recycle over the trickling filter sometimes has reduced sludge bulking from filamentous bacteria, especially with food-processing wastes. Although it sometimes has improved sludge settleability, there is no evidence that sludge recycle improves the oxygen-transfer capability of the biofilter.
7.7.5 Trickling Filter/Activated Sludge The TF/AS process is designed at high organic loads similar to RF/AS or BF/AS. However, the TF/AS process may have an intermediate clarifier between the trickling filter and SGR. The intermediate clarifier removes solids produced in the trickling filter before partially treated wastewater enters the SGR. The reduced oxygen demand afforded by intermediate clarification typically is considered less significant
Trickling Filter and Combined Trickling Filter Suspended-Growth Process
than the savings in capital and operating costs gained by eliminating intermediate clarification. Therefore, most designers prefer to use the RF/AS or TF/SC processes rather than the TF/AS process.
8.0 REFERENCES Albertson, O. E. (1995) Excess Biofilm Control by Distributor-Speed Modulation. J. Environ. Eng. (Reston, Virginia), 121 (4), 330–336. Albertson, O. E. (1989a) Slow Down That Trickling Filter! Water Environ. Technol., 6 (1), 15–20. Albertson, O. E. (1989b) Slow Motion Trickling Filters Gain Momentum! Water Environ. Technol., 6 (8), 28–29. Albertson, O. E.; Davies, G. (1984) Analysis of Process Factors Controlling Performance Plastic Bio-media. Proceedings of the 57th Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Sept 3–Oct. 5; Water Environment Federation: Alexandria, Virginia. Albertson, O. E.; Eckenfelder, W. (1984) Analysis of Process Factors Affecting Plastic Media Trickling Filter Performance. Proceedings of the Second International Conference on Fixed Film Biological Processes, Washington, D.C. Albertson, O. E.; Okey, R. (1988) Design Procedure for Tertiary Nitrification. Surfpac, Inc.: West Chester, Pennsylvania. Andersson, B.; Aspregren, H.; Parker, D. S.; Lutz, M. (1994) High Rate Nitrifying Trickling Filters. Water Sci. Technol., 29 (10–11), 47–52. Arthur, J. W.; West, C. W.; Allen, K. N.; Hedtke, S. F. (1987) Seasonal Toxicity of Ammonia to Five Fish and Nine Invertebrate Species. Bull. Environ. Contam. Toxicol., 38, 324–331. Aryan, A. F.; Johnson, S. H. (1987) Discussion of: A Comparison of Trickling Filter Media. J. Water Pollut. Control Fed., 59, 915. Aspegren, H. (1992) Nitrifying Trickling Filters, A Pilot Study of Malmö, Sweden. Malmö Water and Sewage Works: Malmö, Sweden. ATV (Abwassertechnische Vereinigung) (1983) German ATV Regulations– A135. Grundsätze für die Bemessung von einstufigen Tropfkörpern und Scheibentauchkörpern mit Anschluwerter über 500 Einwohnergleichwerten. D-5205, St. Augustine, Germany (in German).
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Baier, R. E. (1973) Applied Chemistry at Protein Interfaces. Adv. Chem. Ser. Amer. Chem. Soc., 145, 1. Baxter and Woodman Environmental Engineers (1973) Nitrification in Wastewater Treatment: Report of the Pilot Study. Prepared for the Sanitary District of Bloom Township, Illinois. Benzie, W. J.; Larkin, H. O.; Moore, A. F. (1963) Effects of Climactic and Loading Factors on Trickling Filter Performance. J. Water Pollut. Control Fed., 35, 445–455. Biesterfeld, S.; Dane, M.; Dingeman, R.; Freeman, D.; Heppler, P.; Keilbach, K.; Oram, E.; Paterniti, D.; Wadas, D.; Lutz, M. (2005) Optimizing the TF/ SC Process for Nitrification. Proceedings of the 78th Annual Water Environment Federation Technical Exposition and Conference, Washington, D.C., Oct 9–Nov 2; Water Environment Federation: Alexandria, Virginia. Biesterfeld, S.; Figueroa, L. (2002) Nitrifying Biofilm Development with Time: Activity Versus Phylogenetic Composition. Water Environ. Res., 74, 470–479. Boller, M.; Gujer, W. (1986) Nitrification in Tertiary Trickling Filters Followed by Deep Filters. Water Res., 20, 1363–1373. Boller, M.; Gujer, W.; Nyhuis, G. (1990) Tertiary Rotating Biological Contactors for Nitrification. Water Sci. Technol., 22 (1–2), 89–100. Boltz, J. P.; Daigger, G. T. (2010) Uncertainty in Bulk-Liquid Hydrodynamics Creates Uncertainties in Biofilm Reactor Design. Water Sci. Technol., 61 (2), 307–316. Boltz, J. P.; La Motta, E. J. (2007) The Kinetics of Particulate Organic Matter Removal as a Response to Bioflocculation in Aerobic Biofilm Reactors. Water Environ. Res., 79, 725–735. Boltz, J. P.; La Motta, E. J.; Madrigal, J. A. (2006) The Role of Bioflocculation on Suspended Solids and Particulate COD Removal in the Trickling Filter Process. J. Environ. Eng. (Reston, Virginia), 132 (5), 506–513. Bratby, J. R.; Fox, B.; Parker, D. S.; Fisher, R.; Jacobs, T. (1999) Using Process Simulation Models to Rate Plant Capacity. Proceedings of the 72nd Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Oct 9–13; Water Environment Federation: Alexandria, Virginia.
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Bruce, A. M.; Merkens, J. C. (1973) Further Studies of Partial Treatment of Sewage by High-Rate Biological Filtration. Water Pollut. Control, 5, 499–527. Bruce, A. M.; Merkens, J. C. (1970) Recent Studies of High Rate Biological Filtration. Water Pollut. Control, 69, 113–148. Bryan, E. H. (1955) Molded Polystyrene Media for Trickling Filters. Proceedings of the 10th Purdue Industrial Waste Conference, West Lafayette, Indiana, May 9–11; Purdue University: West Lafayette, Indiana, 164–172. Bryan, E. H. (1962) Two-Stage Biological Treatment: Industrial Experience. Proceedings of the 11th South Municipal Industrial Waste Conference; University of North Carolina: Chapel Hill, North Carolina, 136. Bryan, E. H.; Moeller, D. H. (1960) Aerobic Biological Oxidation Using Dowpac. Proceedings of the Conference on Biological Waste Treatment; Manhattan College: Riverdale, New York. Characklis, W. G.; Marshall, K. C. (1990) Biofilms; Wiley and Sons: New York. Chartered Institution of Water and Environmental Management (1988) Unit Processes Biological—Manuals of British Practice in Water Pollution Control. Chartered Institution of Water and Environmental Management: London, United Kingdom. Crine, M.; Schlitz, M.; Vandevenne, L. (1990) Evaluation of the Performances of Random Plastic Media in Aerobic Trickling Filters. Water Sci. Technol., 22 (1/2), 227–238. Curds, C. R.; Hawkes, H. A. (1975) Ecological Aspects of Used-Water Treatment, Vol. I; Academic Press: London, United Kingdom. Daigger, G. T.; Heinemann, T. A.; Land, G.; Watson, R. S. (1994) Practical Experience with Combined Carbon Oxidation and Nitrification in Plastic Media Trickling Filters. Water Sci. Technol., 29 (10–11), 189–196. Daigger, G. T.; Norton, L. E.; Watson, R. S.; Crawford, D.; Sieger, R. B. (1993) Process and Kinetic Analysis of Nitrification in Coupled Trickling Filter Activated Sludge Processes. Water Environ. Res., 65, 750–758. Drury, D. D.; Carmona, J.; Delgadillo, A. (1986) Evaluation of High Density Cross Flow Media for Rehabilitating and Existing Trickling Filter. J. Water Pollut. Control Fed., 58, 364–366.
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Duddles, G. A.; Richardson, S. E.; Barth, E. F. (1974) Plastic Medium Trickling Filters for Biological Nitrogen Control. J. Water Pollut. Control Fed., 46, 937–946. Eckenfelder, W. W. (1961) Trickling Filter Design and Performance. ASCE J. Sanit. Eng. Div., 87, 33–45. Eckenfelder, W. W.; Barnhart, E. L. (1963) Performance of a High Rate Trickling Filter Using Selected Materials. J. Water Pollut. Control Fed., 35, 1535–1551. Everett, J. W., et al. (1995) Slowing Down a Snail’s Pace. Oper. Forum, 20–22. Galler, W. S.; Gotaas, H. G. (1964) Analysis of Biological Filter Variables. ASCE J. Sanit. Eng. Div., 90 (6), 59–79. Germain, J. E. (1966) Economical Treatment of Domestic Waste by Plastic Medium Trickling Filters. J. Water Pollut. Control Fed., 38, 192–203. Grady, L. E.; Daigger, G. T.; Lim, H. (1999) Biological Wastewater Treatment, 2nd ed.; Marcel Dekker: New York. Gray, R.; Ritland, G.; Chan, R.; Jenkins, D. (2000) Escargot…Going…Gone, A Nevada Facility Controls Snails with Centrate to Meet Stringent Total Nitrogen Limits. Water Environ. Technol., 12 (5), 80–83. Gujer, W.; Boller, M. (1986) Design of a Nitrifying Trickling Filter Based on Theoretical Concepts. Water Res., 20, 1353–1362. Gujer, W.; Boller, M. (1984) Operating Experience with Plastic Media Tertiary Trickling Filters for Nitrification. Water Sci. Technol., 16, 201–213. Gullicks, H. A.; Cleasby, J. L. (1990) Cold-Climate Nitrifying Biofilters: Design and Operation Considerations. J. Water Pollut. Control Fed., 62, 50–57. Gullicks, H. A.; Cleasby, J. L. (1986) Design of Trickling Filter Nitrification Tower. J. Water Pollut. Control Fed., 58, 60–67. Harrison, J. R. (2007) Personal communication. Harrison, J. R.; Daigger, G. T. (1987) A Comparison of Trickling Filter Media. J. Water Pollut. Control Fed., 59, 679–685. Harrison, J. R.; Daigger, G. T.; Filbert, J. W. (1984) A Survey of Combined Trickling Filter and Activated Sludge Processes. J. Water Pollut. Control Fed., 56, 1073–1079.
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Harrison, J. R.; Timpany, P. L. (1988) Design Considerations with the Trickling Filter Solids Contact Process. Proceedings of the Joint Canadian Society of Civil Engineers, ASCE National Conference on Environmental Engineering, Vancouver, British Columbia, July 13–15; American Society of Civil Engineers: Reston, Virginia, 753–762. Hawkes, H. A. (1963) The Ecology of Waste Water Treatment; Pergamon Press: Oxford, United Kingdom. Hawkes, H. A. (1955) Film Accumulation and Grazing Activity in the Sewage Filters at Birmingham. J. Proc. Inst. Sew. Purif., 88–110. Kincannon, D. F.; Stover, E. L. (1982) Design Methodology for Fixed-Film Reactors, RBCs and Trickling Filters. Civ. Eng. Pract. Design Eng., 2, 107–124. Kuenen, J. G.; Jørgensen, B. B.; Revsbech, N. P. (1986) Oxygen Microprofiles of Trickling Filter Biofilms. Water Res., 20 (12), 1589–1598. Lacan, I.; Gray, R.; Ritland, G.; Jenkins, D.; Resh, V.; Chan, R. (2000) The Use of Ammonia to Control Snails in Trickling Filters. Proceedings of the 73rd Annual Water Environment Federation Technical Exposition and Conference, Anaheim, California, Oct 14–18; Water Environment Federation: Alexandria, Virginia. La Motta, E. J.; Jimenez, J. A.; Josse, J. C.; Manrique, A. (2003) The Effect of AirEnduced Velocity Gradient and Dissolved Oxygen on Bioflocculation in the TF/SC Process. Adv. Environ. Res., 7 (2), 441–451. La Motta, E. J.; Jiminez, J. A.; Josse, J. C.; Manrique, A. (2004) The Role of Bioflocculation on COD Removal in the Solids Contact Chamber of the TF/SC Process. J. Environ. Eng. (Reston, Virginia), 130, 726–735. Lee, N. M.; Welander, T. (1994) Influence of Predation on Nitrification in Aerobic Biofilm Processes. Water Sci. Technol., 29 (4), 355–363. Lekhlif, B.; Toye, D.; Marchot, P.; Crine, M. (1994) Interactions Between the Biofilm Growth and the Hydrodynamics in an Aerobic Trickling Filter. Water Sci. Technol., 29, 423–430. Levine, A. D.; Tchobanoglous, G.; Asano, T. (1985) Characterization of the Size Distribution of Contaminants in Wastewater: Treatment and Reuse Implications. J. Water Pollut. Control Fed., 57, 805–816.
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Levine, A. D.; Tchobanoglous, G.; Asano, T. (1991) Size Distribution of Particulate Contaminants in Wastewater and Their Impact on Treatability. Water Res., 25 (8), 911–922. Lin, H.; Sansalone, J. (2001) Impact of Snail Infestation and Recirculation on Wastewater Treatment Plant Performance During Drought Conditions in the Gulf Coast. Proceedings of the 74th Annual Water Environment Federation Technical Exposition and Conference, Atlanta, Georgia, Oct 13–17; Water Environment Federation: Alexandria, Virginia. Logan, B. E. (1999) Environmental Transport Processes; Wiley & Sons: New York. Logan, B. E.; Hermanowicz, S. W.; Parker, D. S. (1987a) A Fundamental Model for Trickling Filter Process Design. J. Water Pollut. Control Fed., 59, 1029–1042. Logan, B. E.; Hermanowicz, S. W.; Parker, D. S. (1987b) Engineering Implications of a New Trickling Filter Model. J. Water Pollut. Control Fed., 59, 1017–1028. Logan, B. E.; Wagenseller, G. A. (2000) Molecular Size Distributions of Dissolved Organic Matter in Wastewater Transformed by Treatment in a Full-Scale Trickling Filter. Water Environ. Res., 72, 277–281. Lucero, B.; Foess, G.; Middleton, G.; Kucera, W.; Hoff, A. (2002) Snail Control in Trickling Filters. Presented at the Water Environment Association of Texas Annual Conference. Mabbott, J. W. (1982) Structural Engineering of Plastic Media for Wastewater Treatment by Fixed Film Reactors. Proceedings of the First International Conference on Fixed Film Processes, Kings Island, Ohio, April 20–23. Matasci, R. N.; Benedict, A. H.; Parker, D. S. (1986) Trickling Filter/Solids Contact Process: Full-Scale Studies, EPA-600/S2–86-046; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C. Matasci, R. N.; Clark, D. L.; Heidman, J. A.; Parker, D. S.; Petrik, B.; Richards, D. (1988) Trickling Filter/Solids Contact Performance with Rock Filters at High Organic Loadings. J. Water Pollut. Control Fed., 60, 68–76. Metcalf and Eddy, Inc. (2003) Wastewater Engineering: Treatment and Reuse, Tchobanoglous, G., Burton, F. L., Stensel, H. D. (Eds.); McGraw-Hill: New York. National Research Council (1946) Sewage Treatment at Military Installations. Sew. Works J., 18, 787–1028.
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Neumayer, A. (2002) Accelerated Gravity Removal of Snail Shells from Trickling Filter Plants. Paper presented at the Annual Conference of the Water Environment Association of Utah, St. George, Utah. Norris, D. P.; Parker, D. S.; Daniels, M. L.; Owens, E. L. (1982) High Quality Trickling Filter Treatment without Tertiary Treatment. J. Water Pollut. Control Fed., 54, 1087–1098. Okey, R. W.; Albertson, O. E. (1989a) Diffusion’s Role in Regulating Rate and Masking Temperature Effects in Fixed-Film Nitrification. Water Environ. Res., 61, 500–509. Okey, R. W.; Albertson, O. E. (1989b) Evidence of Oxygen Limiting Conditions During Tertiary Fixed-Film Nitrification. J. Water Pollut. Control Fed., 61, 510–519. Onda, K.; Takeuchi, H.; Okumoto, Y. (1968) Mass Transfer Coefficients Between Gas and Liquid Phase in Packed Columns. J. Chem. Eng. Jpn., 1, 56–62. Parker, D. S. (1998) Establishing Biofilm System Evaluation Protocols. WERF Workshop: Formulating a Research Program for Debottlenecking, Optimizing, and Rerating Existing Wastewater Treatment Plants. Proceedings of the 71st Annual Water Environment Federation Technical Exposition and Conference, Orlando, Florida, Oct 3–7; Water Environment Federation: Alexandria, Virginia. Parker, D. S. (1999) Trickling Filter Mythology. J. Environ. Eng. (Reston, Virginia), 125 (7), 618–625. Parker, D. S.; Bratby, J. R. (2001) Review of Two Decades of Experience with TF/ SC Process. J. Environ. Eng. (Reston, Virginia), 127 (5), 380–387. Parker, D. S.; Butler, R.; Finger, R.; Fisher, R.; Fox, W.; Kido, W.; Merill, S.; Newman, G.; Slapper, J.; Wahlberg, E. (1996) Design and Operations Experience with Flocculator-Clarifiers in Large Plants. Water Sci. Technol., 33 (12), 163–170. Parker, D. S.; Jacobs, T.; Bower, E.; Stowe, D. W.; Farmer, G. (1997) Maximizing Trickling Filter Nitrification Through Biofilm Control: Research Review and Full Scale Application. Water Sci. Technol., 36, 255–262. Parker, D. S.; Lutz, M.; Andersson, B.; Aspegren, H. (1995) Effect of Operating Variables on Nitrification Rates in Trickling Filters. Water Environ. Res., 67, 1111–1118.
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Parker, D. S.; Lutz, M.; Dahl, R.; Berkkopf, S. (1989) Enhancing Reaction Rates in Nitrifying Trickling Filters Through Biofilm Control. J. Water Pollut. Control Fed., 61, 618–631. Parker, D. S.; Lutz, M. P.; Pratt, A. M. (1990) New Trickling Filter Applications in the USA. Water Sci. Technol., 22 (1/2), 215–226. Parker, D. S.; Merrill, D. T. (1984) Effect of Plastic Media Configuration on Trickling Filter Performance. J. Water Pollut. Control Fed., 56, 955–961. Parker, D. S.; Richards, T. (1986) Nitrification in Trickling Filters. J. Water Pollut. Control Fed., 58, 896–902. Schroeder, E. D.; Tchobanoglous, G. (1976) Mass Transfer Limitations on Trickling Filter Design. J. Water Pollut. Control Fed., 48, 771–775. Schulze, K. L. (1960) Load and Efficiency of Trickling Filters. J. Water Pollut. Control Fed., 32, 245–253. Solbe, J. F.; de, L. G.; Williams, N. V.; Roberts, H. (1967) The Colonization of a Percolating Filter by Invertebrates, and Their Effect on the Settlement of Humus Solids. Water Pollut. Control, 66, 423–448. Stenquist, R. J.; Parker, D. S.; Dosh, T. J. (1974) Carbon Oxidation-Nitrification in Synthetic Media Trickling Filters. J. Water Pollut. Control Fed., 46, 2327–2339. Takács, I.; Newbeggin, M.; Stephenson, J.; Romano, L. (1996) Optimizing the TF/ SC Process for West Windsor Using a Comprehensive Modeling Technique. Proceedings of the 69th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 5–9; Water Environment Federation: Alexandria, Virginia. Tekippe, T. R.; Hoffman, R. J.; Matheson, R. J.; Pomeroy, B. (2006) A Simple Solution to Big Snail Problems—A Case Study at VSFCD’s Ryder Street Wastewater Treatment Plant. Proceedings of the 79th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 21–25; Water Environment Federation: Alexandria, Virginia. Thörn, M.; Mattsson, A.; Sorensson, F. (1996) Biofilm Development in a Nitrifying Trickling Filter. Water Sci. Technol., 34 (1/2), 83–89. Trulear, M.; Characklis, W. G. (1982) Dynamics of Biofilm Processes. J. Water Pollut. Control Fed., 54, 1288–1301.
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U.S. Environmental Protection Agency (1991) Assessment of Single-Stage Trickling Filter Nitrification, EPA-430/09–91-005; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C. U.S. Environmental Protection Agency (1993) Nitrogen Control Manual, EPA625/R-93–010; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C. U.S. Environmental Protection Agency (1975) Process Design Manual for Nitrogen Control; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C. Velz, C. J. (1948) A Basic Law for the Performance of Biological Filters. Sew. Works J., 20, 607–617. Wahlberg, E. J.; Keinath, T. M.; Parker, D. S. (1994) Influence of Activated Sludge Flocculation Time on Secondary Clarification. Water Environ. Res., 66, 779–786. Wall, D.; Frodsham, D.; Robinson, D. (2001) Design of Nitrifying Trickling Filters. Proceedings of the 74th Annual Water Environment Federation Technical Exposition and Conference, Atlanta, Georgia, Oct 13–17; Water Environment Federation: Alexandria, Virginia. Water Environment Federation; American Society of Civil Engineers; Environmental and Water Resources Institute (2009) Design of Municipal Wastewater Treatment Plants, 5th ed., WEF Manual of Practice No. 8, ASCE Manuals and Reports on Engineering Practice No. 76; McGraw-Hill: New York. Welty, J. R.; Wicks, C. E.; Wilson, R. E. (1976) Fundamentals of Momentum, Heat and Mass Transfer, 2nd ed.; Wiley & Sons: New York. Williams, N. V.; Taylor, H. M. (1968) The Effects of Psychoda alternata (Say) (Diptera) and Lumbricillus rivalis (Levinsen) (Enchytraeidae) on the Efficiency of Sewage Treatment in Percolating Filters. Water Res., 2, 139–150.
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Chapter 4
Rotating Biological Contactors 1.0 Introduction
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2.0 Process design considerations
3.6 Predicted Performance versus Full-Scale Data 191
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3.7 Temperature Correction192
2.1 Media Surface Area
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2.2 pH and Nutrient Balance
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4.0 Rotating biological contactor nitrification models 194
2.3 Oxygen Transfer
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2.4 Flow and Loading Variability
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2.5 Operating Temperature 183
5.0 Denitrification Application
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6.0 Physical design features 198 6.1 Physical Layout
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6.2 Tank Volume
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6.3 Hydraulics and Flow Control
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6.4 Media
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6.5 Drive Systems
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3.1 Monod Kinetic Model 184
6.6 Covers
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3.2 Second-Order Model
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6.7 Biomass Control
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3.3 Empirical Model
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7.0 Rotating biological contactor design examples
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2.6 Solids Production
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2.7 Toxic and Inhibitory Substances
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3.0 Rotating biological contactor design methods
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3.4 Manufacturers’ Design Curves 189 3.5 Comparison of Model Predictions
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7.1 Secondary Treatment Design Example
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(continued)
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7.2 Advanced Secondary Treatment Design Example
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8.0 Problems and corrective actions
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8.4 Loping of Air-Drive Systems
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8.5 High Clarifier Effluent Suspended Solids 207
8.1 Inadequate Treatment Capacity
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8.2 Excessive First-Stage Loadings
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9.0 Pilot-plant studies
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8.3 Excessive Biomass Growth
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10.0 References
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8.6 Corrosion of Media Supports
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1.0 Introduction The rotating biological contactor (RBC) process is a biological process that has been used widely for secondary and advanced secondary wastewater treatment. The process has been used for five-day biochemical oxygen demand (BOD5) removal, nitrification, or combined BOD5 removal and nitrification. As a secondary treatment process, it has been applied widely, where the average effluent limitations of 30 mg/L BOD5 and total suspended solids (TSS) are required, by permit. As an advanced secondary treatment process, it has been used often in conjunction with effluent filtration, where there are limits of ≤10 mg/L BOD5 and TSS and for effluent ammonia-nitrogen (NH3-N) levels as low as 1 mg/L. The RBC process has been used for aerobic pretreatment of industrial process wastewater, and it also has been used in an anoxic mode for denitrification. The RBC process is a form of fixed-film or attached-growth biological treatment. A circular disk of polystyrene or polyvinyl chloride is mounted on a horizontal shaft and partially submerged (typically 40%) in a tank holding the wastewater undergoing treatment. Plastic packing material, in a cylindrical basket or cage mounted on a horizontal shaft, also has been used. The media is slowly (1 to 1.6 rpm) rotated in the direction of the influent flow, to expose the biological film to food and nutrients and to provide necessary oxygen. Excess biological cell mass is “sloughed” from the media and removed from the wastewater forward-flow by clarification. Figure 4.1 provides a graphic representation of the process.
Rotating Biological Contactors
Cover Rotating biological contactor Interstage baffle
Oxygen
Influent Degradation products Food
Sludge
Effluent
Nutrients
Figure 4.1 General representation of RBC process.
A modification of the RBC process is the submerged biological contactor (SBC), in which the RBC is submerged up to 70 to 90%. Advantages may include lower shaft loadings, ability to retrofit the process to existing activated sludge aeration tanks, use of a dual air-header scouring system (which reduces biomass thickness), and ability to use larger-diameter media bundles. The design approach recommended is similar to that recommended for air-driven RBC systems. The SBCs also have been considered for application in denitrification processes where the shaft is mechanically driven and the introduction of air into the bulk liquid is minimized. In general, however, the SBC has seen limited application. Rotating biological contactor units commonly are manufactured in standard units, with a media bundle of approximately 3.5 m (12 ft) in diameter and 7.5 m (25 ft) in length. Standard-density units have a media surface area of approximately 9300 m2 (100 000 sq ft) per shaft, whereas high-density units typically have a media surface area of approximately 13 900 m2 (150 000 sq ft) per shaft. Standard-density units are used for BOD5 removal, where biological growth is relatively thicker, and more open media are necessary to maintain open passageways for wastewater flow throughout the media. High-density units most commonly are used in nitrification applications, where biofilm growths are relatively thin. Some designs will use lower surface area in the initial stages, followed by higher surface-area media. A more recent development
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is the use of plastic packing material instead of circular discs. The packing material, contained in a cylindrical cage mounted on the shaft, provides a very high specific surface area for biofilm growth. Rotating biological contactor units are covered to prevent algae growth, excessive heat loss in cold weather, and UV deterioration of polyethylene media. Prefabricated fiberglass, Quonset-hut types of covers commonly have been used, although conventional building structures also are used for covering the units. The RBC process typically is configured with several stages arranged in series, with one or more shafts composing each stage and with one or more parallel trains of shaft stages to provide the needed media surface area. Each train typically consists of several shafts installed in a common tank, with baffles installed between shafts to separate the shafts into distinct stages. The number of stages required depends on the degree of treatment desired, with one or two stages provided for roughing applications and six or more stages provided where advanced treatment with nitrification is necessary. Shafts typically are installed perpendicular to the direction of wastewater flow, with baffles for staging. For plants designed to treat relatively small flows, the shafts may be oriented in the direction of wastewater flow, with baffles provided to allow more than one stage per shaft. Figure 4.2 shows the general arrangements typically used for the RBC process. Shaft rotation most typically has been provided by mechanical-drive systems. Air-drive systems using diffused air and an array of cups, fixed to the periphery of the media, which capture some of the diffused air resulting in rotation from the buoyancy effect, also have been used. The mechanical drive system may be considered a constant-speed/variable-torque system, where shaft revolutions per minute are relatively constant, and torque requirements of the drive system vary with biomass growth and other conditions. The air-drive system may be considered a constanttorque/variable-speed system, where, at a given air rate, the torque applied to the shaft is relatively constant, and the shaft revolutions per minute vary with biomass growth and other conditions. The RBC process has a number of potential advantages, including simplicity and the need for minimal operator attention, low energy costs, low overall costs, and rapid recovery from shock loadings. However, there have been numerous examples of RBC process failure that have resulted from the following: • Structural failure of the shaft, media, or media support systems, • Less-than-anticipated treatment performance,
Rotating Biological Contactors
Typical parallel train configuration
Typical secondary treatment
Typical advanced treatment
End flow for small plants
Figure 4.2 Rotating biological contactor process general flow arrangements.
• Excessive development of nuisance organisms, • Development of excessive or uneven biomass growth, • Inadequate performance of air-drive systems for shaft rotation, and • Misapplication of pilot plant data. Primarily, these failures have been attributed to the following: • Misapplication of the process or use of inadequate design criteria; • Original designs for shafts and media support systems that were inadequate; • Inadequate upstream treatment; • Failure to consider the effect of sidestream loadings in the process design; • Lack of a thorough understanding of the long-term performance efficiency and characteristics of the process in full-scale use;
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• Development of design standards and criteria on the basis of limited data, much of which was developed using small-scale systems; and • General lack of a conservative approach in developing design relationships for the process. Many of the problems that were associated with earlier installations of the RBC process are now better understood, and the process has been used successfully at numerous locations. However, some regulatory agencies have established policies that restrict or prevent the use of the RBC process in their areas of authority. Therefore, the designer should investigate the applicability of the process in the area of the project before considering the RBC process. In addition to the anticipated benefits or advantages of the RBC process, it is essential that the designer consider the limitations of the process, long-term performance of the process that can be expected in full-scale applications, and other practical knowledge that is based on experience with the RBC process during the last 2 decades. As with any process, proper sizing and process design, a conservative and flexible design approach, and careful consideration to the specification of materials and components are necessary, if the process is to be applied successfully.
2.0 Process design considerations Soluble and particulate carbonaceous BOD5 (cBOD5) components of wastewater are reduced in the RBC process, by a combination of oxidation by the biofilm and synthesis into new cell mass. Soluble waste components, nutrients, and oxygen are transferred to the biofilm from the bulk liquid, and degradation products are transferred from the biofilm to the bulk liquid. Particulate waste components are enmeshed by the biofilm and hydrolyzed before degradation can occur. For the overall process to work successfully, the proper environmental and process conditions must be present. Principal RBC process design considerations include the following: • Media surface area, • pH and nutrient balance, • Oxygen transfer, • Flow and loading variability, • Operating temperature, • Excess biomass production, and • Toxic and inhibitory substances.
Rotating Biological Contactors
2.1 Media Surface Area Adequate media surface area is essential to the process, to provide sufficient biofilm to affect treatment and operate at specific substrate use rates (e.g., the mass of BOD5/ unit and time/unit media surface area). The substrate removal rate must be matched by the oxygen-transfer capability of the system. Sufficient media and proper operational controls are necessary to reduce the potential for nuisance organism development and limit excessive biofilm buildup, which could lead to structural damage to the shaft or media. Selecting the media surface area to be provided to meet design objectives is a balance between a conservative assessment of the capability of the process and economics. Because higher specific substrate use rates occur at higher bulk wastewater substrate concentrations, staging the process can provide the same degree of treatment that could be provided by a single-stage system having a larger surface area. A single-stage system would require that the bulk BOD5 concentration be limited to a relatively low value. This would reduce BOD5 removal rates, because, in the lower range of BOD5 values, specific substrate-use rates are related to concentration (first-order kinetics). Therefore, a more economical design can be provided by staging, where the earlier stages are loaded more heavily organically and have relatively high bulk fluid BOD5 concentrations and higher substrate-removal rates. However, care should be taken so that loadings to the first stages are not high enough to cause inadequate oxygen transfer, nuisance organism development, overweight shafts, and associated process problems. Early graphical methods (prepared by RBC suppliers) for determining media surface area requirements for the RBC process typically have yielded surface areas less than those required to meet design requirements for BOD5 removal based on actual long-term plant operating data. However, graphical methods evolved from a minimal database and pilot units, which later proved to be optimistic in the sizing of RBCs. Recently developed models, which are more reliable for defining surface-area requirements for BOD5 removal and nitrification, are presented and discussed later in this chapter. The reader also is referred to Chapter 3 for a discussion of modeling techniques for trickling filters, which also are applicable to the design of the RBC process. In addition, Chapter 11 provides a general understanding of modeling biofilm systems. A lack of adequate consideration of the effect of loadings from recycle streams, such as solids-processing unit operations or filter backwash, have been reported to be factors in unsatisfactory process performance at several plants. However, this is not
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a concern that is particular to RBC plants, and it is not caused by the RBC process. Therefore, a designer should develop a complete mass balance of the process, so that the effect of recycle sidestreams on influent flow and organic loadings to the RBC process are considered, and their effect on the determination of required media surface area is incorporated to the design.
2.2 pH and Nutrient Balance As with any biological treatment process, a proper pH and nutrient balance must be provided for effective treatment by the RBC process. The optimum pH for bacterial growth has been reported to be 6.5 to 7.5 (Metcalf and Eddy, 1979). Where nitrification is required, optimum performance has been observed at a slightly alkaline pH, with the optimum pH being 7.5 to 8.5 for suspended-growth systems (Sawyer et al., 1973). For RBC systems, nitrification rates have been reported to fall off rapidly, with the pH decreasing from 7.0 to 6.0 (Brenner et al., 1984). Well-acclimated RBC systems, where the pH is maintained in a narrow range, may be able to tolerate pH values somewhat below 7.0. However, with the minimal buffering at this pH, a stable pH is difficult to maintain. For systems required to nitrify, the alkalinity of wastewater is important because of the consumption of alkalinity. The autotrophic organisms responsible for nitrification consume inorganic carbon (alkalinity). The conversion of the base ammonia (NH3) to nitrate (NO3-) releases hydrogen ions (H+), thereby producing an acid (ammonium oxide [NHO3]). Approximately 7 mg of calcium carbonate (CaCO3) alkalinity are required for each milligram of ammonia-nitrogen oxidized. It has been recommended that a residual alkalinity of 50 to 100 mg/L be maintained in the effluent of a nitrifying plant (Envirex, Inc., 1989). This alkalinity level typically produces a pH of 7.0 or higher. Chemical addition may be necessary to wastewater where natural alkalinity is insufficient. As microorganisms providing treatment require nutrients for cell growth, an adequate supply of available nutrients is necessary for proper performance of the process. A minimum mass ratio of 100:5:1 (BOD: nitrogen:phosphorus) typically is used.
2.3 Oxygen Transfer The rate of oxygen transfer must be sufficient to maintain fully aerobic conditions in the RBC process. Organic loadings above the oxygen-transfer capability of the system result in reduced performance, odors, and development of nuisance organisms.
Rotating Biological Contactors
Maximum oxygen-transfer rates of 6.8 to 7.3 g oxygen/m 2·d (1.4 to 1.5 lb/d/1000 sq ft) have been reported (Scheible and Novak, 1980), based on tests conducted with full-scale, mechanical-drive systems. Studies of supplemental aeration applied to mechanical-drive RBCs indicated an oxygen-transfer efficiency of 2 to 2.5% for applied air (Chou, 1978). If it is assumed that the effective oxygen-transfer rate for an air-drive RBC is 2.5%, a 9300-m2 (100 000-sq ft) air-drive unit supplied with 7 m3/min (250 scfm) would have an oxygen-transfer rate of approximately 8.3 g oxygen/m2·d (1.7 lb/d/1000 sq ft). Maximum oxygen-transfer rates from 6.8 to 8.3 g oxygen/m2·d (1.4 to 1.7 lb/d/1000 sq ft) are consistent with the maximum reported ammonia-nitrogen removal rate of approximately l.5 g/m 2·d (0.3 lb/d/1000 sq ft), based on 4.6 kg oxygen/kg NH3-N oxidized. Historically, pilot testing of the RBC process was performed using small-scale units operating at the same media peripheral velocity (approximately 18 m/min [60 ft/min]) as full-scale units. Data on process performance collected on small-diameter units can be misleading because of the higher oxygen-transfer capability of the smaller units compared with full-scale units. Reh et al. (1977) estimated that a 0.5-mdiameter (1.5-ft-diameter) test unit provided an oxygen-transfer rate 1.6 times that of a 3.2-m-diam (10.5-ft-diam) unit that also was tested. Similar scale-up problems also were noted by other researchers (Brenner et al., 1984). For these reasons, data collected on small-scale systems should be used only for treatability considerations and not to establish sizing criteria. Manufacturers of RBC equipment historically have recommended maximum organic loadings to the first stage of RBC systems of approximately 20 g soluble BOD5 [sBOD5]/m2·d (4 lb/d/1000 sq ft) of media. This equates to an organic loading of 39 g BOD5/m2·d (8 lb/d/1000 sq ft) on a total BOD5-basis for a typical domestic wastewater with 50% soluble BOD5/BOD5 in the primary effluent. However, oxygen-transfer capabilities were reported to have been exceeded at first-stage loadings of more than 15 g sBOD5/m2·d (3 lb/d/1000 sq ft) (McCann and Sullivan, 1980). A survey of 23 plants indicated nuisance organism development because of oxygenlimiting conditions for first-stage loadings in excess of approximately 31.2 g BOD5/ m2·d (6.4 lb/d/1000 sq ft) of media (Weston, Inc., 1985). One RBC manufacturer currently recommends maximum first-stage loadings of approximately 12 to 15 g sBOD5/m2·d (2.5 to 3.0 lb/d/1000 sq ft) for unaerated RBC5 (Envirex, Inc., 1989). This rate equates to 24 to 29 g total BOD5/m2·d (5 to 6 lb/d/1000 sq ft), assuming 50% sBOD5/BOD5 (15 g sBOD5/m2·d [3 lb/d/1000 sq ft]). The manufacturer (Envirex, Inc., 1989) indicates that loadings below these levels provide a safety margin against
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Biofilm Reactors
unexpected loads and prevent sulfide formation. Another manufacturer (Lyco, Inc., 1992) recommends maximum first-stage loadings of 12 to 20 g sBOD5/m2·d (2.5 to 4.0 lb/d/1000 sq ft). The designer should consider limiting organic loading to approximately 29 g BOD5/m2·d (6 lb/d/1000 sq ft) (total BOD5-basis) to any RBC stage. This criterion will limit loadings to levels within the oxygen-transfer capability of the system and reduce the potential for growth of nuisance organisms.
2.4 Flow and Loading Variability Typically, higher effluent concentrations during the higher loading periods of the day are counterbalanced by lower effluent values during low-loading periods. However, a high degree of variability in influent flow or organic loading may disrupt the RBC process. Fixed-film processes, in general, are more susceptible to breakthrough of substrate, as a result of high peak loadings, than are suspended-growth processes. Loadings can increase to the point where breakthrough from the RBC process occurs. At this point, the substrate loading per unit surface area exceeds the substrate removal rate per unit surface area. This potentially could occur as a result of diurnal variations, in which case, it may be more cost-effective to provide flow equalization tanks for load equalization rather than provide additional RBC units. Another potential solution is to store peak day loads for feeding into the process at night, when the influent loads are minimized. It has been recommended (Envirex, Inc., 1989) that flow equalization is incorporated to the design if peak-to-average flow ratios exceed 2.5. The designer also should be aware of variability in organic loadings, including the effect of recycle loads from solids-processing unit operations. Particular attention should be paid to the sBOD concentration in the recycle from dewatering operations. Providing flow equalization may compensate for both flow and loading variability and eliminate the requirement for additional RBC tankage and media. This is especially true if stringent effluent limitations must be met. In the absence of equalization, a high variation in hydraulic or organic loading to the process would need to be addressed by providing additional RBC media and tankage. In addition to daily variations, the designer should consider peak daily loadings, which may result from industrial operations or other variable conditions. Because effluent limitations must be achieved during peak-loading periods and not just under average conditions, appropriate safety or peaking factors should be incorporated to the design. As with any process, the degree of safety factor incorporated to the design
Rotating Biological Contactors
also is influenced by the details of the discharge permit (instantaneous limit, monthly average, annual waste load allocation, etc.).
2.5 Operating Temperature Rates for the removal of BOD5 and nitrification are reduced by cold wastewater temperatures in RBC systems. It is important that the designer carefully review long-term data on wastewater temperatures and select a design temperature representative of critical cold-weather operating periods. For situations in which effluent limitations vary seasonally, each seasonal condition should be checked to determine which period is critical for design purposes. Biochemical oxygen demand removal and nitrification rates typically are reported by RBC equipment manufacturers as being relatively constant above 13°C (55°F) and declining significantly below this temperature (Envirex, Inc., 1989; Lyco, Inc., 1992; Walker Process, Inc., 1992). Relationships developed to adjust media-sizing criteria for reduced temperature operation are presented later in this chapter.
2.6 Solids Production Excess solids production must be estimated to size the solids-handling facilities of the plant. Solids production in the RBC process is a function of the synthesis of new cell mass, cell mass decay caused by endogenous respiration, inert solids present in the wastewater, and solids loss from the system caused by TSS in the clarified effluent. In general, solids production from the RBC process can be expected to be similar to other fixed-film processes. One manufacturer (Envirex, Inc., 1989) reports that waste solids production, based on clarifier underflow, ranges from approximately 0.4 to 0.6 kg/kg total BOD5 removed by the process. Lower values represent lightly loaded systems with higher endogenous oxidation of cell mass, and higher values represent more highly loaded systems. Net solids yields typically will range from 0.5 to 0.8 kg/ kg BOD5 removed from the process. The designer should use a conservative approach to solids production estimates, to account for periods of high biomass release from media and cold-temperature operation, where endogenous respiration will be lower. Typical secondary sludge concentrations for the RBC process are reported to be 2.5 to 3% (Envirex, Inc., 1989), with thickening in a continuous, low-volume solids withdrawal clarification system. Envirex, Inc. (1989) also reports concentrations of 4 to 5% achievable by gravity thickening of the RBC secondary sludge or co-thickening with primary sludge in primary clarification. As with other aerobic biological wastewater treatment processes, solids volatility can be expected to range from 80 to 95%.
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Biofilm Reactors
It generally is good practice to use clarifiers with conventional mechanical solidsremoval mechanisms with the RBC process. Spiral blade sludge collector mechanisms also should be considered. Siphon-type clarifiers typically are not required and will unnecessarily dilute settled solids. Thickening of RBC solids in the clarifier can lead to performance problems. Because the quantity of solids to be removed is relatively low and thicker, uptake pipes or openings are prone to plugging.
2.7 Toxic and Inhibitory Substances Toxic or inhibitory substances, if present in sufficient concentration in plant influent, will have a negative effect on wastewater treatment. In addition, excessive quantities of such substances may have a deleterious effect on solids management and effluent quality because of pass-through. The designer should review current technical literature for information on such substances if they are suspected of being present in concentrations that are deleterious to biological systems.
3.0 Rotating biological contactor design methods Numerous predictive models and equations have been proposed for estimating the removal of BOD5 in RBC systems. Several of these proposed relationships are described in this section. Actual plant performance data also are compared with the predictions of one model and with design curves published by two RBC manufacturers. Some models and manufacturers’ design approaches are based on sBOD 5, and others are based on total BOD5. For a consistent presentation, all models and design approaches are presented graphically on a total BOD5-basis, assuming that sBOD5 of primary effluent makes up 50% of the total BOD5 for both RBC influent and settled effluent.
3.1 Monod Kinetic Model As part of a study funded by the U.S. Environmental Protection Agency (Washington, D.C.) (U.S. EPA) (Weston, Inc., 1985), a design approach was presented based on Monod growth kinetics, as described by Clark et al. (1978). Using a mass-balance approach, assuming steady-state conditions, and making other simplifying assumptions, the following relationship was obtained:
R = (Fi/Ai)(So – Si) = PiSil(Ki + Si)
(4.1)
Rotating Biological Contactors
Where R = substrate removed per unit media surface area per unit time, FI = wastewater flowrate to stage i (gpd), Ai = area of stage i (sq ft), So = influent sBOD5 to stage i (mg/L), Si = effluent sBOD5 from stage i (mg/L), Pi = area capacity constant for stage i (gpd/sq ft-mg/L), and Ki = half-velocity constant for stage i (mg/L). Based on an analysis of interstage sBOD5 data from 11 RBC facilities, values for the maximum removal rate, P, and the half-velocity constant, K, were determined by plotting the data in the linearized form of eq 4.1, as follows: [1/R = (K/PS + 1/P)]. The majority of the facilities were air-driven, and none of the facilities were considered organically overloaded (first-stage total BOD5 loadings were below 31 g BOD5/m2·d [6.4 lb/d/1000 sq ft]). The P and K values were determined for stages 1 through 4, as follows:
(1) Stage 1—P1 = 1000 gpd/sq ft-mg/L and K1 = 161 mg/L (2) Stage 2—P2 = 667 gpd/sq ft-mg/L and K2 = 139 mg/L (3) Stage 3—Ps = 400 gpd/sq ft-mg/L and K3 = 82 mg/L (4) Stage 4—P4 = 100 gpd/sq ft-mg/L and K4 = 25 mg/L.
Using the above coefficients, eq 4.1 can be used to solve for the Si value exiting a given stage, as follows: (Stage 1)
Si =
(Stage 2)
(Stage 3)
{[ HL1 (So − K1 )] − P1 } + {[ HL2 (So − K 2 ) − P2 ]2 + [4( HL3 )2 (K 3 × So )]}0.5 2( HL1 )
(4.2)
where HL1, HL2, HL3 = hydraulic loading rate for stage 1, etc. (m3/m2∙d [gpd/sq ft]). In using the model for design purposes, information must be available, or assumptions must be made, for the ratio of soluble to total BOD5. The work presented in the U.S. EPA study (Weston, Inc., 1985) assumed a sBOD5 fraction of 0.5 for both the RBC influent and settled effluent, which is a typical value for domestic wastewater undergoing RBC treatment. The U.S. EPA study (Weston, Inc., 1985) reported that the model was typically conservative (predicted lower organic removal than observed), with the closest correlation for RBC influent values of approximately 100 mg/L BOD5 or less. The study was reported as being a “first step in collecting input data, which should be used to test the model over a variety of flows and loadings.”
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100 Roatating biological contactor influent total BOD5
90
BOD5 removal, %
186
80 75
70
100
125
150
50 60 50 40 0.5
1.0
1.5
2.0
2.5
3.0
3.5
Organic loading, lb BOD5/d/1000 sq ft
Figure 4.3 Clark RBC predictive model, Monod kinetics, BOD5 removal: (1) values are total BOD5; (2) 50% sBOD5 assumed; (3) >13°C (>55°F) temperature; (4) three stages, with 50% of media area in the first stage; and (5) standard-density media (lb/d/1000 sq ft × 4.882 = g/m2·d).
Figure 4.3 presents a graph of solutions of the Clark model for various RBC influent concentrations, as a function of organic loading rate for the assumptions noted. Soluble BOD5 is assumed to be 50% of total BOD5 in the development of this figure.
3.2 Second-Order Model A second-order kinetic model was developed (Opatken, 1980) based on an analysis of interstage data from two full-scale facilities. Predictions of the model were compared with data from nine other full-scale plants, and there was general agreement (Brenner et al., 1984). An equation was developed based on the Levenspiel equation (Levenspiel, 1972), having the following form:
Cn =
−1 + [1 + 4kt(Cn− 1 )]0.5 2kt
(4.3)
Rotating Biological Contactors
100
Rotating biological contactor influent BOD5
BOD5 removal, %
90 80
150 125
100
70
75 60 50
50
40 30 0.5
1.0
1.5
2.0
2.5
3.0
3.5
Organic loading, lb BOD5/d/1000 sq ft
Figure 4.4 Opatken RBC predictive model, second-order kinetics, BOD5 removal: (1) values are total BOD5; (2) >13°C (>55°F) temperature; (3) 50% sBOD5 assumed; (4) three stages, with 50% of media area in the first stage; and (5) standard-density media (lb/d/1000 sq ft × 4.882 = g/m2·d). Where Cn = sBOD5 effluent from stage n (mg/L); k = reaction constant (L/mg·h); and t = average hydraulic retention time (HRT) in stage n (hours). A k rate of 0.083 L/mg·h was determined to be appropriate for municipal wastewater, and a tank volume/media area relationship of 4.9 L/m2 (0.12 gal/sq ft) of media was used to determine HRTs. Figure 4.4 is a graph of solutions for Opatken’s model (1980) for various influent BOD5 concentrations and the assumptions indicated as a function of organic loading. Again, 50% sBOD5 was assumed for both RBC influent and effluent.
3.3 Empirical Model Because the RBC process is an attached-growth process similar to the aerobic biotower and trickling filter processes, mathematical relationships developed for those processes can be used to predict the performance of RBC systems. Benjes (1977) presented an empirical relationship to predict the performance of RBCs. The relationship
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was based on the Velz equation (1948) (see Chapter 3) and Schulze (1960), as applied to trickling filters and biotowers. The relationship predicts the fraction of influent BOD5 remaining in the effluent from the process, as a function of media volume and hydraulic loading, as follows: 0.5 Se = e − k (V /Q ) Si
(4.4)
Where Se = effluent total BOD5 (mg/L), Si = influent total BOD5 (mg/L), e = the natural number 2.7183, V = media volume (m3 [cu ft]), Q = average flowrate (m3/m2∙d [gpm]), and k = reaction constant, 0.30.
To establish the reaction constant, Benjes (1977) reviewed operating data from a number of plants. Figure 4.5 is a graph of solutions to the Benjes model for various influent concentrations, for the assumptions shown, as a function of organic loading rate. 95 Rotating biological contactor influent total BOD5
90 85 BOD5 removal, %
188
80 75
150
70
125 100
65
75
60 50
55 50 45 0.5
1.0
1.5
2.0
2.5
3.0
3.5
Organic loading, lb BOD5/d/1000 sq ft
Figure 4.5 Benjes RBC predictive model, BOD5 removal: (1) values are total BOD5; (2) >13°C (>55°F) temperature; and (3) standard-density media (lb/d/1000 sq ft × 4.882 = g/m2·d).
Rotating Biological Contactors
3.4 Manufacturers’ Design Curves Manufacturers of RBC equipment use design relationships that were developed from experimental and full-scale studies. Envirex, Inc. (1989) uses a design approach based on a family of curves to predict effluent BOD5 as a function of hydraulic loading rate and influent sBOD5. Slightly higher performance is predicted for air-drive units compared with mechanical-drive units. Design curves were generated from an experimentally developed relationship between effluent sBOD5 and organic loading (kgsBOD5/m2·d) of the media. Figure 4.6 shows these design relationships, assuming 50% sBOD5 and the other assumptions shown. Figure 4.6 also shows the relationship recommended by Envirex, Inc. (Waukesha, Wisconsin), which applies to the effluent from upstream biological treatment processes and relatively low RBC influent BOD5 (<100 mg/L). Figure 4.6 also shows the design relationships used by Lyco, Inc. (Marlboro, New Jersey), another RBC manufacturer. Lyco, Inc., uses total BOD5 loading as the basis for its design methodology. A third manufacturer, Walker Process, Inc., does not publish design curves and relies on a computerized design technique (Walker Process, Inc., 1992). In addition to the general application of their design procedures to establish media surface-area requirements, manufacturers use other factors, including maximum 50
Effluent total BOD5, mg/L
45 Mechanical drive
40 35
Air drive
30 25 20 15 10
2.0
2.5
3.0 4.0 3.5 4.5 Organic loading, lb BOD5/d/1000 sq ft
5.0
5.5
Figure 4.6 Manufacturers’ RBC design basis, BOD5 removal: (1) values are total BOD5; (2) >13°C (>55°F) temperature; and (3) 50% sBOD5 assumed (Envirex, Inc., 1989) (lb/d/1000 sq ft × 4.882 = g/m2·d).
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organic loadings to individual stages of the system, temperature corrections for cold-weather operation, wastewater characteristics, and other site-specific information. In some cases, site-specific factors dictate the application of safety factors in determining media area requirements (Envirex, Inc., 1992).
3.5 Comparison of Model Predictions With respect to predictive relationships presented in the preceding sections, a comparison can be made of the organic loading that would result in a given level of treatment for each of the relationships. Table 4.1 shows this comparison for typical domestic wastewater of 125 mg/L total BOD5 in the primary effluent, 50% of which is soluble, and for two different levels of treatment—76% total BOD5 removal to achieve secondary treatment and 90% total BOD5 removal. As indicated in Table 4.1, there is a wide range in predicted organic loading rates applicable to the two treatment levels, as predicted by the models or design approaches. The Benjes model predicts the largest surface-area requirement in each case and would require approximately 1.9 times the media area for the secondary treatment application and more than 3.0 times the media area for the advanced treatment application than that predicted by the model or design approach that yields the lowest media-area requirement. Therefore, the design engineer should use caution in the selection and application of design models in this situation. Table 4.1 Comparison of organic loadings for models. Loading (kg total BOD5/m2·d [lb/d/sq ft]) 76% removal
90% Removal
Clark model
12.7 (2.6)
6.8 (1.4)
Opatken model
15.1 (3.1)
3.9 (0.8)
Benjes model
11.2 (2.3)
3.4 (0.7)
Envirex, Inc., mechanical drive
19.5 (4.0)
10.7 (2.2)
Envirex, Inc., air drive
21.0 (4.3)
12.2 (2.5)
Envirex, Inc., <100 mg/L influent
13.2 (2.7)
7.3 (1.5)
Predictive model or design approach
Lyco, Inc.
17.6 (3.6)
8.8 (1.8)
Walker Process, Inc.
15.1 (3.1)
8.3 (1.7)
Rotating Biological Contactors
3.6 Predicted Performance versus Full-Scale Data Figure 4.7 compares performance predicted by the manufacturer’s design relationships (Envirex, Inc., 1989; Lyco, Inc., 1992) with data from 16 full-scale plants (Weston, Inc., 1985) and data from 11 additional full-scale plants (Doran, 1994). As indicated, the performance of the 27 full-scale plants was significantly lower than predicted by the manufacturer’s design method (Envirex, Inc., mechanical and air drive and Lyco, Inc.) at a given organic loading rate. The Envirex, Inc., design curve for biological effluent and low influent BOD5 concentrations (<100 mg/L total BOD5) compares more favorably with the plant data, but still would predict greater BOD5 removal than achieved at a given organic loading. Figure 4.8 compares performance predicted by the Benjes relationship with data from the same 27 plants as those in Figure 4.7. Figure 4.8 shows the Benjes solutions for 75, 125, and 175 mg/L RBC influent total BOD5 and the data from the 27 plants shown in subsets of 50 to 100, 100 to 150, and greater than 150 mg/L total BOD5 in full-scale plant influent. As indicated in Figure 4.8, predictions from the Benjes relationship compare well with data from the 27 plants over a wide range of conditions. As shown in Figure 4.8, some of the plant data indicate better performance, and some of the data indicate worse performance, compared with the performance 120 Plant data
Effluent total BOD5 mg/L
100
80
60 Mechanical drive
40
Air drive
20
0
0
1
2
3
4
5
6
7
8
Organic loading, lb BOD5/d/1000 sq ft
Figure 4.7 Equipment manufacturers’ predictions versus full-scale plant data, BOD5 removal (lb/d/1000 sq ft × 4.882 = g/m2·d).
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100 90 BOD5 removal, %
192
Rotating biological contactor influent BOD5
80 70
175 60 50 40
125
Data 50–100 mg/L 100–150 mg/L > 150 mg/L 0
1
2
75
3
4
5
6
7
8
Organic loading, lb BOD5/d1000 sq ft
Figure 4.8 Benjes RBC model predictions versus full-scale plant data, BOD5 removal: (1) values are total BOD5; (2) k = 0.30; (3) >13°C (>55°F) temperature; and (3) standard-density media (lb/d/1000 sq ft × 4.882 = g/m2·d). predicted by the Benjes relationship. The Benjes model, with k = 0.3, would appear to indicate expected average performance, but would not be suitable for a conservative design approach, without applying a safety factor. Figure 4.9 shows a family of curves representing k rates from 0.2 to 0.4 superimposed on the same plant data in Figures 4.7 and 4.8. A k of 0.27 would encompass approximately 75% of the data in Figure 4.9 and would be analogous to a safety factor of approximately 25% in media volume compared with the use of a k of 0.30. A k of 0.25 would encompass approximately 85% of the data of Figure 4.9. Larger safety factors may be dictated by more stringent maximum daily or weekly effluent limitations, high peak-to-average organic-loading ratios, and other site-specific factors.
3.7 Temperature Correction Cold wastewater temperatures can be expected to reduce the rates of microbial growth and substrate use. The van’t Hoff-Arrhenius relationship commonly has been applied to biological systems to adjust kinetic constants with temperature.
KT = θ ( Τ − 20) K 20
(4.5)
Rotating Biological Contactors
0.6 0.20
Benjes, Inc., k Plant data
Se /Si
0.5
0.25
0.4
0.30 0.35
0.3
0.40
0.2 0.1 0
0.5
1
1.5
2
2.5
3
3.5
4
Hydraulic loading rate, gpd/sq ft
Figure 4.9 Benjes RBC predictive model, effect of k rate on predicted performance (standard-density media assumed) (Envirex, Inc., 1989) (gpd/sq ft × 40.74 = L/m2·d).
Where KT = reaction rate at temperature, T (°C); K20 = reaction rate at 20°C; and θ = temperature coefficient. Theta values from 1.01 to 1.05 for BOD5 removal have been applied to attachedgrowth biological treatment systems. For the RBC process, it has been reported (Weston, Inc., 1985) that actual full-scale field data are sparse and have not provided adequate verification of the temperature coefficient. Benjes (1977) evaluated data from several full- and pilot-scale plants to determine the effect of temperature on the k rate for eq 4.4. From a typical k of 0.3 at temperatures of more than 13°C, the k was reduced to typical values of approximately 0.2 at 7°C and 0.15 at 5°C. All major manufacturers of RBC equipment (Envirex, Inc., 1989; Lyco, Inc., 1992; Walker Process, Inc., 1992) provide adjustment factors in their design approaches for adjusting media surface area requirements for design temperature, as shown in Table 4.2.
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Table 4.2 Media surface adjustment factors. Temperature correction factor for BOD5 removal ratios Envirex, Inc.
Lyco, Inc.
Walker Process, Inc.
18 (64)
1.00
1.00
1.00
13 (55)
1.00
1.00
1.00
10 (50)
0.87
0.83
0.87
7 (45)
0.76
0.71
0.73
6 (42)
0.67
0.66
0.65
4 (40)
0.65
0.62
—
(°C [°F])
With the correction factors in Table 4.2, the media surface area determined for 13°C (55°F) is divided by the applicable correction factor to determine the required media surface area at the design temperature. Nitrifying systems are predicted to be more sensitive to cold temperatures than systems designed for BOD5 removal. This is a similar situation to experiences with suspended-growth systems. Cold-temperature solids retention times are more than twice as high as those required at warmer temperatures for effective nitrification, depending on system conditions (U.S. EPA, 1977).
4.0 Rotating biological contactor nitrification models Rotating biological contactor systems have been designed to provide nitrification in combined BOD5 removal and nitrification systems, where sufficient media surface area is provided to accomplish both BOD 5 removal and ammonia oxidation. Less commonly, the RBC process has been used to nitrify effluent from other types of biological treatment processes. The total Kjeldahl nitrogen (TKN) concentration should be used when determining the ammonia concentration, because most of the organic nitrogen will be converted aerobically to ammonia. For typical domestic wastewater, RBC manufacturer design information (Envirex, Inc., 1989; Lyco, Inc., 1992; Walker Process, Inc., 1992) indicates that nitrification begins in the RBC process when the total BOD 5 concentration has been reduced to 30 mg/L or less (15 mg/L sBOD5 or less). Brenner et al. (1984), based
Rotating Biological Contactors
on an analysis of plant data, reported that maximum nitrification rates may not be achieved by the RBC process in some cases until sBOD5 values are reduced to approximately 5 mg/L. A relationship based on Monod kinetics was proposed by Pano et al. (1981), as follows:
Ci Z = kN K n + Ci
(4.6)
Where Z = NH3-N removal rate in RBC stage (g/m2·d [lb/d/1000 sq ft]), kn = maximum NH3-N removal rate (g/m2·d [lb/d/1000 sq ft]), Ci = NH3-N concentration in RBC stage (mg/L), and Kn = NH3-N removal half-saturation constant (mg/L). The values for kn and Kn of 0.478 and 0.4, respectively, were developed from pilotscale data at 15°C (59°F) (Brenner et al., 1984). The design basis used by one RBC manufacturer (Envirex, Inc., 1989) is based on a zero-order ammonia oxidation rate at 13°C (55°F), of 1.5 g/m2·d (0.3 lb/d/1000 sq ft) of media for substrate concentrations of more than 5 mg/L NH3-N, with a first-order ammonia removal rate for substrate concentrations of less than 5 mg/L. Envirex, Inc. uses a family of design curves, as shown in Figure 4.10, to determine the required media surface area for nitrification. Based on an evaluation of full-scale data at elevated temperatures, Brenner et al. (1984) report that the maximum zero-order rate of 1.5 g/m2·d (0.3 lb/d/1000 sq ft) does not increase with increasing temperatures above 13°C (55°F). This agrees with the recommendations of two manufacturers (Envirex, Inc., 1989; Lyco, Inc., 1992) that surface-area adjustments should not be made for temperatures of more than 13°C. It is possible that ammonia removal rates of more than 1.5 g/m 2·d (0.3 lb/d/1000 sq ft) are limited by the oxygen-transfer capabilities of the process. Also, high predation may be a factor in reducing nitrification rates at elevated temperatures. Based on an evaluation of limited data, Brenner et al. (1984) concluded that the average performance of full-scale data evaluated was close to the design basis of one manufacturer (Envirex, Inc.), although the plant data were scattered widely, with some data exhibiting ammonia removal rates substantially less than 1.5 g/m2·d. This suggests that a safety factor is needed for design.
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Wastewater temperature > 13 °C (>55 °F) Influent ammonia–nitrogen, mg/L 30
6
Effluent ammonia–nitrogen, mg/L
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22 20 19 18 17
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14 13
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3 2 1 0
0
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Figure 4.10 Manufacturers’ design basis, nitrification of domestic wastewater (Envirex, Inc., 1989) (gpd/sq ft × 40.74 = L/m2·d).
Lyco, Inc. (1992), recommends that the following specific nitrification rates be used for the indicated ranges of ammonia removal: removal range of NH3-N = 45 to 5, 5 to 3, and 3 to 2 mg/L for removal rates of 1.5, 1.2, and 0.75 g/m2·d, respectively. With the above approach, Lyco, Inc. further recommends that the mass of ammonia (kilograms per day) in each of the ranges be computed; that the required media area (square meters) be computed for each range, by dividing the appropriate mass by the removal rate; and that the required media area for each range be summed to determine the total media surface area requirement for nitrification. This approach is similar to one from Envirex, Inc. (1989), in that a maximum ammonia removal rate of 1.5 g/m2·d is assumed, and the removal rate is assumed to decrease for ammonia concentrations of less than 5 mg/L. In sizing RBC systems for nitrification, it is important to recognize that organic nitrogen present in the process also may be available for nitrification because of its hydrolysis to ammonia during the process. Randtke et al. (1978) determined that soluble organic nitrogen from full-scale activated sludge plants averaged 1.5 mg/L and
Rotating Biological Contactors
that approximately 1 mg/L of this was refractory organic material present in the raw wastewater. The amount of refractory organic nitrogen can vary significantly, if there are industrial contributions to the wastewater. On the basis of their studies, Barth and Bunch (1979) proposed that the following quantity of influent TKN is assumed to be available for nitrification and used as the basis of the RBC sizing: TKNa = TKNt – 1 mg/L – (0.055 BOD5r)
(4.7)
Where TKNa = TKN available for nitrification (mg/L), TKNf = total TKN in RBC influent (mg/L), and BOD5r = BOD5 removed in RBC process (mg/L). This relationship assumes that approximately 1 mg/L of the influent TKN is refractory and not available for nitrification and that 0.055 times the BOD5 removed by the process is used for the synthesis of new cell mass. The remaining TKN (the TKNa value) would be available for nitrification and be used as the basis of the RBC sizing. Brenner et al. (1984) also stressed the importance of considering the TKN content of solids processing recycle streams (digester supernatant, dewatering filtrate, heat processing liquors, etc.) and inclusion of these loadings in the process design. Chapter 3 presents additional information regarding various kinetic models used to design fixed-film systems.
5.0 Denitrification Application Because an SBC is almost entirely submerged, oxygen transfer can be minimized. Thus, an anoxic environment suitable for denitrification can be established. Applications of SBCs for denitrification are limited, but one interesting example is the Wallingford Wastewater Treatment Plant (WWTP) in Connecticut (Bradstreet et al., 2009). The existing 30 000 m 3/d (8.0-mgd) WWTP incorporated a predenitrification system to their existing RBC plant by recycling settled solids from the final clarifiers to a basin with 2.1 hours HRT at a capacity of 25 000 m 3/d (6.5 mgd). Pilot studies confirmed the ability of the process to establish a mixed liquor in the basins. After a period of operation during which the performance was optimized, an average annual effluent total nitrogen concentration of less than 8.0 mg/L was achieved. Efforts to improve performance focused on minimizing the dissolved oxygen recycle to the pre-denitrification zones, providing adequate wastewater carbon for denitrification, and establishing the optimum mixed-liquor suspended
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solids (MLSS) levels (1800 to 2000 mg/L in winter and 3800 to 4000 mg/L in summer). Higher MLSS levels tend to stress the final clarifiers, because, as is typical with RBC plants, they were not designed for such loadings. A post-denitrification system using SBCs with supplemental carbon addition was planned to decrease the effluent total nitrogen concentration to less than 5.0 mg/L. The design recommendation from the manufacturer was based on a loading of 14.4 g/m3 (0.9 lb NO3/1000 cu ft) safety factor of submerged biological contactor (SBC) media surface area.
6.0 Physical design features 6.1 Physical Layout In planning physical layout, a designer should consider equipment access for maintenance purposes and future plant-expansion requirements. Access to bearings, load cells, drives, air valves, and other equipment requiring routine inspection and maintenance should be provided for safe and convenient operation and maintenance by WWTP staff. The designer should consider the potential need for shaft removal for repair and replacement and access isles between trains for crane access. Tank drains should be provided to facilitate dewatering of tankage for inspection and maintenance. Portable blowers or other provisions for adequate ventilation should be provided for safe entry to any confined space. The future construction of parallel trains or other processes to provide future capacity or enhanced treatment should be considered during design. Adequate space should be reserved on-site for these facilities, as appropriate. “Knockouts” can be provided in channels, division boxes, and so on, to facilitate future plant construction.
6.2 Tank Volume The tank volume to contain the RBCs typically is sized at 4.9 L/m2 (0.12 gal/sq ft) of media for low-density units. Tanks for high-density units have been designed to be of the same size and configuration as for a low-density shaft having the same physical dimensions to simplify layout and for economic construction.
6.3 Hydraulics and Flow Control Flow through the RBC process is by gravity from the process influent to the clarifier effluent. In hydraulic design, the designer should consider providing effective
Rotating Biological Contactors
flow-splitting between parallel units, the ability to isolate process trains or to bypass tanks arranged in series to facilitate maintenance, and the ability to step-feed the first stages of the process to control first-stage organic loading. The ability to provide stepfeeding may be of particular importance to multi-shaft first stages, to allow for more uniform loading to the shafts making up the stage. Flow commonly is split to parallel trains by providing a low headloss influent channel perpendicular to the train flow path and by providing free discharge weirs at the influent end of each train. Aeration of the influent channel often is provided in such designs to reduce the potential for septicity and maintain solids in suspension in low-velocity influent channels. Feed to the influent channel may be best accomplished in a symmetrical fashion, to reduce potential problems with uneven distribution of loading to process trains. Isolation of process trains can be accomplished readily by using gates or stop logs at the influent and effluent ends of the trains. A free discharge on the effluent end and wide low-head influent weirs improve the capability of effective hydraulic isolation for maintenance. This is because a free discharge will prevent the effluent from backing up into an isolated channel (should isolation gates not be installed), and a wide low-head weir will minimize the increase in the water-surface elevation in the RBC trains remaining in service.
6.4 Media Rotating biological contactor units are provided in low-, medium-, and high-density configurations, with low-density units having a media surface area of approximately 118 m2/m3 (36 sq ft/cu ft), and high-density units having a media surface area of 180 m2/m3 (55 sq ft/cu ft). This compares with 100 m 2/m3 (30 sq ft/cu ft) (standard density) and 140 m2/m3 (42 sq ft/cu ft) (high density) for plastic biotower media. Low-density media are used in the first stages of RBC systems designed for BOD5 removal, to reduce potential media clogging problems and overweight problems caused by higher biomass accumulations. High-density media typically are used in the latter stages of systems designed to achieve relatively low BOD5 values and in portions of systems designed for nitrification. Medium-density media are recommended by one manufacturer (Lyco, Inc., 1992) as a transitional media to be installed following the initial stages of the process, where low-density media are used, and preceding high-density media used in low BOD5 loading stages and for nitrification.
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6.5 Drive Systems Both mechanical- and air-drive systems are available for RBC systems. Mechanicaldrive systems consist of an electric motor, speed reducer, and belt- or chain-drive components separately provided for each shaft. Multistage speed reducers also are used instead of belt or chain drives. Typically, 3.7-kW (5-hp) mechanical drives have been provided for full-size units (media approximately 3.5 m in diameter and 7.5 m in length). Air-drive units involve remote blower facilities delivering air to header systems located under each shaft. Air headers are equipped with coarse-bubble diffusers, and approximately 4.2 to 11.3 m3/min (150 to 400 scfm) of air are provided per standard size shaft, depending on media density, revolutions per minute, wastewater temperature, slime thickness and characteristics, and other conditions. The quantity of air required for air-drive systems is difficult to predict and can vary considerably from shaft to shaft and over time, depending on environmental and loading conditions. Mechanical-drive units have been designed for operation from 1.2 to 1.6 rpm, and air-drive units have been designed for operation from 1.0 to 1.4 rpm. Rotational speeds should be consistent, and an even biomass growth should be maintained, to avoid uneven shaft weight. Unbalanced growth may cause cyclical loadings in mechanical-drive systems and loping (uneven rotation) in air-drive systems. A loping condition often accelerates and, if not corrected, may lead to inadequate treatment and the inability to maintain rotation in the unit. For mechanical-drive systems, materials, and motor designs should be selected that will resist the corrosive effect of a humid operating environment. Rotationalspeed flexibility should be provided by alternative pulley or sprocket ratios or by providing variable-speed drives. Air-drive systems should provide ample reserve air supply to maintain rotational speeds, restart stalled shafts, and provide short-term, relatively high rotational speeds to control excessive or unbalanced biomass growth. Available data indicate that an air rate of 11.3 m3/min (400 scfm) or more per shaft may be needed to maintain 1.2 rpm in a heavily loaded situation (Brenner et al., 1984). Large-capacity air cups (150 mm) typically are provided in the first stages of the process. These air cups provide the capability to exert a greater torque on the shaft and reduce loping problems. Smaller-capacity air cups (100 mm) typically are provided on those stages receiving low BOD5 loadings and those designed for nitrification (Envirex, Inc., 1992). With air-drive systems, manual adjustments to the air rate to a given shaft may affect the air delivered to other units and their rotational speed. This problem has been overcome at some installations, by providing rotational-speed sensors and automatic
Rotating Biological Contactors
adjustment of valve settings to individual shafts to maintain selected revolutions per minute. Such systems also have been designed to provide sensing and alarms for uneven rotation or loping conditions.
6.6 Covers The RBC process requires covering for protection from atmospheric conditions, media deterioration from UV light, and algal growth. Rotating biological contactor systems have been installed in buildings or under prefabricated fiberglass-reinforced plastic (FRP) covers. The FRP covers should be equipped with doors, to allow inspection of the growth on the media. Buildings have been constructed of masonry, treated wood, and pre-engineered metal. The FRP covers typically are designed in sections to facilitate shipment to the job site and allow the subsequent removal of cover sections for RBC repairs or shaft removal. For designs using a building structure to house the RBC units, provisions also should be made for removing roof sections, shaft removal and replacement, or other major repairs. The designer also should consider adequate ventilation, humidity and condensation control, heat loss, and corrosion caused by the humid atmosphere within the building or cover. For this reason, some FRP cover designs provide bearing and drive locations outside the environment of the RBC.
6.7 Biomass Control Excessive biofilm thickness can result in process impairment because of excessive or uneven shaft weight, loping in air-drive systems, media clogging, excess energy consumption, and nuisance growths and odors caused by anaerobic conditions within the slime layer. Facilities should be provided in the design of the system to monitor shaft weight, as an indication of biomass buildup, and as a means to control shaft weight, if biomass develops beyond the range recommended by the manufacturer. Load cell devices commonly are provided to allow manual weighing of the shaft with a hand hydraulic pump and a pressure-sensing device. Electronic strain gauge load cells also are available. Excess biomass may be controlled by removing interstage baffles or step-feeding to reduce the stage organic loading; increasing rotational speed; temporarily removing a train from service and starving for a period of time; supplemental aeration; alternately reversing the rotational direction; or chemically stripping the media. The design should allow one or more means for biomass control.
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7.0 Rotating biological contactor design examples 7.1 Secondary Treatment Design Example Preliminary work has established the following design loadings (influent to the RBC system), including internal recycle loadings: • Average flow = 19 000 m3/d (5.0 mgd), influent BOD5 =125 mg/L; • Peak flow = 37 900 m3/d (10.0 mgd), minimum temperature = 10°C (50°F); and • Effluent requirements = 30 mg/L total BOD5. (1)
From eq 4.4, for k – 0.30 (warm weather), In(Se/Si) = −k(V/Q)0.5 In(30/125) = −1.43 = −0.30(V/Q)0.5 4.77 = (V/Q)0.5 22.7 =V/Q, cu ft/gpm V = 22.7(5.0)(106/1440) = 78 900 cu ft
(2) Number of shafts (2750 cu ft/shaft), (78 900 cu ft)/(2750 cu ft/shaft) = 28.7 shafts. Add approximately 25% for factor of safety, 28.7(1.25) = 35.9 shafts Adjust for 10°C (50°F) temperature, using 0.87 correction factor, 35.9(1/0.87) = 41.3 shafts (4 150 000 sq ft) (3) Size first stage for a maximum of 6 lb BOD/d/1000 sq ft, 5 mgd (8.34 lb/gal) (125 mg/L) = 5212 lb BOD5/d (5212 1b BOD5)/(6 lb BOD5/d/1000 sq ft/d) = 868 700 sq ft or nine lowdensity (100 000 sq ft each) media shafts.
(4) Choose 10 trains of four (would reduce safety factor to 21%). 5212 lb Bod 5/d = 5.2 lb Bod 5/d/1000 sq ft [(10 shafts) (100 000 sq ft/shaft)]
This is an acceptable loading. The baffle between the first and second shafts could be removed to reduce loading further, and three stages would remain.
Rotating Biological Contactors
7.2 Advanced Secondary Treatment Design Example A facility having the above design loads and flows, with RBC influent ammonia and TKN of 25 and 35 mg/L, respectively, should meet the following treatment standards: effluent BOD5 = 20 mg/L and effluent NH4-N = 5 mg/L.
(1) Shafts required for BOD5 removal down to 30 mg/L (onset of nitrification) equal 40 (from above example).
(2) Determine ammonia to use for design purposes. 35 mg/L influent TKN – 1 mg/L refractory effluent TKN – 6 mg/L TKN removed by synthesis [0.055] [125 – 20] = 28 mg/L
(3) Because effluent ammonia is 5 mg/L, base is on 0.3 lb NH 4-N/d/1000 sq ft. 5(8.34) (28 to 5)lb/d = 3 197 000 sq ft (0.3 lb/1000 sq ft/d)
Add approximately 25% safety factor. (3 197 000 sq ft) (1.25) = 4 000 000 sq ft 10°C (50°F) temperature correction factor = 0.78 (4 000 000 sq ft) / (0.78) = 5 150 000 sq ft Number of 150 000-sq ft shafts (5 150 000 sq ft)/(150 000 sq ft/shaft) = 34 shafts (4) Select arrangement. Select 11 trains of four low-density and three high-density shafts per train. This will provide a total of 869 000 m2 (9350 000 sq ft) of media, which compares with the computed total of 864 000 m2 (9 300 000 sq ft) (386 000 + 478 000 m2 [4 150 000 + 5 150 000 sq ft]). (5) Check media required to achieve BOD5 reduction to 20 mg/L. %BOD5 removed = [(125 to 20)/125]100% = 84% From Figure 4.5, choose an organic loading rate of 1.25 lb BOD 5/d/l000 sq ft. Adjust for temperature. (1.15 lb BOD5/d/1000 sq ft) (0.87) = 1.00 lb BOD5/d/1000 sq ft Include 25% safety factor.
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(1.00 lb BOD5/d/1000 sq ft) / (1.25) = 0.80 lb BOD5/d/1000 sq ft Compute loading based on above design. (5212 lb BOD5/d)·[(9350)1000 sq ft] = 0.57 lb BOD5/d/1000 sq ft Therefore, the selected design is acceptable from an overall BOD5-removal standpoint.
8.0 Problems and corrective actions Significant problems that have been experienced in applying the RBC process and corrective actions that have been used to reduce those problems are discussed in this section. Problems addressed in the following section include inadequate treatment capacity that must be resolved by loading reduction or increased plant capacity, high-effluent suspended solids and associated BOD5, and corrosion of media-supporting systems. Where treatment performance is inadequate, the cause or causes of the inadequate treatment should be understood fully before designing process improvements. For example, • High effluent soluble substrate concentrations may indicate inadequate RBC capacity, high loadings, or process upset; • Low-effluent soluble substrate concentrations, but high total effluent substrate concentrations, indicate inadequate solids removal; • Nutrient imbalances, excessive pH swings, and highly variable organic loadings indicate problems caused by industrial discharges; • Depressed pH and low alkalinity in systems designed for nitrification indicate the need for supplemental alkalinity; • High instantaneous flows may disrupt the process, by limiting retention times in the RBC system and interfering with clarification processes. Plant surveys should be performed to allow the RBC process performance to be characterized fully, and data should be collected from significant sources of plant loading, as appropriate, to provide the necessary information for implementing an effective remedy for poor process performance. Before proceeding with significant plant expansion, the designer should investigate the feasibility of influent loading controls, enhanced preliminary-primary treatment, controls or separate treatment of recycle streams, or improved suspended-solids control, to improve the overall performance of the RBC system.
Rotating Biological Contactors
8.1 Inadequate Treatment Capacity A number of RBC facilities have been constructed that have not achieved the degree of treatment anticipated during design. Typically, this has been traced to inadequate consideration of recycle loadings to the plant from solids processing, performance lower than that predicted by the RBC vendor’s design curves, or unanticipated industrial or other loadings. The resolution of treatment capacity problems has been accomplished by constructing additional treatment units or taking other appropriate steps to reduce system organic loading. The following steps potentially could apply: • Control or separate treatment of recycle sidestream loadings; • Implementation of industrial waste pretreatment programs; • Improvement of the degree of preliminary and primary treatment provided; • Installation of additional RBCs; • Construction of biotowers in series or parallel with the RBCs; and • Construction of activated sludge facilities, in series with or parallel to the RBCs, Additional RBC units would reduce the organic and hydraulic loading on the RBC facilities and increase the degree of treatment provided. Reduced loadings also would tend to mitigate problems with oxygen deficiency, nuisance microorganism growth, and excessive shaft weight, which are symptomatic of overloaded systems. The construction of other biological treatment systems, either in parallel or upstream of the RBCs, will accomplish the same results. With any of the above augmentations of the RBC process, the maximum capability of the RBCs would need to be assessed, and the most economical and practical process additions defined. Blending the effluents from a down-rated RBC system and another process, such as activated sludge, may apply, in some situations.
8.2 Excessive First-Stage Loadings The following actions should be considered to resolve excessive first-stage loadings: • Remove interstage baffles to increase the media area provided in the first RBC stage; • Increase RBC rotational speed to increase oxygen transfer and encourage media sloughing;
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• Add supplemental aeration to mechanical-drive systems; • Step-feed process influent to bypass the first stage, with some of the flow and loading, thereby spreading the organic loading between the first stages of the process; and • Add chemicals to the raw wastewater to enhance BOD5 removal during primary treatment.
8.3 Excessive Biomass Growth Excessive biomass growth typically results from excessive organic loadings. If the system capacity is inadequate to allow biomass control by the methods described, the following actions should be considered: • Chemical stripping of the media with caustic soda or other appropriate chemical treatment to remove slime buildup, • Alternate reversal of the RBC rotational direction, and • Discontinuation or reduction of flow to the overweight train to starve biomass and reduce mass by endogenous metabolism. Using the above procedures may have a negative short-term effect on the treatment efficiency. Also, the above procedures temporarily will increase loadings to the units remaining in service and may accelerate biomass growth on those units. In facilities subject to chronic overloading conditions, it is important to carefully monitor shaft weights, by making routine load cell measurements. This is particularly true during the winter (lower endogenous rate) and during periods of high plant loadings.
8.4 Loping of Air-Drive Systems Loping results from uneven biomass growth, which causes a non-uniform rotational speed. When established, a loping condition may be difficult to control in a given shaft. Decreasing the loading to the RBC stage or increasing the air rate and rotational speed may prove successful. Installing water feed to the air cups (water-assist) at a point past their uppermost rotation is a manner of increasing rotational torque. This practice will increase the effective loading on the shaft. For a shaft with repeated loping problems and with 100-mm (4-in.) air cups, the air cups may be replaced with larger 150-mm (6-in.) cups. Another solution, such as chemical cleaning or starving, may be necessary to resolve a serious loping problem.
Rotating Biological Contactors
8.5 High Clarifier Effluent Suspended Solids Fine colloidal solids typically will be flocculated and enmeshed during biological treatment, resulting in their removal in secondary clarification. Overflow of fine suspended solids may be a chronic problem in maintaining compliance with effluent limitations. Success in enhancing suspended solids removal in clarification has been reported in Lancaster, Wisconsin, where recycle from the final clarifiers is recirculated to the influent of air-drive RBCs to form a mixed liquor. The flocculent nature of the mixed liquor is reported to have enhanced the control of fine suspended solids, resulting in a substantial improvement in effluent quality (Doran, 1994). Certain RBC process modifications for enhanced effluent suspended-solids control, such as installation of RBCs in activated sludge aeration tanks and recirculation of secondary solids to a RBC contact zone, may be subject to patent restrictions. The use of polymer and other coagulant aids during clarification and filtration of the clarifier effluent also may be used for enhanced suspended-solids control. An excessive overflow rate or the design of the clarifier may be the cause of a high effluent suspended solids concentration.
8.6 Corrosion of Media Supports Corrosive failure of media support systems has been reported at some WWTPs. This situation may be the result of site-specific water chemistry, biological degradation products, original materials selection, incompatible welding materials, or other factors. The fact that some WWTPs report that this problem only occurs in particular portions or stages of the process suggests that biological activity may be an important factor. Welding repairs or selective material replacements may be required to maintain the integrity of the system. Cathodic protection also has been used, with apparent success. Manufacturers have continued to improve the structural design of the RBC, by modifying the materials of construction and the details of construction.
9.0 Pilot-plant studies Pilot-plant studies are recommended before application of the process to industrial or municipal wastewater containing a significant industrial waste component. Pilot-plant studies should be conducted using full-scale media, to minimize scale-up problems. Pilot-plant studies should be conducted for extended periods, to include seasonal effects, long-term effects of deposition of organic and inorganic materials on the RBC media, and other factors important to process application.
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Pilot-plant investigations should be designed to include all of the factors anticipated with full-scale installation, including daily changes in waste strength and flow. In addition, the effects of solids-processing sidestreams and other in-plant waste streams should be included. To provide the data needed to fully evaluate the process performance, the following data collection should be included, as appropriate: • Influent and effluent total BOD5 and influent and effluent sBOD5 (from each stage); • Dissolved oxygen, TKN, ammonia, and nitrate/nitrite stage-to-stage profiles; • Temperature trends and profiles; • Flow and flow variability; • Shaft weight trends; and • Air requirements for air-drive systems.
10.0 References Barth, E. F.; Bunch, R. L. (1979) Biodegradation and Treatability of Specific Pollutants, EPA-600/9-79-034; U.S. Environmental Protection Agency: Washington, D.C. Benjes, H. H., Jr. (1977) Small Community Wastewater Treatment Facilities— Biological Treatment Systems. Prepared for U.S. Environmental Protection Agency (Washington, D.C.) Technology Transfer National Seminar on Small Wastewater Treatment Systems, Culp/Wesner/Culp, El Dorado Hills, California; U.S. Environmental Protection Agency: Washington, D.C. Bradstreet, K., et al. (2009) Proceedings of the 82nd Annual Water Environment Federation Technical Exposition and Conference, Orlando, Florida, Oct 17–21; Water Environment Federation: Alexandria, Virginia, 1255–1276. Brenner et al. (1984) Design Information on Rotating Biological Contactors, EPA600/2-84-106; U.S. Environmental Protection Agency: Cincinnati, Ohio. Chou, C. C. (1978) Oxygen Transfer Capacity of Clean Media Pilot Reactors at South Shore. Autrotrol Corporation: Milwaukee, Wisconsin. Clark, J. H.; Moseng, E. M.; Asano, T. (1978) Performance of a Rotating Biological Contactor Under Varying Wastewater Flow. J. Water Pollut. Control Fed., 50, 896–911.
Rotating Biological Contactors
Doran, M. D., Strand Associates, Inc., Madison, Wisconsin (1994) Personal communication. Envirex, Inc. (1989) Specific RBC Process Design Criteria. Envirex, Inc.: Waukesha, Wisconsin. Envirex, Inc.: Waukesha, Wisconsin (1992) Personal communication. Levenspiel, O. (1972) Chemical Reaction Engineering; Wiley & Sons: New York. Lyco, Inc. (1992) Rotating Biological Surface (RBS) Wastewater Equipment: RBS Design Manual; Lyco, Inc.: Marlboro, New Jersey. McCann, K. J.; Sullivan, R. A. (1980) Aerated Rotating Biological Contactors: What are the Benefits? Proceedings of the 1st National Symposium on Rotating Biological Contactor Technology, Vol. I, EPA-600/9-80-046a; Champion, Pennsylvania. Metcalf and Eddy, Inc. (1979) Wastewater Engineering: Treatment, Disposal, and Reuse; McGraw-Hill: New York. Opatken, E. J. (1980) Rotating Biological Contactors—Second Order Kinetics. Proceedings of the 1st National Symposium on Rotating Biological Contactor Technology, Vol. I, EPA-600/9-80-046a; Champion, Pennsylvania. Pano, A., et al. (1981) The Kinetics of Rotating Biological Contactors Treating Domestic Wastewater, Water Quality Series UWRL/Q-8104; Utah State University, Logan, Utah. Randtke, S. J.; Parkin, G. F.; Keller, J. V.; Leckie, J. O.; McCarty, P. L. (1978) Soluble Organic Nitrogen Characteristics and Removal, EPA-600/2-78-030; U.S. Environmental Protection Agency: Cincinnati, Ohio. Reh, C. W.; et al. (1977) An Approach to Design of RBCs for Treatment of Municipal Wastewater. Paper presented at American Society of Civil Engineers National Environmental Engineering Conference, Nashville, Tennessee. Sawyer, C. N.; Wild, H. E., Jr.; McMahon, T. C. (1973) Nitrification and Denitrification Facilities, Wastewater Treatment, U.S. EPA Technology Transfer; U.S. Environmental Protection Agency: Cincinnati, Ohio. Scheible, O. K.; Novak, J. J. (1980) Upgrading Primary Tanks with Rotating Biological Contactors. Proceedings of the 1st National Symposium on Rotating Biological Contactor Technology, Vol. II, EPA-600/9-80-046b, Champion, Pennsylvania. Schulze, K. L. (1960) Load and Efficiency of Trickling Filters. J. Water Pollut. Control Fed., 32, 245–253.
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U.S. Environmental Protection Agency (1977) Process Control Manual for Aerobic Biological Wastewater Treatment Facilities, EPA-III/A-524-77; U.S. Environmental Protection Agency: Washington, D.C. Velz, C. J. (1948) A Basic Law for the Performance of Biological Filters. Sew. Works J., 20, 607–617. Walker Process, Inc., Aurora, Illinois (1992) Personal communication. Weston, Inc. (1985) Review of Current RBC (Rotating Biological Contactor) Performance and Design Procedures, EPA-600/2-85-033; U.S. Environmental Protection Agency: Cincinnati, Ohio.
Chapter 5
Moving-Bed Biofilm Reactors 1.0 INTRODUCTION
3.1.6.3 Combined Pre-/ Post-Denitrification Moving-Bed Biofilm Reactors 230 3.2 Mixers 231 3.3 Pretreatment 232
212
2.0 Moving-Bed Reactors 213 3.0 Design Considerations for Moving-Bed Reactors 217 3.1 Carrier Biofilms 218 3.1.1 Carbonaceous Matter Removal 220 3.1.2 High-Rate Designs 220 3.1.3 Normal-Rate Designs 221 3.1.4 Low-Rate Designs 221 3.1.5 Nitrification 224 3.1.6 Denitrification 229 3.1.6.1 PreDenitrification Moving-Bed Biofilm Reactors 229 3.1.6.2 PostDenitrification Moving-Bed Biofilm Reactors 230
4.0 Solids Separation
233
5.0 General considerations for moving-bed biofilm reactors 234 5.1 Approach Velocity 234 5.2 Foaming 234 5.3 Media Transfer and Inventory Management 234 6.0 CASE STUDIES 235 6.1 Moa Point Wastewater Treatment Plant, Wellington, New Zealand 235 6.2 Harrisburg Wastewater Treatment Plant, Harrisburg, Pennsylvania 238 (continued)
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6.3 Moorhead Wastewater Treatment Facility, Moorhead, Minnesota (Adapted from Zimmerman et al., 2004) 240 6.4 Williams Monaco Wastewater Treatment Plant, Henderson, Colorado 241
6.5 Klagsham Wastewater Treatment Plant, Malmö, Sweden (Adapted from Taljemark et al., 2004) 246 6.6 Gardemoen Wastewater Treatment Plant, Gardemoen, Norway 250 7.0 References
253
1.0 INTRODUCTION In the past 20 years, the moving-bed biofilm reactor (MBBR) has been established as a simple yet robust, flexible, and compact technology for wastewater treatment. The MBBR technology has demonstrated success with biochemical oxygen demand (BOD) removal, ammonia oxidation, and nitrogen removal applications, in a variety of different treatment configurations, designed to meet a wide range of effluent quality standards, including stringent nutrient limits. Moving-bed biofilm reactors use specially designed plastic carrier elements for biofilm attachment held in suspension throughout the reactor by turbulent energy imparted by aeration, liquid recirculation, or mechanical mixing energy. In most applications, the reactor is filled between one-third and two-thirds full with carriers. Perforated plates or sieves located on the effluent-end of the reactor allow treated water to pass through to the next treatment step, but retain the media inside the reactor. Perhaps the most impressive aspect of an MBBR is in its versatility, allowing creative solutions by design engineers. The key differentiator for moving-bed technology when compared with other biofilm systems is that it combines many of the advantages of activated sludge with the advantages offered with biofilm systems, while, at the same time, trying to minimize the drawbacks of each. • Like other submerged-bed biofilm processes, MBBRs help to promote a highly specialized active biofilm that is well-suited for the particular
Moving-Bed Biofilm Reactors
conditions in a reactor. This highly active specialized biomass results in high volumetric efficiencies and increased process stability, resulting in a more compact reactor. • Unlike most other submerged-bed biofilm processes, the MBBR is a continuous flow-through process, eliminating the need for backwashing of the media to maintain throughput and performance; thus, headloss and operational complexity of the treatment step is minimized. • Moving-bed reactors can offer much of the same flexibility and flow-sheet simplicity as activated sludge processes, allowing multiple reactors to be configured in a flow-through series arrangement to achieve multiple treatment objectives (i.e., BOD removal, nitrification, and pre- and post-denitrification). This occurs without the need for intermediate pumping. • Unlike suspended-growth processes, biological performance in the MBBR does not depend on the solids separation step, because most of the active biomass is retained continually in the reactor. The solids concentration leaving the reactor with the treated flow is at least an order of magnitude lower in concentration. As a result, MBBRs are compatible with a variety of different separation techniques—not just conventional clarifiers. • The versatility of MBBRs allows the technology to be considered in a variety of different potential reactor geometries. For upgrades at existing plants, this makes MBBRs well-suited for retrofit installation to existing tanks.
2.0 Moving-Bed Reactors Numerous concepts, including a variety of different carrier types (mediums) combined with different reactor types (configuration and driving force), have been developed and applied in a mobile-bed (carrier) reactor format for wastewater treatment (Lazarova and Manem, 1994). However, currently, much of the interest in mobile-bed reactors for municipal wastewater treatment is with systems that use specially designed plastic carriers to create the moving-bed reactor. Much of the original research and development of plastic-carrier moving-bed reactors was conducted by staff at the Norwegian University of Science and Technology (NTNU) in Trondheim, Norway. This original work was initiated in the mid-1980s, motivated by an international effort to reduce point-source discharges of nitrogen to the North Sea. In Norway, engineers and researchers recognized that, in many instances, the
213
214
Biofilm Reactors
most cost-effective upgrade option would need to rely on a compact biofilm process (Odegaard et al., 1991). From this, the plastic-carrier moving-bed reactor was developed. In the late 1980s, the Japanese corporation NKK (Tokyo) also worked on development of a cylindrical hollow-tube carrier, coined Biotube. The original Kaldnes Moving Bed process was developed from the work conducted at NTNU in the late 1980s. Following its development, the Moving Bed biofilm technology was patented and commercialized in 1989, through an agreement with Kaldnes (Tønsberg, Norway), creating the subsidiary Kaldnes Miljøteknolgi AS. In 1994, the British company Anglian Water Services, Ltd. (Huntingdon, United Kingdom) bought the subsidiary and introduced the technology to the global market. In 2002, the Swedish company Anox AB (Lund, Sweden) acquired Kaldnes from Anglian Water. Before the acquisition, Anox AB had focused principally on its own variant of a moving-bed carrier aimed principally toward the pulp and paper industry. AnoxKaldnes and a number of other manufacturers have developed a number of different carriers over the years using different geometries, different materials, and different manufacturing techniques (See Table 5.1). To date, most AnoxKaldnes MBBR installations worldwide use K-type carriers. The AnoxKaldnes approach for carrier retention sieves has evolved based on the firm’s practical experience. Early designs used vertically mounted stainless-steel flat-panel sieves. AnoxKaldnes now uses horizontally configured stainless-steel wedge-wire sieves in aerated MBBR applications, because the position and orientation of the sieve takes advantage of the media and process aeration grid for scouring. Vertical-mounted flat-panel sieves still are used in applications where an aeration grid is absent (e.g., anoxic reactors). In these instances, an air-sparge system is mounted along the bottom edge of the panels to periodically pulse the sieves and dislodge any accumulated matter. Figure 5.1 shows a typical horizontal and vertical sieve installation. For aerated applications, AnoxKaldnes uses its custom fabricated aeration grid assembly consisting of distribution piping and small-diameter diffusers with 4-mm holes drilled along the underside. The rugged yet simple design of the aeration grids allows them to stand up to the abusive conditions in the reactor both while the reactor is in service and when the reactor is taken out of service and drained with the weight of the media resting on the reactor floor. The coarse-bubble design, use of stainless materials, and rugged construction eliminates any need for routine maintenance or periodic replacement of diffusers elements. Figure 5.2 shows a typical aeration grid.
Table 5.1 Plastic biofilm carriers. Bulk specific surface area*
Nominal carrier dimensions Carrier (depth; diameter) photograph
Manufacturer
Name
Veolia, Inc.
AnoxKaldnes™ K1
500 m2/m3
7 mm; 9 mm
AnoxKaldnes™ K3
500 m2/m3
12 mm; 25 mm
AnoxKaldnes™ biofilm chip (M)
1200 m2/m3
2 mm; 48 mm
AnoxKaldnes™ biofilm chip (P)
900 m2/m3
2 mm; 48 mm
ActiveCell™ 450
450 m2/m3
15 mm; 22mm
ActiveCell™ 515
515 m2/m3
15 mm; 22 mm
ABC4™
600 m2/m3
14 mm; 14 mm
ABC5™
660 m2/m3
12 mm; 12 mm
Bioportz™
589 m2/m3
14 mm x 18 mm
Infilco Degremont, Inc.
Siemens Water Technologies Corp.
Entex Technologies, Inc.
As reported by manufacturer.
*
215
216
Biofilm Reactors
(a)
(b)
Figure 5.1 (Left) Vertical mounted flat-panel sieves with air sparge system, and (right) horizontal sieves located over aeration grid.
(a)
(b)
Figure 5.2 (a) Aeration grid and distribution piping, and (b) stainless diffuser with 4-mm aeration-holes along underside. In recent years, several additional plastic-carrier moving-bed systems have emerged on the United States market. United States-based suppliers offer a product developed in Canada in the 1990s and a product developed in Israel in the early 2000s. The historical development of these two products is provided below. Additional manufacturers have come on to the market in the last ten years as the MBBR process has gained popularity and acceptance (see Table 5.1). In 1994, a study supported by Environment Canada (New Brunswick, Canada) began research on wastewater treatment technologies suited to provide cost-effective upgrades of municipal wastewater facilities to nitrify municipal wastewater in the typical cold wintertime climate in Canada. The study resulted in a biofilm-type
Moving-Bed Biofilm Reactors
upgrade to conventional activated sludge processes (i.e., integrated fixed-film activated sludge [IFAS]), which was deemed to offer a balance between cost and performance. The carrier that was ultimately developed was first used in full-scale as part of an IFAS nitrification study at the Waterdown, Ontario Sewage Treatment Plant (Canada) in 1997 to 1998. In 2003, Aqwise–Wise Water Technologies (Herzliya, Israel), with its Attached Growth Airlift Reactor (AGAR) process, was brought to the U.S. market. Founded in 2000, Aqwise developed the original AGAR process and commercialized the process with several small installations in Israel. A focus of the AGAR process has been to develop and use a plastic carrier that maximizes the protected area, while, at the same time, promoting efficient mass transfer from the bulk liquid to the biofilm. The first few installations of the AGAR process attempted to use a specially designed baffle wall combined with a hydraulic barrier to retain the carriers in the reactor, thus avoiding the need for sieves or screens to retain the media. However, this since has been dropped in favor of screen assemblies.
3.0 Design Considerations for Moving-Bed Reactors Process design with MBBRs is based on the concept that treatment is achieved with several reactors in series and that each individual reactor is designated for a specific treatment function within the greater treatment scheme. This is appropriate because each reactor promotes the development of a specialized biofilm oriented toward a treatment goal based on the conditions set within the reactor (e.g., available electron acceptor and available electron donor). This compartmentalized approach results in a rather simple and straightforward design, whereby one or more complete-mix reactors are oriented in series, each with a specified treatment purpose. Compare this with the design of a suspended-growth system, whereby competing reactions are always occurring and an overall system solids residence time (SRT) must be maintained to provide the mixture of bacterial cultures (based on growth rates and character of the raw wastewater) necessary to achieve the desired treatment goal within the hydraulic residence provided in each zone (aerated and unaerated) of the tank. The simplicity of MBBRs has allowed for a strong empirical understanding of biofilms with MBBRs based on observations made by researchers, engineers, and plant operating staff. Much of this section is dedicated to drawing on notable examples of observations that help demonstrate the key principals and considerations that are
217
218
Biofilm Reactors
important for proper design and operation of MBBRs. Chapter 11 discusses process models and plant simulators available for modeling MBBR systems. These models can be used for process design, process optimization, and process simulation.
3.1 Carrier Biofilms Critical to the success of any biofilm process is to maintain a high proportion of active biomass in the reactor. When the biomass concentration on MBBR carriers is presented in terms of an equivalent suspended solids concentration, values typically are approximately 1000 to 5000 mg/L suspended solids. Yet, when performance is assessed on a volumetric basis, results show that removal rates can be much higher than those compared with suspended-growth systems (Rusten et al., 1995). This added volumetric efficiency with MBBRs can be attributed to the following:
(1) High overall biomass activity resulting from effective control of biofilm thickness on the carrier resulting from the shear imparted on the carriers by the mixing energy (e.g., aeration); (2) Ability to retain highly specialized biomass specific to the conditions within each reactor, independent of an overall system SRT (see Figure 5.3); and (3) Acceptable diffusion rates resulting from the turbulent conditions in the reactor.
r LC
Figure 5.3 Representative carriers removed from each of four moving-bed reactors in series showing the variation in biofilm color (specialization and biofilm concentration (active biofilm) dependent on the operating condition and treatment function of each reactor
Moving-Bed Biofilm Reactors
Moving-bed reactors can be arranged in a number of different continuous-flow treatment schemes for BOD removal, nitrification, and denitrification. Table 5.2 provides a summary of the general potential treatment schemes with MBBRs. Selection of the most effective treatment scheme depends on the following:
(1) Site-specific conditions, including site layout and plant hydraulic profile; (2) Existing treatment scheme and potential for retrofit of existing facilities/ tanks; and (3) Targeted effluent quality objectives.
The net effective biofilm area is a key design parameter with moving-bed reactors (Odegaard et al., 2000), and loading and reaction rates can be expressed as a function of the surface area offered by the carriers. Hence, carrier surface area is convenient and commonly used for expressing performance and loadings of MBBRs often presented as surface area removal rate (SARR) and surface area loading rate (SALR), respectively. The substrate removal rate in MBBRs is zero-order-dependent when bulk substrate concentrations are high (i.e., S>>K) and first-order-dependent when the bulk substrate concentration is low (S<
L r = rmax K + L
(5.1)
Table 5.2 General summary of treatment schemes offered by MBBRs. Treatment objective
Process description
Carbonaceous matter removal
Stand-alone MBBRs Roughing MBBR preceding suspended growth process
Nitrification
Stand-alone MBBRs MBBR following secondary treatment IFAS (see Chapter 6)
Nitrogen removal (denitrification)
Stand-alone MBBRs with pre-denitrification Stand-alone MBBRs with post-denitrification Stand-alone MBBRs with pre- and post-denitrification Post-denitrification MBBR of nitrified effluent
219
220
Biofilm Reactors
Where r = removal rate (g/m2·d), rmax = maximum removal rate (g/m2·d), L = loading rate (g/m2·d), and K = half-saturation constant.
3.1.1 Carbonaceous Matter Removal The design SALR for a moving-bed reactor designated for carbonaceous matter removal will depend on the overriding treatment objective and the method selected for solids separation. Table 5.3 shows a range of typical design values for BOD loading based on application and treatment objective. The reactor should be designed using a low loading rate when nitrification is the objective of the downstream process. Higher loading rates can be considered when only removal of carbonaceous matter is required. Experience has shown that, for carbonaceous matter removal, a bulk dissolved oxygen level of 2 to 3 mg/L is sufficient and that higher bulk dissolved oxygen levels do not help to improve the SARR.
3.1.2 High-Rate Designs The moving-bed reactor can be considered when a compact high-rate system is needed to meet basic secondary treatment standards. In high-rate applications, the MBBR is operated under high SALR conditions, with the main purpose of removing the soluble and easily degraded BOD from the influent stream. The settling character of the biofilm slough diminishes under high loading conditions (Odegaard et al., 2000). As a result, the high-rate MBBR is combined with chemical coagulation and flocculation of the treated effluent, flotation, or with a solids contact process step to remove the solids; however, overall, this results in a compact system capable of meeting basic secondary treatment standards in a short hydraulic retention time (HRT) (Melin et al., 2004). Figure 5.4 shows results from a high-rate treatment study. Figure 5.4a shows that chemical oxygen demand (COD) removal was efficient and essentially linear across a wide range of high loading conditions. However, as Figure 5.4b shows, Table 5.3 Typical BOD design loading criteria. Application (objective)
BOD SALR (g/m2·d)
High rate (75 to 80% BOD removal)
>20
Normal rate (80 to 90% BOD removal)
5 to 15
Low rate (preceding nitrification)
5
Moving-Bed Biofilm Reactors
150
(b) 100 %
100%
125
SS-removal in settling tank
Obtainable removal rate [g COD/m2*d]
(a)
100 75 50 25 0
0
50
100
150
200
Total COD loading rate [g COD/m2*d]
v=0.05 m/h v=0.35 m/h
80 %
v=0.65 m/h
60 % 40 % 20 % 0%
0
20 40 Bioreactor loading [g COD/m2*d]
60
Figure 5.4 (a) COD removal efficiency at high loading conditions, and (b) poor settling character of biofilm slough under high loading rates. settling performance was poor, even at low surface overflow rates—hence the need for an enhanced solids-capture strategy. Figure 5.5 shows soluble BOD removal as a function of influent total BOD loading across a high-rate MBBR, which is part of the MBBR/solids contact process at the Mao Point Wastewater Treatment Plant (WWTP), New Zealand. The figure shows that typical BOD removal performance at high loadings was in the range 70 to 75%. Biological flocculation and polishing with the solids contact treatment step helps to meet basic secondary treatment standards (see Case Study on Moa Point Wastewater Treatment Plant, Wellington, New Zealand).
3.1.3 Normal-Rate Designs The moving-bed reactor offers a technology option when conventional basic secondary treatment approaches are being considered. Typically, two reactors in series are considered to achieve this level of treatment. Table 5.4 summarizes 7-day BOD (BOD7) removal performance at four normal-rate MBBR WWTPs designed for BOD removal followed by chemical removal of phosphorus. In these cases, the MBBRs were designed with an organic loading rate of 7 to 10 g BOD7/m2·d at 10°C and are combined with follow-up chemical addition and flocculation for phosphorus removal and enhanced solids separation.
3.1.4 Low-Rate Designs A low loading-rate design should be considered for reactors designated for organic matter removal upstream of a nitrification reactor. This helps to ensure that a high
221
Biofilm Reactors
Moa Point WWTP, New Zealand Surface Area Loading Rate vs. Removal Rate in MBBR from 1/1/99 - 4/5/00 Based on Influent Total BOD & MBBR Effluent Soluble BOD
100.00 90.00 80.00 Removal Rate (g/m2/day)
222
70.00 60.00 50.00 40.00 30.00 20.00 10.00 0.00 0.00
10.00
20.00
30.00
40.00
50.00
60.00
70.00
80.00
90.00
100.00
Surface Area Loading Rate (g/m2/day)
Figure 5.5 Soluble BOD removal rate as a function of total BOD loading in roughing MBBR application.
Table 5.4 Average results from normal-rate MBBRs with chemical phosphorus removal (adapted from Odegaard et al., 2004). BOD7 WWTP Steinsholt Tretten
b
Svarstad Frya
b
a b
b
a
COD
Total phosphorus
In (mg/L)
Out (mg/L)
In (mg/L)
Out (mg/L)
In (mg/L)
Out (mg/L)
398
10
833
46
7.1
0.3
361
4
–
–
7.3
0.1
–
–
403
44
5.1
0.25
181
5
–
–
8.6
0.21
1996–1997. Data from 2000–2002.
nitrification rate can be achieved in the downstream nitrifying moving-bed reactor, which, in turn, results in the most economical design. In instances where the BOD load to the nitrification MBBR is not sufficiently reduced, the nitrification rate will be reduced significantly, and the reactor may be ineffective.
Moving-Bed Biofilm Reactors
Figure 5.6a depicts the effect of increasing BOD load on carrier nitrification rates. This would be the case in a situation where a BOD load that was too high was selected for the prior organic matter removal stage of treatment, resulting in carryover to the nitrification stage of treatment. As an example, the figure shows that a nitrification rate of 0.8 g/m2·d could be achieved at a BOD load of 2 g/m2·d and a bulk dissolved oxygen of 6 mg/L, but would be reduced by roughly 50% if the BOD load increased to 3 g/m2·d. The operator could adjust by using a higher bulk dissolved oxygen level to help compensate for the impeded condition, or the fill fraction could be increased to reduce the surface area loading rate. However, it is important to note that neither of these options results in an economical and effective design approach. Rather, the preferred design approach would be to size the MBBRs designated for BOD removal by using a conservative low loading rate, so that the maximum efficiency can be attained in the downstream nitrifying MBBR. Figure 5.6b shows nitrification rates observed during a study involving three aerobic MBBRs in series. During the study, two separate bench-scale nitrification rate tests were conducted 6 weeks apart on carriers taken from each of the three MBBRs. For each test, nearly identical conditions (i.e., dissolved oxygen, temperature, pH, and initial ammonia concentration) were maintained in each of the three bench-scale vessels. The rate test results show that the first reactor, which was loaded at the highest soluble COD (SCOD) loading rate (5.6 g SCOD/m2·d), provided little, in terms of
(b) 14 15 °C
2.4
2 /d
2.1
OD
1.8
d=
1.5
nic
1.2
a rg
O
0.9
loa
0.
B 0g
/m
5
0 1.
0
2.
0
3.
0.6 0.3
0 4. 0 5.
0
6.
0.0 0
6 8 2 4 Oxygen concentration, mg O2/L
10
NOx-N Concentration (mg-N/L)
Nitrification rate, lg NH4-N/m2/d
(a) 2.7
R-1 SCOD = 5.6 g/m2/d R-2 SCOD = 2.8 g/m2/d R-3 SCOD = 2.2 g/m2/d
12 10
R-3 RMAX(20) = 0.78 g/m2/d
8 R-2 RMAX(20) = 0.75 g/m2/d
6
R-1 RMAX(20) = 0.29 g/m2/d
4 2 0
0
20
40
60
80
100
120
140
Time (min)
Figure 5.6 (a) Effect of BOD loading and dissolved oxygen on nitrification rates at 15°C (adapted from Hem et al. [1994]), and (b) difference in nitrification rate between multiple MBBRs in series.
223
224
Biofilm Reactors
nitrification, but was successful in its function of reducing the SCOD load before the second reactor. This is exhibited by the following:
(1) High nitrification rates observed for the second reactor, which were close to the rates observed in the third reactor, and (2) Small difference between the soluble COD loading condition in the second reactor compared with the soluble COD in the third reactor.
For low loading rate designs where it is important to select a conservative SALR, the following temperature correction factor can be used to adjust the SALR based on wastewater temperature (Sen et al., 2000): LT = L10 1.06(T − 10 ) Where LT = loading rate at temperature, T, and L10 = 4.5 g/m2·d at 10°C.
(5.2)
3.1.5 Nitrification A number of conditions within a nitrification-stage MBBR must be considered for proper design, as these have a significant effect on reactor performance. The most significant factors are
(1) Organic loading, (2) Dissolved oxygen concentration, (3) Ammonia concentration, (4) Wastewater temperature, and (5) pH/alkalinity.
Figure 5.6 illustrated the importance of removing the organic matter present in the wastewater before reasonable removal rates will be achieved in the downstream nitrification MBBR. Otherwise, heterotrophic biofilm competition for space and oxygen reduces (or eliminates) the nitrifying activity of the biofilm. As the organic loading is reduced, the nitrification rate increases, until the dissolved oxygen concentration becomes rate-limiting. Substrate availability (ammonia) only becomes ratelimiting when concentrations in the reactor are low (<2 mg N/L) and therefore is a concern when substantially complete nitrification is required. In these cases, two reactors in series may be considered, whereby the first reactor is oxygen-limited and the second reactor is ammonia-limited. As with all biological processes, temperature has a marked effect on the nitrification rate, but this can be offset with MBBRs by
Moving-Bed Biofilm Reactors
operating with a higher bulk dissolved oxygen concentration. Finally, at low alkalinity levels, the nitrification rate within the biofilm will begin to show inhibition. The following subsections discuss each of these important design factors. In applications where sufficient alkalinity and ammonia is present (at least initially), the nitrification rate will increase as organic loading is reduced, until the dissolved oxygen concentration becomes rate-limiting. In well-established nitrifying MBBR biofilms, the availability of oxygen limits the rate of nitrification on the carriers, as long as the O2-to-NH4-N ratio is below 2.0 (Hem et al., 1994; Odegaard et al., 1994). Thus, unlike suspended-growth systems, the reaction rate in a moving-bed reactor exhibits a linear or near-linear dependence on the bulk dissolved oxygen concentration under oxygen-limited conditions. This observed behavior is likely from the rate at which oxygen diffuses through the stagnant liquid layer and penetrates the biofilm (Hem et al., 1994). A higher bulk oxygen concentration increases the concentration gradient through the biofilm. The increased mixing energy under higher aeration velocities also helps to improve transfer from the bulk liquid to the biofilm. From Figure 5.6a, if organic loading is held constant (i.e., biofilm thickness and composition), the nitrification rate can be expected to increase linearly with an increasing dissolved oxygen concentration. Figure 5.7 illustrates that until fairly low bulk liquid ammonia concentrations are observed in the reactor, an increase in bulk dissolved oxygen level helps increase the nitrification rate.
Nitrification rate, g NH4-N/m2d
1.4 1.2 1.0
15 deg. C 0.4 g BOD5/m2/d
DO = 6 mg/L
DO = 4 mg/L
0.8 0.6
DO = 2 mg/L
0.4 0.2 0.0 0.0
0.5 1.0 1.5 2.5 2.0 Ammonium concentration, mg NH4-N/L
3.0
Figure 5.7 Influence of dissolved oxygen at low ammonia concentrations (Rusten and Odegaard, 2007).
225
226
Biofilm Reactors
In a well-established “pure” nitrifying biofilm, bulk liquid ammonia concentration does not begin to affect the reaction rate until the O2:NH4-N ratio enters the range 2 to 5. Examples of reported O2:NH4-N ratios are listed in Table 5.5. The transition value of 3.2 often is used as a starting point for the basis of design. The ammonia concentration at this adjusted substrate-limited transition value then is used with the following overall reaction rate equation to estimate an appropriate rate for the basis of design:
rNH3 − N = k × (SNH3 − N )( n )
(5.3)
Where rNH3-N = nitrification rate (g NH3-N/m2·d), k = reaction rate constant (site-/temperature-dependent), SNH3-N = rate-limiting substrate concentration (mg N/L), and n = reaction order (site-/temperature-dependent). The reaction rate constant (k) at a given dissolved oxygen concentration is dependent on biofilm thickness and on the diffusion coefficient of the limiting substrate. The reaction order (n) is dependent on the liquid layer adjacent to the biofilm. The reaction order tends toward 0.5 when turbulence is high and the stagnant liquid layer is thin and tends toward 1.0 when turbulence is low and the stagnant liquid layer is thick, resulting in diffusion being rate-limiting (Salvetti et al., 2006). The ammonia concentration at the transition (SNH3-N) can be estimated using the transition ratio and design bulk dissolved oxygen level (see below). A higher design bulk dissolved oxygen will help to reduce the transition ratio, but just slightly (Hem et al., 1994). In addition, an allowance for oxygen depletion across the heterotrophic
Table 5.5 Examples of reported O2:NH4-N. Reference
O2:NH4-N
Hem et al. (1994)
Less than 2 (oxygen-limited) 2.7 (transition: O2 = 9 to 10 mg/L) 3.2 (transition: O2 = 6 mg/L) Greater than 5 (ammonia-limited)
Bonomo et al. (2000)
Greater than 3 to 4 (ammonia-limited) Less than 1 to 2 (oxygen -limited)
Moving-Bed Biofilm Reactors
layer on the biofilm should be considered based on heterotrophic competition for space resulting from reactor loading and mixing conditions. (SNH3 − N ) = 1.72 mg − N / l =
6 mg O 2 / l − 0.5 mg O 2 / l O2 3.2 NH 3 − N
Based on a SNH3-N of 1.72 and assuming an example reaction rate constant k = 0.5 and reaction order of 0.7 for the reactor condition, the above presented overall reaction rate equation gives the following:
rNH3 − N = 0.73 g / m 2 / d = 0.5 × (1.72)( 0.7 )
Several temperature-related factors are important when examining the effect of temperature on nitrification in MBBR systems. Consider that wastewater temperature in an MBBR will affect intrinsic biological nitrification kinetics; affect the rate of diffusion of substrate into and out of the biomass; affect the viscosity of the liquid, which, in turn, may influence the effect of shear-energy on biofilm thickness; and affect the solubility of oxygen in the liquid. The effect of temperature on the observed overall reaction rate presented above can be adjusted by the following relationship: kT2 = kT1 ⋅ θ (T2 − T1 ) Where kT1 = reaction rate constant at T1, kT2 = reaction rate constant at T2, and θ = temperature coefficient.
(5.4)
Examples of reported temperature-dependency coefficients are listed in Table 5.6. Though dependency of nitrification kinetics on temperature can reduce nitrification rates at the governing “wintertime” design temperature, this effect can be offset
Table 5.6 Examples of reported temperature-dependency coefficients. Reference
Temperature coefficient
Rusten et al. (1995)
Θ = 1.09 (ammonia-limited)
Salvetti et al. (2006)
Θ = 1.098 (ammonia-limited)
Salvetti et al. (2006)
Θ = 1.058 (oxygen-limited)
227
Biofilm Reactors
with MBBRs by the combined effect of higher attached biofilm concentrations on the carriers typically observed at colder temperatures and by maintaining higher bulk dissolved oxygen levels in the reactor. At colder wastewater temperatures, a higher biomass (g/m2) commonly is observed. Also, higher bulk dissolved oxygen levels can be attained without an increase in aeration velocity, as a result of its higher solubility at lower liquid temperatures. In net, the nitrification activity per unit surface area of carrier can be maintained effectively despite a reduction in the specific biofilm activity (g NH3-N /m2·d ÷ g SS/m2). Figure 5.8a shows the seasonal observation of increasing and decreasing biomass as a function of wastewater temperature at a tertiary nitrifying MBBR installation. As shown, the biomass concentration drops significantly between May and June as the wastewater temperature warms from <15°C to >15°C. Two data clusters are grouped by wastewater temperature conditions (>15°C and <15°C) and are shown on Figure 5.8b. Though the specific biofilm activity is lower for the <15°C data set, overall performance is maintained by a higher overall biomass concentration and by a higher bulk dissolved oxygen level as a result of an increase in gas solubility at lower temperatures. The observation helps depict that an overall surface area reaction rate on the carriers can be maintained during colder temperature conditions offset by an adapted biofilm, despite a reduction in the nitrifier growth rate.
(b) 25
Biomass Temp
20 15 10 5 0
D
J
F
M
A
M
J Month
J
A
S
O
N
D
Nitrification Activity (g-NH3-N/m2/d / g-SS/m2)
(a)
Biomass (g/m2), Temperature (C)
228
0.25 < 15 C
> 15 C
< 15 C 2 NH3-N SALR (AVG) = 1.1 g/m /d Effluent NH3-N (AVG) = 2.4 mg-N/L >15 C 2 SALR (AVG) = 1.2 g/m /d Effluent NH3-N (AVG) = 6.9 mg-N/L
0.20 0.15 0.10 0.05 0.00
5.0
6.0
7.0
8.0
9.0
10.0
11.0
Bulk Dissolved Oxygen (mg/L)
Figure 5.8 (a) Seasonal biomass concentration and temperature in tertiary nitrifying MBBR (Zimmerman, 2007), and (b) nitrification activity grouped by temperature condition as a function of dissolved oxygen level (Zimmerman, 2007).
Moving-Bed Biofilm Reactors
3.1.6 Denitrification Moving-bed reactors have been used successfully for denitrification in pre-, post-, and combined configurations. As with other biological denitrification processes, the main factors that must be considered in design are the following:
(1) Presence of a suitable carbon source in a proper carbon-to-nitrogen ratio in the reactor, (2) Level of denitrification required, (3) Wastewater temperature, and (4) Oxygen carryover from upstream process steps or recirculation flows.
3.1.6.1 Pre-Denitrification Moving-Bed Biofilm Reactors Pre-denitrification with MBBRs is typically well-suited in instances where BOD removal, nitrification, and a moderate level of nitrogen removal are needed. The influent to the reactor should have a favorable ratio of easily biodegradable COD and ammonia (C:N) to make efficient use of the anoxic reactor volume. The dissolved oxygen returned in the recycle flow can have a significant influence on performance with MBBR systems, because elevated bulk dissolved oxygen levels can be required for the nitrification stage of the MBBR process. This can place a practical upper limit on the most effective recirculation ratio (Qrcy:Qinf), whereby a further increase in the recirculation rate ends up reducing the overall denitrification efficiency. Where influent wastewater characteristics are suitable for pre-denitrification, nitrogen removal performance typically can range between 50 and 70%, with a Qrcy:Qinf ratio of 1:1 to 3:1 (see Williams Monaco WWTP [Henderson, Colorado] case study below). Some examples of typical pre-denitrification rates that have been observed with municipal wastewater are listed in Table 5.7. Table 5.7 Examples of typical pre-denitrification rates observed with municipal wastewater. Reference
Denitrification rate (NO3-N equivalents)
Rusten and Odegaard (2007)
0.40 to 1.00 g/m2·d (full-scale at Gardemoen WWTP, Ullensaker, Norway)
Rusten et al. (2000)
0.15 to 0.50 g/m2·d (pilot testing at FREVAR WWTP, Fredrikstad, Norway)
McQuarrie and Maxwell (2003)
0.25 to 0.8 g/m2·d (pilot testing at Crow Creek WWTP, Cheyenne, Wyoming)
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Biofilm Reactors
In practice, variation in the observed rate can be the result of site-specific and seasonal differences in wastewater character (i.e., C:N), dissolved oxygen carryover into the reactor, and wastewater temperature.
3.1.6.2 Post-Denitrification Moving-Bed Biofilm Reactors Post-denitrification with MBBRs may be considered in instances where the degradable carbon naturally present in the wastewater is insufficient or has already been depleted by an existing upstream process, or when limited space availability at the plant-site calls for compact high-rate denitrification. Because performance is not limited by factors such as internal recycle rate or carbon availability, high levels of nitrogen removal (>80%) can be achieved within a minimum HRT. If there are stringent limits on effluent BOD and low nitrate levels are required, it may be necessary to follow post-denitrification with a small post-aeration MBBR. Practical experience has shown that post-denitrification performance in MBBRs can be inhibited when precipitation treatment upstream of the reactor results in phosphorus-limiting conditions for cell synthesis (see Klagsham WWTP [Malmo, Sweden] case study below). Maximum nitrate SARRs with external carbon addition can be greater than 2 g/ m2·d when carbon is added in excess. An indication of rates for different carbon types and the effect of temperature are shown in Figure 5.9.
3.1.6.3 Combined Pre-/Post-Denitrification Moving-Bed Biofilm Reactors oving-bed reactors can be combined to take advantage of the economy proM vided with pre-denitrification and the performance of post-denitrification. The Ethanol Methanol Monopropylene glycol
4 Denitrification rate g NO-Nequiv./m2d
230
3 2 1 0
4
6
8
10
12
14
16
Temperature, °C
Figure 5.9 Denitrification rate as a function of temperature with different external carbon types (Rusten et al., 1996).
Moving-Bed Biofilm Reactors
re-denitrification reactor can be designed so that it can be operated as an aerated p reactor during the winter months when
(1) Additional aerated reactor volume would be helpful for improving nitrification performance; and (2) The cold wastewater temperature results in high dissolved oxygen levels and lower soluble CODs, which otherwise would impede pre-denitrification efficiency.
During the winter months, the post-denitrification reactor would be relied on to provide all required nitrate removal (see Gardermoen WWTP [Gardermoen, Norway] case study below).
3.2 Mixers In denitrification MBBR applications, rail-mounted submersible mechanical mixers have been used to circulate and mix the contents of the reactor and mix the carrier media. Some special considerations are required for a proper design. These are (1) location and orientation of mixers, (2) type of mixer, and (3) mixing energy. With a specific gravity of approximately 0.96, the biofilm carriers will float when there is no energy added to the reactor. This is different from suspendedgrowth systems where, in the absence of energy, the solids will settle. As a result, the mixers needed to be located toward the surface of the reactor, but not so close as to create a vortex on the surface, which would entrain air into the reactor. The mixer unit needs to have a slight negative inclination, to help push the media down into the lower depths of the reactor, as shown in Figure 5.10. In typical applications, unaerated MBBRs require between 25 and 35 W/m3 to mobilize the floating bed of media. Special mixing considerations are required for denitrification MBBRs. Not all mixers are suitable for long-term use in an MBBR. The mixer manufacturer (ABS, Meriden, CT) collaborated with the AnoxKaldnes and utility at several MBBR installations to develop a mixer suited to withstand the abrasive nature of the carriers on the mixer propeller. The ABS 123K mixer was developed from this work, specifically designed to handle moving-bed reactor applications. The mixer uses a stainless-steel backward-curve-design propeller with a 12-mm round bar welded along its leading edge to avoid damage to the media and wear to the propeller. The mixer also uses a fairly low rotational speed (90 rpm at 50 Hz and 105 rpm at 60 Hz) in this application.
231
232
Biofilm Reactors
(a)
(b)
Figure 5.10 (a) ABS 123K mixers installed toward the surface of the reactor oriented with a 30-degree negative inclination to “push” the media down into the depths of the reactor, and (b) post-denitrification MBBR reactor in operation at the Sjölunda WWTP (Malmö, Sweden). The mixing energy required to properly mix a denitrification MBBR depends on the fill fraction of media and the biofilm growth anticipated in the reactor. Practical experience has shown that reactor mixing is more effective at lower fill fractions (e.g., <55%). A high fill fraction can make it difficult for a mixer to successfully circulate the media and should be avoided. Instead, a lower fill fraction with higher surface area loadings—and hence a higher biomass concentration—on the carriers helps to weigh down the media, making it easier for the mixers to circulate the media within the reactor. For this reason, it is important to select a size of denitrification reactor such that the fill fraction is amenable to mechanical mixing.
3.3 Pretreatment As with many of the other submerged biofilm technologies, MBBRs need to have proper preliminary treatment upstream of the reactor. Good screening and grit removal is required to prevent long-term accumulation of undesirable inert materials in the MBBR, such as rags, plastics, and sand. These materials are difficult to remove once they have entered the reactor, because the reactor is partly filled with carrier media. Manufacturers generally recommend no larger than 6-mm spacing if primary treatment also is provided. Still, finer screening at 3mm or less is mandatory in instances where primary treatment is not provided. On the other hand, add-on MBBR processes, where significant treatment already occurs upstream, likely do not need additional screening. Table 5.8 provides a summary of sample screening installations at several facilities where MBBRs are incorporated to the treatment scheme.
Moving-Bed Biofilm Reactors
233
Table 5.8 Screening examples at MBBR installations. MBBR facility
Pretreatment
Screen details
Lillehammer WWTP (Norway)
Step screens, grit removal, sedimentation
15 mm followed by 3 mm
Gardemoen WWTP (Norway)
Step screen, grit removal, sedimentation
6 mm
Crow Creek WWTP (Cheyenne, Wyoming)
Filter screen, grit removal, sedimentation
10 mm × 15 mm
Western WWTP
Step screen, grit removal
3 mm
Mao Point WWTP (New Zealand)
Step screen, grit removal, sedimentation
3 mm
Table 5.9 Solids separation examples at MBBR installations. MBBR facility
Follow-up separation
Design rate (average peak)
Lillehammer WWTPa
Flocculation/settling
1.3 to 2.2 (m3/m2·h)
Nordre Follo WWTPa
Flocculation/flotation
5 to 7.5 (m3/m2·h)
Gardemoen WWTPa
Flocculation/flotation
3.1 to 6.4 (m3/m2·h)
Crow Creek WWTPa
Reused existing clarifiers
1.1 to 2.2 (m3/m2·h)
Moorhead WWTP (Minnesota)b
None
a
Multi-stage MBBRs.
b
Tertiary nitrification MBBR.
N/A
4.0 Solids Separation Compared with suspended-growth systems, the moving-bed process offers considerable flexibility, in terms of the type of process used for follow-up solids separation. Biological process performance is independent of the solids separation step, which expands the array of options available for consideration. Plus, the concentration of solids in the reactor effluent is at least an order of magnitude lower in concentration. As a result, a variety of different solids separation processes have been paired successfully with MBBRs. Some representative examples are summarized in Table 5.9 In cases where space is a factor, MBBRs can be combined with a compact high-rate solids separation technology, such as dissolved air flotation (DAF) or inclined plate settlers. In the case of a retrofit, existing clarifiers may be reused for solids separation.
234
Biofilm Reactors
5.0 General considerations for moving-bed biofilm reactors The following general considerations are important for design of an MBBR system.
5.1 Approach Velocity The expected peak velocity (flowrate divided by cross-sectional area of reactor) during high-flow conditions through an MBBR reactor must be considered. At lower approach velocities (e.g., 20 m/h), the carriers will stay evenly distributed throughout the reactor. However, at higher velocities (e.g., >35 m/h), the carriers will tend to migrate and stack against the media retention sieves, resulting in high headloss, creating an unacceptable condition. In some designs, reactor geometry and number of trains may be dictated by peak hydraulic conditions. It is important to consult with the manufacturer regarding acceptable design approach velocities. The length-towidth (L:W) ratio of the reactor can be a factor also. In general, lower reactor L:W (e.g., 1:1 or less) will help to reduce the amount of media migration that occurs at peak flow conditions.
5.2 Foaming Though uncommon, foaming episodes occasionally may occur in MBBRs, especially during startup or process upset. The foam will become trapped in the MBBR reactor, because the partition wall between two successive reactors extends above the water surface. It is recommended that a defoaming agent is used when foam must be controlled in a reactor. Use of a defoamer will coat the carrier media, impeding diffusion to the biofilm, which may affect performance. It is extremely important to use defoamers that do not use silica-compound agents, as these types of defoamers are incompatible with plastic-carrier media.
5.3 Media Transfer and Inventory Management Though there is little to go wrong inside a well-constructed and designed moving-bed reactor, it is still prudent to consider how one might manage transferring and storing the media on-site should a reactor ever need to be taken out of service for in-basin maintenance. The contents of a reactor, including carriers, can be transferred using a 10-cm (4-in.) recessed impeller pump. Depending on the design fill fraction, it may be possible to transfer the media to another reactor. The drawback with this is that it is difficult to re-apportion the media once two reactors are combined. After pumping
Moving-Bed Biofilm Reactors
the media back to the reactor, the only reasonably accurate method for measuring the media fill fraction is to dewater the reactor and measure the height of the media in both reactors to check the fill fraction. Preferably, another tank or idle unit process could be used for temporary storage of the media. This way, it is easier to ensure that the proper fill fraction has been restored to the reactor.
6.0 CASE STUDIES Six case study examples of moving-bed reactors help provide a cross-section of how the technology has been applied to meet a variety of different treatment objectives and situations. These cases help to highlight the innovation and ingenuity offered by the technology.
(1) BOD removal: • Moa Point WWTP (Wellington, New Zealand)—MBBR/solids contact
(2) Nitrification: • Harrisburg WWTP (Harrisburg, Pennsylvania)—tertiary MBBR (pilot study) • Moorhead WWTP (Moorhead, Minnesota)—tertiary MBBRs
(3) Nitrogen removal: • Williams Monaco WWTP (Henderson, Colorado)—multiple MBBRs with pre-denitrification • Klagsham/Sjölunda WWTPs (Malmö, Sweden)—post-denitrification MBBRs • Gardemoen WWTP (Gardemoen, Norway)—multiple MBBRs with preand post-denitrification
6.1 Moa Point Wastewater Treatment Plant, Wellington, New Zealand In 1998, the MBBR/solids contact process at the Moa Point WWTP was placed into operation. This innovative high-rate process was selected because it offered a viable compact solution for providing 80% BOD removal and TSS of approximately 71 000 m3/d (18.75 mgd) of flow (259 000-m3/d [68.5-mgd] peak) on an extremely spaceconstrained site. The process approach offered a second benefit, in that the small size of the reactors could be covered for odor control. The plant treats wastewater
235
236
Biofilm Reactors
from the North Island City of Wellington before discharge by a 1.9-km outfall pipe to the Cook Strait. During wet-weather events, peak flows above approximately 259 000-m3/d (68.5-mgd) are bypassed around the MBBR/solids contact process. A simplified process schematic of the Mao Point WWTP is provided in Figure 5.11, and Table 5.10 provides a summary of the MBBR/solids contact system. The MBBR provides a roughing step to quickly remove a significant fraction of the readily biodegradable soluble and easily hydrolyzed organic matter from the
RE-AERATED SLUDGE
PRIMARY SEDIMENTATION
INFLUENT
SCREEN
SECONDARY TREATMENT
GRIT REMOVAL MBBR
SETTLED SLUDGE TO SLUDGE REAERATION OR DEWATERING PLANT
EFFLUENT
SOLIDS CONTACT AERATION SETTLED SLUDGE TO SLUDGE REAERATION OR DEWATERING PLANT
Figure 5.11 Simplified liquid treatment schematic of Moa Point WWTP (New Zealand).
Table 5.10 MBBR/solids contact process at Moa Point WWTP (New Zealand). Description
Design details
MBBR/solids contact process trains Number of basin trains
3
Number of MBBR reactors per train
1
Volume per MBBR reactor
919 980 L (32 485 cu ft)
Number of solids contact reactors per train
2
Volume of solids contact reactors per train
1 940 000 L (68 500 cu ft)
Media Type Reactor fill fraction (%)
Kaldnes K1 (AnoxKaldnes, Inc.) 30
Moving-Bed Biofilm Reactors
primary effluent before the solids contact reactors. The biological slough, which is carried with the unsettled MBBR effluent, is mixed with the return sludge from the secondary clarifiers as it enters the solids contact reactors. The treated effluent mixed liquor flocculates and settles in the clarifiers. Secondary treatment standards require that the geometric mean of 90 consecutive samples shall not exceed 20 mg/L for BOD and 30 mg/L for TSS, and no more than 10% of those consecutive values can exceed 45 mg/L for BOD and 68 mg/L for TSS. As the facilities came on-line, there was a period from January 1999 through March 2000 when two of the three MBBR/solids contact process trains were in operation handling all of the incoming flow. During this time period, organic loading conditions on the MBBRs offered an opportunity to observe system performance at a high surface loading rate (see Figure 5.5). Figure 5.12 shows the BOD profile across the stages of treatment at Moa Point WWTP from January 1999 through March 2000.
Moa Point WWTP, New Zealand Influent Total BOD, MBBR Effluent Total BOD & Soluble BOD and Final Effluent Total BOD from 1/5/99 - 4/11/00
500
Raw BOD MBBR Effluent Total BOD MBBR Effluent Soluble BOD Final Effluent Total BOD
450 400
BOD (mg/L)
350 300 250 200 150 100 50
Date
Figure 5.12 BOD profile results at Moa Point WWTP (New Zealand).
4/11/2000
3/28/2000
3/14/2000
2/29/2000
2/1/2000
2/15/2000
1/4/2000
1/18/2000
12/7/1999
12/21/1999
11/9/1999
11/23/1999
10/26/1999
9/28/1999
10/12/1999
9/14/1999
8/31/1999
8/3/1999
8/17/1999
7/6/1999
7/20/1999
6/8/1999
6/22/1999
5/25/1999
5/11/1999
4/27/1999
4/13/1999
3/30/1999
3/2/1999
3/16/1999
2/2/1999
2/16/1999
1/5/1999
1/19/1999
0
237
Biofilm Reactors
6.2 Harrisburg Wastewater Treatment Plant, Harrisburg, Pennsylvania The Harrisburg Authority conducted an extended pilot-scale study to evaluate MBBR as an add-on unit process for tertiary nitrification. The study was performed as an option to upgrade the utility’s existing 88 900-m3/d (23.5-mgd) high-purity-oxygen (HPO) activated sludge plant designed for high-rate secondary treatment to comply with a pending stringent effluent ammonia requirement (<3 mg N/L). To conduct the pilot study, settled secondary effluent from the full-scale HPO plant was pumped to a pilot unit consisting of two MBBR reactors in series. The roughly 6-month study timeline was divided into nine test periods. The biofilm carrier fill fraction was reduced after test period 3 to increase the surface area loading rate condition and the flowrate to the pilot unit was adjusted up and down several times, changing the HRTs and loading rates accordingly. Figure 5.13 shows ammonia profile data for the two reactors over the nine test periods. Note that, during the first test period (Figure 5.13a), the two reactors were loaded lightly, and effluent ammonia from the first reactor typically was already below 1 mg N/L. The loading rate was increased for test periods 2 and 3, resulting in some carryover of ammonia from the first to second reactor, but with no effect on overall effluent ammonia levels. For test periods 4 through 9, the reactors were operated at a reduced fill fraction, to increase loading rates on the carriers. At the higher carrier loading conditions, the two reactors in series provided the level of nitrification required to comply with the target effluent quality objective of 3 mg NH3-N/L.
2
6 3
0 DATE
Figure 5.13 (a) Test periods 1, 2, and 3, and (b) test periods 4 through 9.
24 31 6/ 7 5/
17
0
5/
16 3/
2
9
3/
3/
23
16 2/
2/
2
9
2/
2/
26
19
1/
1/
5
12
1/
1/
12
/2
2
0
9
0
9
10
FLOW (GPM)
1
4
12
5/
5
15
5/
2
6
18
19 26 5/ 3
10
21
4/
3
10 8
12
15
FLOW
24
4/
4
R2 NH3-N
27
4/
20
R1 NH3-N
30
29 4/ 5
5
INFLUENT NH3-N
33
22
25
DATE
(b)
6
3/
FLOW
3/
R2 NH3-N
NH3-N CONCENTRATION (mg/L)
R1 NH3-N
FLOW (GPM)
INFLUENT NH3-N
30
/2
NH3-N CONCENTRATION (mg/L)
(a)
12
238
Moving-Bed Biofilm Reactors
Figure 5.13 helps to highlight the benefits of having two nitrification reactors in series, if the application requires compliance with stringent effluent ammonia concentrations. As the figure shows, with two reactors in series, the first reactor can be highly loaded, while the second reactor is capable of absorbing a wide range of influent ammonia concentrations, while consistently producing a substantially nitrified effluent. Paired data from the nine test periods is presented in Figure 5.14 to show SARR as a function of SALR for reactors 1 and 2 and both reactors combined. The figure helps to show the following:
(1) Substantially complete nitrification was achieved within a single reactor (reactor 1) at low loadings up to a rate of 0.7 g NH3-N/m2·d; (2) With two reactors in series, substantially complete nitrification was achieved with the second reactor (reactor R2) being loaded up to 1.2 g NH3-N/ m2·d; and (3) The maximum removal rate under high ammonia loadings was approximately 1.8 NH3-N/m2·d.
2.2 2.0
Combined Reactors
Reactor 1
Reactor 2
Removed - SARR (g NH3-N/m2/d)
1.8 1.6 1.4 1.2 1.0 0.8 0.6 0.4 Average Temperature = 14 C 90th Percentile = 18.7 C 10th Percentile = 10.7 C
0.2 0.0
0.0 0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.0 2.2 2.4 2.6 2.8 3.0 3.2 3.4 3.6 3.8 4.0 4.2 4.4 Applied - SALR (g NH3-N/m2/d)
Figure 5.14 Ammonia removal rate as a function of loading rate in a tertiary MBBR (Kaldate, 2007).
239
240
Biofilm Reactors
6.3 Moorhead Wastewater Treatment Facility, Moorhead, Minnesota (Adapted from Zimmerman et al., 2004) In 2003, the City of Moorhead, Minnesota, completed full-scale implementation of a tertiary nitrifying MBBR added as an upgrade to its existing 23 000 m 3/d (6-mgd) wastewater treatment facility (WWTF). The upgrade was driven by a seasonal requirement to comply with a moderate level of ammonia removal (approximately 8 mg N/L) imposed for discharges to the Red River of the North during low-flow conditions. The MBBR receives non-nitrified settled secondary effluent from the WWTF’s HPO treatment process. The moving-bed process was selected over other potential upgrade options, because it maintained the existing plant capacity, conserved site space, can be expanded easily to meet future demands by adding more carriers, used an existing tank and fit within the plant’s existing hydraulic profile, and was comparable in cost with other feasible alternatives. A simplified process schematic of the Moorhead WWTF is provided in Figure 5.15, and Table 5.11 provides a summary of the MBBR system. cl
2
Influent
Screen
Secondary Treatment
Grit removal Hpo activated sludge Settled sludge to anaerobic digestion
OX Polishing pond
sog
OX
N-MBBR
Effluent Polishing pond
Polishing pond
Settled sludge to anaerobic digestion
Figure 5.15 Simplified liquid treatment schematic of Moorhead WWTF (Minnesota).
Table 5.11 Tertiary MBBR process at Moorhead WWTF (Minnesota). Description
Design details
MBBR trains Number of basin trains
1
Number of MBBR reactors per train
1
Volume per MBBR reactor
2 970 200 L (104 880 cu ft)
Dimensions (L × W × D)
42 m × 24 m × 2.9 m (138 ft × 80 ft × 9.5 ft)
Media Type
IDI ActiveCell (Richmond, Virginia)
Reactor fill fraction (%)
32
Moving-Bed Biofilm Reactors
Figure 5.16 shows ammonia profile performance of the MBBR. The tertiary nitrifying MBBR system was installed as a retrofit inside a single existing tank located between two existing polishing lagoons.
6.4 Williams Monaco Wastewater Treatment Plant, Henderson, Colorado In 2003, the South Adams County Water and Sanitation District placed into operation its multistage MBBRs at the Williams Monaco WWTP. This installation provides an example of a design that uses multiple MBBRs in series to provide basic secondary treatment, nitrification, and partial nitrogen removal. The new MBBRs were part of an upgrade project that increased the overall capacity of the plant to approximately 26 000 m 3/d (7 mgd) and helped the Williams Monaco WWTP meet a new permit regulation requiring removal of ammonia down to approximately 10 mg N/L (30-day average). Equally important was to select a process that would improve the plant’s ability to reliably comply with a BOD limit of
45.0
NH3-N Concentration, mg-N/L
40.0 35.0 30.0
MBBR Influent
25.0 20.0 15.0 10.0 MBBR Effluent 5.0 0.0
D
J
F
M
A
M
J
J
A
S
O
N
D
Figure 5.16 Ammonia profile results across nitrifying MBBR at the Moorhead WWTF (Minnesota).
241
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Biofilm Reactors
25 mg/L and TSS of 30 mg/L. The District elected to include pre-denitrification MBBRs to help position the facility for meeting future limits on nitrate and nitrogen. Before the upgrade, the Williams Monaco WWTP consisted of three parallel treatment trains (one rock media trickling filter, one biotower, and one train of rotating biological contactors). Only the biotower was capable of providing seasonal nitrification. As a result, the District needed to explore upgrade options to meet the pending limit on ammonia. A number of alternatives were considered, including activated sludge. However, MBBRs ultimately were selected, because the technology offered a cost-effective approach that was compatible with the existing secondary clarifiers and fit with staff experience with attached-growth processes facilities. Yet, in terms of performance, the MBBR option positioned the facility for both its near-term ammonia and future nitrogen requirement. A simplified process schematic of the Williams Monaco WWTP is provided in Figure 5.17. Figure 5.18 shows an overview of the reactors. The rotating biological contactors were demolished during the upgrade project whereas the trickling filters remain as part of the treatment scheme. Table 5.12 provides a summary of the multistage MBBR system. Each MBBR train consists of four reactors in series—two anoxic reactors followed by two aerated reactors. Nitrate-rich recirculation flow from the effluent end
Secondary clarifier Trickling filter
Influent
Screen
Primary clarifier
Settled sludge to primary clarifier
Primary effluent splitter
Effluent
Secondary clarifier Biotower
Settled sludge to anaerobic digestion
AX
AX
OX
OX
Settled sludge to primary clarifier
C/N/DN MBBRS
Figure 5.17 Simplified liquid treatment schematic of Williams Monaco WWTP (Colorado).
Moving-Bed Biofilm Reactors
Figure 5.18 Two MBBR trains with four reactors each: anoxic (right) and aerated (left).
Table 5.12 Multistage MBBR process at Williams Monaco WWTP (Colorado). Description
Design details
MBBR trains Number of basin trains
2
Number of MBBR reactors per train
4 (2 anoxic, 2 aerated)
Volume per MBBR reactor
579 990 L (20 480 cu ft)
Media Type
Kaldnes K1 (AnoxKaldnes, Inc.)
Reactor fill fraction (%)
55 (anoxic), 60 (aerated)
of each train is recycled to the front end of the train and mixed with the influent entering the train. The rock filter treatment train is part of the original facility and has been out of service since the MBBRs came on-line. The MBBRs were designed to be operated in series or in parallel with the biotower. In the first mode (series), the biotower provides a roughing step to reduce the amount of organic load to
243
Biofilm Reactors
the MBBRs. In this mode of operation, the MBBRs can be operated to achieve substantially complete nitrification. Some primary effluent can be bypassed around the biotower and sent directly to the MBBRs to facilitate pre-denitrification. The MBBR effluent is sent to the original secondary clarifiers for settling. In the second mode (parallel operation), primary effluent flow is split between the biotower and the MBBRs and rejoins before the secondary clarifies. The biological slough from the MBBRs naturally flocculates and settles in the clarifiers without the use of chemicals. Figures 5.19 and 5.20 show settled effluent BOD and TSS quality from the MBBRs from startup through the middle of 2007. The MBBRs were operated with the biotower providing a roughing stage of treatment for most of the primary effluent flow for the majority of 2004. However, plant staff started to notice that snails from
50
Biotower Parallel to MBBR
Biotower / Post MBBR
Biotower Out of Service (All Flow to MBBR)
40 Plant Effluent CBOD (mg/L)
7-day average limit = 40 mg/L
30 30-day average limit = 25 mg/L 20
10
Daily Composite CBOD
30-day Moving Avg (CBOD)
06/15/07
03/16/07
12/15/06
09/15/06
06/16/06
03/17/06
12/16/05
09/16/05
06/17/05
03/18/05
12/17/04
09/17/04
06/18/04
03/19/04
0
12/19/03
244
7-day Moving Avg (CBOD)
Figure 5.19 BOD profile results across MBBRs at Williams Monaco WWTP (Colorado) (2004 to 2007).
Moving-Bed Biofilm Reactors
60
Biotower Parallel to MBBR
Biotower / Post MBBR
Plant Effluent TSS (mg/L)
50
Biotower Out of Service (All Flow to MBBR)
7-day average limit = 45 mg/L
40
30 30-day average limit = 30 mg/L 20
10
Daily Composite TSS
30-day Moving Avg (TSS)
06/15/07
03/16/07
12/15/06
09/15/06
06/16/06
03/17/06
12/16/05
09/16/05
06/17/05
03/18/05
12/17/04
09/17/04
06/18/04
03/19/04
12/19/03
0
7 per. Mov. Avg. (Daily Composite TSS)
Figure 5.20 Settled effluent TSS at Williams Monaco WWTP (Colorado) (2004 to 2007).
the biotower were taking up residence in the MBBR carriers. Though no effects on performance were observed, it was decided to operate the biotower in parallel with the MBBRs, to avoid colonization of the reactors with snails. Within a short period of time after the change to parallel operation, the snails disappeared from the MBBRs, suggesting that the snails were unable to reproduce in the MBBRs. In March 2005, plant staff decided to simplify plant operations and remove the biotower from service completely, thus sending all primary effluent flow directly through the MBBRs. This operational change significantly increased BOD loadings to the system. The operational data from this point forward shows performance of the MBBRs at highly loaded conditions. During the summer months, average effluent ammonia concentrations typically are below 5 mg N/L, while, during the winter months, effluent ammonia concentrations typically are between 5 and 10 mg N/L.
245
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6.5 Klagsham Wastewater Treatment Plant, Malmö, Sweden (Adapted from Taljemark et al., 2004) In the mid-1990s, the City of Malmö, Sweden, proceeded with a plan to add post-denitrification MBBRs to its Klagsham and Sjölunda WWTPs to help comply with pending enhanced nitrogen removal standards. Post-denitrification MBBRs were selected over other nitrogen upgrade alternatives, because the technology provided the most cost-effective and flexible upgrade for both WWTPs. In both cases, the upgrade treatment scheme fit well with the existing facilities and within the constraints of the site. The post-denitrification MBBRs at the 24 000-m3/d (6.3-mgd) (42 400 m3/d [11.2mgd] peak) Klagsham WWTP have been in operation since 1997. Before the eventual installation of post-denitrification MBBRs, full-scale testing was conducted with the plant’s activated sludge process, which showed that the carbon naturally present in the primary effluent would limit nitrogen removal performance in a pre-denitrification system. Alternatively, the pending effluent nitrogen standard could be reached by running activated sludge as a single-sludge process with nitrification and post-denitrification using external carbon. However, by adding post-denitrification MBBRs, the process design could be optimized by running the activated sludge process with pre-denitrification for partial nitrogen removal, followed by polishing removal with the MBBRs using external carbon. This strategy allowed for some savings in the amount of external carbon required. Further, the reactor volume designated for pre-denitrification can be varied, depending on the seasonal change in reactor volume required for nitrification. At the time of the upgrade, downflow filters of the conventional two-media type also were added to Klagsham WWTP for improved phosphorus removal. The post-denitrification MBBRs were inserted between the activated sludge process and the filters. The sand and anthracite filters have proven to be effective in accepting and removing the suspended solids carried in the MBBR effluent. The effluent standards for the plant are rather stringent, requiring less than 0.3 mg/L total phosphorus, 12 mg/L total inorganic nitrogen, and 12 mg/L BOD. A simplified process schematic of the Klagsham WWTP is provided in Figure 5.21 and Table 5.13 provides a summary of the post-denitrification MBBR system. The post-denitrification MBBR process is comprised of two trains, each with two compartments in series. The design figure for the process is 440 kg NO3-N/d at 12°C, corresponding to a loading rate of 1.7 g N/m2·d. The relatively low fill fraction is advantageous, in that it mixes well within the contents of the reactor. The mixing
Moving-Bed Biofilm Reactors
Ferric chloride
Influent
Screen
Grit removal
Ferric chloride pax
Ethanol
Primary treatment
Secondary treatment Aeration basins
Settled sludge to anaerobic digestion
AX
Effluent
AX
DN MBBR
Two-media filter
Settled sludge to anaerosic digestion
Figure 5.21 Simplified liquid treatment schematic of Klagsham WWTP (Sweden).
Table 5.13 Post-denitrification MBBR process at Klagsham WWTP (Sweden). Description
Design details
MBBR trains Number of basin trains
2
Number of MBBR reactors per train
2
Total volume of MBBR reactors
~1 440 000 L (~51 000 cu ft)
Media Type
Kaldnes K1 (AnoxKaldnes, Inc.)
Reactor fill fraction (%)
36
is performed by two submerged propeller mixers in the first compartment and one mixer in the second compartment because of the different sizes of the compartments. The mixing units draw 31 W/m3 for mixing. External carbon is added at the inlet of the first compartment in each train. In the 1990s, extensive research was performed at the Klagsham WWTP regarding different external carbon sources for denitrification. Ethanol offered an advantage, in that a higher COD content is provided compared with methanol (2.1 versus 1.5 g COD/g) and yields a higher denitrification rate. Furthermore, the bacteria were adapted more easily to ethanol. Ethanol obtained as a waste product from pharmaceutical production is used at the Klagsham plant, covering the total need. A control system based on incoming wastewater flow and an on-line nitrate analyzer was installed to calculate the incoming nitrate load. The desired amount of carbon source
247
248
Biofilm Reactors
then is based on an adjustable set-point value of the amount of COD added, in proportion to the amount of incoming NO3-N, on a weight basis. Each flow train has its separate dosage pump and flow meter for carbon dose/control. Through the control system, a low effluent nitrate concentration could be maintained and the risk of bleed-through of carbon minimized at the same time. This is important, because the plant must comply with a stringent BOD limit. The carbon dosage set-point is somewhat higher than the theoretical need for denitrification, and the level has been found by practical experience. Sometimes, the dissolved oxygen concentration in the influent to the MBBR process can be rather high at both plants. The process is operated for being BOD limited, because the effluent standards for BOD are rather stringent. Consequently, the influent nitrate is not denitrified fully, and the average effluent nitrate concentration is between 4 and 6 mg/L. The amount of phosphorus present in the influent to the MBBR is found to be crucial for the denitrification process. A breakpoint might be distinguished at approximately 0.3 mg/L dissolved phosphorus. Experience has shown that this corresponds to a concentration of approximately 0.1 mg/L in the effluent. Higher effluent concentrations do not increase the denitrification rate noticeably. However, lower concentrations have been observed to hamper denitrification performance. In these cases, phosphoric acid can be added to the MBBR process to occasionally give the process an extra impetus. If the phosphorus limitation situation is not handled in a proper way, an increase in effluent BOD concentration is obtained, as a result of an unbalanced COD:N:P. As a consequence, the denitrification rate also decreases. Adding phosphoric acid has a quick effect, and improved results could be achieved in a couple of days. At first, all of the added phosphate is consumed to counteract the gained depletion of phosphorus. Gradually, the need becomes satisfied, and a buffer is built up, leading to a slow increase of dissolved phosphorus in the effluent. Soon, the balance is fully restored, assuming that further phosphorus deficiency has been avoided by adjustment of the coagulant dosage in the pre-precipitation stage. Figures 5.22 through 5.24 show the effluent quality from 2000 through 2003 for BOD, total inorganic nitrogen, and total phosphorus, respectively. As shown, the process has been able to comply with its nitrogen removal requirement, while, at the same time, complying with a stringent limit on BOD and phosphorus. Phosphorus is removed by chemical addition at the primary clarifiers and the MBBR effluent just before the multi-media filters. Because of the stringent effluent requirement for phosphorus, the filters need to capture most of the suspended solids. The filter effluent normally contains less than 3 mg/L suspended solids. The loading to the
mg/l 40 36 32 28 24 20 16 12 8 4 0 2000–01–01
2000
2001
2002
2003 2004–01–01
Figure 5.22 Effluent BOD7 results at Klagsham WWTP (Sweden). mg/l 20 18 16 14 12 10 8 6 4 2 0 2000–01–01
2000
2001
2002
2003 2004–01–01
Figure 5.23 Effluent total nitrogen results at Klagsham WWTP (Sweden). mg/l 1 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0 2000–01–01
2000
2001
2002
2003 2004–01–01
Figure 5.24 Effluent total phosphorus results at Klagsham WWTP (Sweden). 249
250
Biofilm Reactors
post-denitrification process during this 2000-to-2003 time period averaged 310 kg NO3-N/d, corresponding to a loading rate of 1.05 g N/m 2·d. The typical C:N averaged 3.9 g COD/g NO3-N.
6.6 Gardemoen Wastewater Treatment Plant, Gardemoen, Norway The Gardemoen WWTP offers an example of a multistage MBBR system combined with chemical addition and high-rate solids separation to provide a high level of nutrient removal within a small footprint. The MBBR design consists of two treatment trains, with each train having seven reactors in series. Several of the reactors were designed to operate in either an aerated or an anoxic mode, which offers the plant operating staff quite a bit of flexibility in how to most effectively and economically meet treatment goals, despite changing seasonal conditions. This installation has been in operation since 1998, handling wastewater generated from the Oslo Airport and several of the surrounding municipalities. The WWTP is rated to handle an average dry-weather flow of 22 000 m3/d (5.8 mgd) and designed to meet stringent discharge requirements for BOD (10 mg/L), total nitrogen (>70% removal), and total phosphorus (0.2 mg/L). A simplified process schematic of the Gardemoen WWTP is provided in Figure 5.25 and Table 5.14 provides a summary of the postdenitrification MBBR system. The Gardemoen WWTP has two MBBR treatment trains, each train with seven MBBRs in series. Reactor 1 is anoxic, designated to denitrify the nitrate-rich recycle flow from reactor 5, while using incoming primary effluent as a carbon source. Reactor 2 can be operated either in anoxic mode or aerated mode, depending on wastewater temperature and the amount of aerated volume required to achieve substantial nitrification. Reactors 3 and 4 are aerated, designated for BOD removal and nitrification. Reactor 5 may be operated with mixers for oxygen depletion or as an aerated reactor, depending on the wastewater temperature and amount of aerated volume required to achieve substantial nitrification. Reactor 6 is used for post-denitrification, with external carbon addition. This reactor can be operated with supplemental carbon in slight excess (high C:N), to favor better reaction rates, because reactor 7 provides post-aeration of the treated effluent, to consume any excess carbon remaining in the effluent from reactor 6. For phosphorus removal, coagulant is added to the unsettled effluent from the MBBRs. After chemical addition, the flow enters flocculation tanks before final solids separation by DAF. The WWTP also seasonally handles the spent aircraft deicing
Moving-Bed Biofilm Reactors
Carbon
Influent
Screen
Grit removal
Primary treatment
AX
AX/ OX
OX
OX
AX/ OX
C/N/DN MBBRS
AX
PAC
Sludge to anaercbic digestion
Effluent
OX Rapid mix Flocculation
DAF
Settled sludge to anaerobic digestion
Figure 5.25 Simplified liquid treatment schematic of Gardemoen WWTP (Norway).
Table 5.14 Multistage MBBRs at Gardemoen WWTP (Norway). Description
Details
Number of treatment trains
2
Number of reactors per train
7
Reactor 1 Reactor 2 Reactor 3 Reactor 4 Reactor 5 Reactor 6 Reactor 7
420 000 L (14 830 cu ft) (anoxic) 420 000 L (14 830 cu ft) (anoxic or aerated) 695 000 L (24 540 cu ft) (aerated) 695 000 L (24 540 cu ft) (aerated) 180 100 L (6360 cu ft) (O2 depletion or aerated) 375 000 L (13 240 cu ft) (anoxic) 109 900 L (3880 cu ft) (anoxic)
Carrier type
Kaldnes K1 (AnoxKaldnes, Inc.)
Fill fraction Reactors 1, 2, 3, 4, and 6 Reactor 5 Reactor 7
60% 42% 51%
fluid collected at the airport and uses this stream when available, as an external carbon source for post-denitrification. The following figures help to summarize representative MBBR performance targeting greater than 70% nitrogen removal. Figure 5.26 presents the influent and effluent ammonia concentrations across the MBBRs over a 12-month period. With the exception of a few data points during the middle of the year, substantial nitrification was achieved. Figure 5.27 shows influent and effluent nitrogen concentrations across the MBBRs over the same 12-month period. The incoming wastewater temperature at
251
80 70
Ammonia, mg-N/L
60 MBBR Influent Ammonia
50 40 30 20 10 0
MBBR Effluent Ammonia
J
F
M
A
M
J
J
A
S
O
N
Date
Figure 5.26 Ammonia profile results across MBBRs at Gardemoen WWTP (Norway) (2001).
90 80
Total Nitrogen, mg-N/L
70
MBBR Influent Nitrogen
60 50 40 30 20
MBBR Effluent Nitrogen
10 0
J
F
M
A
M
J
J
A
S
O
N
Date
Figure 5.27 Nitrogen profile results across MBBRs at Gardemoen WWTP (Norway) (2001). 252
Moving-Bed Biofilm Reactors
the WWTP can drop as low as 10°C or lower during winter. Table 5.15 provides a summary of annual performance, showing that the process meets performance requirements. It should be noted that the WWTP is operated to meet the current standard for >70% nitrogen removal. Still lower nitrogen removal is limited by the availability of carbon substrate in the post-denitrification reactors. During the first 6 months of 2000, guarantee testing was conducted using supplemental carbon, which demonstrated that the process could achieve extremely low effluent total nitrogen concentrations and perform significantly better than what was required to meet permit requirements. Plant effluent total nitrogen showed an average concentration of 2.2 mg N/L, with a median concentration of 1.33 mg N/L (Rusten and Odegaard, 2007).
Table 5.15 Summary of annual performance showing that the process meets erformance requirements. p Parameter
Influent
Final effluent
Total COD (mg/L)
605
29
Suspended solids (mg/L)
279
8
Ammonia (mg N/L)
30.6
0.6
Total nitrogen (mg N/L)
44.4
8
Total phosphorus (mg P/L)
7.1
0.2
7.0 References Bonomo, L.; Pastorelli, G.; Quinto, E.; Rinaldi, G. (2000) Tertiary Nitrification in Pure Oxygen Moving Bed Biofilm Reactors. Water Sci. Technol., 41, 361–368. Hem, L.; Rusten, B.; Odegaard, H. (1994) Nitrification in a Moving Bed Reactor. Water Res., 28, 1425–1433. Kaldate, A., Infilco Degremont, Inc., Richmond, Virginia (2007) Personal communication. Lazarova, V.; Manem, J. (1994) Advances in Biofilm Aerobic Reactors Ensuring Effective Biofilm Activity Control. Water Sci. Technol., 29, 319–327. McQuarrie, J.; Maxwell, M. (2003) Pilot-Scale Performance of the MBBR Process at the Crow Creek WWTP, Cheyenne, Wyoming. Proceedings of the 76th Annual
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Water Environment Federation Technical Exposition and Conference, Los Angeles, California, Oct 11–15; Water Environment Federation: Alexandria, Virginia. Melin, E.; Odegaard, H.; Helness, H.; Kenakkala, T. (2004) High-Rate Wastewater Treatment Based on Moving Bed Biofilm Reactors. Chemical Water and Wastewater Treatment VIII; IWA Publishing: London, United Kingdom, 39–48. Odegaard, H.; Gisvold, B.; Strickland, J. (2000) The Influence of Carrier Size and Shape in the Moving Bed Biofilm Process. Water Sci. Technol., 41, 383–391. Odegaard, H.; Paulsrud, B.; Bilstad, T.; Pettersen, J. (1991) Norwegian Strategies in the Treatment of Municipal Wastewater Towards Reduction of Nutrient Discharges to the North Sea. Water Sci. Technol., 24, 179–186. Odegaard, H.; Rusten, B.; Wessman, F. (2004) State of the Art in Europe of the Moving Bed Biofilm Reactor (MBBR) Process. Proceedings of the 77th Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Oct 2–6; Water Environment Federation: Alexandria, Virginia. Odegaard, H.; Rusten, B.; Westrum, T. (1994) A New Moving Bed Biofilm Reactor—Application and Results. Proceedings of the 2nd International Specialized Conference on Biofilm Reactors, Paris, France, Sept 29–Oct 1; International Association on Water Quality: London, United Kingdom, 221–229. Rusten, B.; Hellstrom, B. G.; Hellstrom, F.; Sehested, O.; Skjelfoss, E.; Svendsen, B. (2000) Pilot Testing and Preliminary Design of Moving Bed Biofilm Reactors for Nitrogen Removal at the FREVAR Wastewater Treatment Plant. Water Sci. Technol., 41, 13–20. Rusten, B.; Hem, L.; Odegaard, H. (1995) Nitrification of Municipal Wastewater in Novel Moving Bed Biofilm Reactors. Water Environ. Res., 67, 75–86. Rusten, B.; Odegaard, H. (2007) Design and Operation of Nutrient Removal Plants for Very Low Effluent Concentrations. Proceedings of the Water Environment Federation Nutrient Removal Workshop, Baltimore, Maryland; Water Environment Federation: Alexandria, Virginia, 1307–1331. Rusten, B.; Wien, A.; Skjefstad, J. (1996) Spent Aircraft Deicing Fluid as External Carbon Source for Denitrification of Wastewater: From Waste Problem to Beneficial Use. Proceedings of the 51st Purdue Industrial Waste Conference, West Lafayette, Indiana, May 6–8; Purdue University: West Lafayette, Indiana.
Moving-Bed Biofilm Reactors
Salvetti, R.; Azzellino, A.; Canziani, R.; Bonomo, L. (2006) Effects of Temperature on Tertiary Nitrification in Moving-Bed Biofilm Reactors. Water Res., 40, 2981–2993. Sen, D.; Copithorn, R.; Randall, C.; Jones, R.; Phago, D.; Rusten, B. (2000) Investigation of Hybrid Systems for Enhanced Nutrient Control, Project 96-CTS-4; Water Environment Research Foundation: Alexandria, Virginia. Taljemark, K.; Aspegren, H.; Gruvberger, N.; Hanner, N.; Nyberg, U.; Andersson, B. (2004) 10 Years of Experiences of an MBBR Process for Post-Denitrification. Proceedings of the 77th Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Oct 2–6; Water Environment Federation: Alexandria, Virginia. Zimmerman, R. (2007) Personal communication. Zimmerman, R. A.; Richard, D.; Costello, J. M. (2004) Design, Construction, Start-Up, and Operation of a Full-Scale Separate Stage Moving Bed Biofilm Reactor Nitrification Process. Proceedings of the 77th Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Oct 2–6; Water Environment Federation: Alexandria, Virginia.
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Chapter 6
Hybrid Processes 1.0 OVERVIEW OF INTEGRATED FIXED-FILM ACTIVATED SLUDGE SYSTEMS 260 1.1 Advantages
261
1.2 Disadvantages
262
2.0 Media Types
262
2.1 Fixed-Media Systems
263
2.2 Free-Floating Media Systems
264
2.2.1 Plastic
264
2.2.2 Sponge
264
3.0 History of Process
265
4.0 Application of INTEGRATED FIXED-FILM ACTIVATED SLUDGE Systems 267 4.1 Fixed Media
267
4.1.1 General Requirements
267
4.1.2 Growth on Media 268 4.1.3 Kinetics
269
4.1.4 Worm Growth
270
4.1.5 Media Breakage
270
4.1.6 Dissolved Oxygen Level 270 257
4.1.7 Mixing
271
4.1.8 Access to Diffusers 271 4.1.9 Odor
271
4.2 Free-Floating Media— Sponge Media 272 4.2.1 General Requirements 272 4.2.2 Screen Clogging
272
4.3 Control of Biomass Growth 273 4.3.1 Loss of Sponges
273
4.3.2 Taking Tank Out-ofService 274 4.3.3 Loss of Solids
274
4.3.4 Air Distribution System 274 4.3.5 Plastic Media
274
4.3.5.1 General Requirements 274 4.3.5.2 Biomass Growth 275 4.3.5.3 Media Mixing 276 4.3.5.4 Screens
276
4.3.5.5 Foaming
276
(continued)
258
Biofilm Reactors
4.3.5.6 Media Replacement 276 4.3.5.7 Taking Tank Out-of-Service 276 4.3.5.8 Worm Growth 277 4.3.5.9 Startup 277 5.0 Process Design 277 5.1 Introduction 277 5.2 Parameters Influencing Organics Removal in the Biofilm of Integrated FixedFilm Activated Sludge Systems 278 5.2.1 Biofilm Flux Rates 278 5.2.2 Removals in Biofilm per Unit of Tank Volume 278 5.3 Parameters Influencing Removals in the MixedLiquor Suspended Solids 281 5.4 Interaction Between the Mixed-Liquor Suspended Solids and the Biofilm 282 5.5 Interaction Between Heterotrophs and Nitrifiers 284 5.6 Design Tools/Procedures 284 5.6.1 Empirical Methods 285 5.6.1.1 EquivalentSludge-Age Approach 285
5.6.1.2 Quantity (Length or Web Surface Area) of Media Approach 286 5.6.2 Rates Based on Pilot Studies 286 5.6.3 Biofilm Rate Model 287 5.6.3.1 Define Range of Flux Rates 287 5.6.3.2 Quantify Removal at Different Mixed-Liquor Suspended Solids Mean Cell Residence Times 287 5.6.3.3 Select Flux Rates Based on Location Along Aerobic Zone 287 5.6.3.4 Calculate the Quantity of Media Required 288 5.6.3.5 Additional Analysis to Finalize a Design 288 5.6.3.6 Application of Kinetics-Based Approach with (continued)
Hybrid Processes
Integrated Fixed-Film Activated Sludge Design Software 288 6.0 Case Studies 288 6.1 Annapolis Water Reclamation Facility, Anne Arundel County, Maryland 288 6.1.1 Original Wastewater Treatment Plant 289 6.1.2 Pilot Study (1993 to 1996) 289 6.1.3 Full-Scale Upgrade for Biological Nutrient Removal (1997 to 2000) 291 6.1.3.1 Pilot Study 291 6.1.3.2 During Construction (1997 to 2000) 293 6.1.3.3 PostConstruction (2000 to 2003) 299 6.2 Westerly Wastewater Treatment Plant, Westerly, Rhode Island 299
6.2.1 Introduction 299 6.2.2 Description of Original Facilities 299 6.2.3 Description of Upgrade 299 6.2.4 Design Criteria 302 6.2.5 Performance of the Integrated Fixed-Film Activated Sludge System 302 6.2.6 Operational Issues 306 6.2.7 Costs 307 6.3 Broomfield Wastewater Treatment Plant, Broomfield, Colorado 307 6.3.1 Introduction 307 6.3.2 Full-Scale Plant Results 308 6.4 Colony Wastewater Treatment Plant, Colony, Texas 308 6.4.1 Introduction 308 6.4.2 Changing Design Conditions 316 6.4.3 Plant Construction and Operation 318 6.4.4 System Flexibility 320 6.4.5 Redworm Predation 321 7.0 References
321
259
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1.0 OVERVIEW OF INTEGRATED FIXED-FILM ACTIVATED SLUDGE SYSTEMS The word hybrid can describe any type of treatment process that combines the features of several different technologies. The focus of this chapter on hybrid processes is on the integrated fixed-film activated sludge (IFAS) process, which combines fixed-film and conventional suspended-growth activated sludge treatment processes. The basic intent of an IFAS process is to provide additional biomass within the reactor volume of an activated sludge process, for the purpose of increasing the capacity of the system or upgrading its performance, as illustrated in Figure 6.1. In fact, the effective mixed-liquor suspended solids (MLSS) concentration essentially can be doubled by using media in an IFAS process. Because the biomass is fixed on a media system, the suspended-growth mixed liquor concentrations are not increased, and the performance of the downstream final clarifiers is not negatively affected by an increase in the solids loading rate. In fact, in many cases, clarifier performance is improved by a reduction in the sludge volume index (SVI), as a result of the fixedfilm growth. Therefore, the IFAS process typically has been considered as an upgrade option in existing treatment plants that must incorporate nutrient removal. The media, and the biomass it supports, allows the aerobic treatment processes to be completed within a reduced volume and thus allows a portion of the existing tank volume to be converted to an anoxic zone or to incorporate an anaerobic zone for biologically enhanced phosphorus removal. Increased capacity is also possible, because the Media Systems
Additional biomass retained within basin
MEDIA RAS
2,500 mg/L MLSS
Conventional Activated Sludge
RAS
2,500 mg/L MLSS
IFAS
Figure 6.1 IFAS process versus conventional activated sludge process.
Hybrid Processes
clarifiers are not subjected to the increased mixed-liquor concentration, although there would be hydraulic limits to an increase in capacity. Thus, IFAS offers a practical and often cost-effective approach to upgrade treatment facilities that are located on tight sites and must improve their level of performance. The IFAS process can be applied to almost any type of process flow schematic and reactor configuration. It has been used primarily in the aerobic zones of treatment processes to enhance biochemical oxygen demand (BOD) removal and nitrification. Thus, many of the applications of IFAS have been in the modified Ludzack-Ettinger (MLE or enhanced MLE)-type process. However, depending on the type of media, IFAS also has been applied to anoxic zones to enhance denitrification. Although IFAS can be incorporated to a process for biologically enhanced phosphorus removal (BEPR), media has not been used in an anaerobic zone, because the mechanism of BEPR relies on exposing the biomass to alternating anaerobic and aerobic conditions. However, there has been some research into the use of IFAS in an anaerobic zone to improve the production of the volatile fatty acids needed for BEPR. As mentioned, the type of reactor configuration that can be adapted to IFAS also is flexible. The IFAS process has been used in both complete-mix and plug-flow reactors, although, as will be discussed later, each type of reactor has its own special design considerations, depending on the typed of media used. The IFAS process also has been applied to lagoons and sequencing batch reactors. The IFAS process sometimes is confused with the moving-bed biofilm reactor (MBBR) process, because both processes use the same type of media. However, the MBBR does not incorporate a return activated sludge (RAS) and thus is a pure fixedfilm process. The IFAS process does have a return sludge and maintains mixed-liquor concentrations that are typical of a conventional activated sludge process. Some of the general advantages and disadvantages of IFAS systems that have been documented through previous experience are summarized below. Some of these are media-dependent and will be explained in greater detail later in this chapter.
1.1 Advantages Advantages of IFAS systems include the following: • Ability to phase-in additional capacity or improve performance by adding more media;
261
262
Biofilm Reactors
• Additional biomass for treatment without increasing the solids loading on final clarifiers; • Higher-rate treatment processes possible, thus allowing greater treatment in a smaller space; • Improved settling characteristics (reduced SVIs); • Reduced sludge production; • Simultaneous nitrification and denitrification; and • Improved recovery from process upsets.
1.2 Disadvantages Disadvantages of IFAS systems include the following: • Potential for odor (when tank dewatered), • Additional operating appurtenances, • Need to relocate media, and • Increased headloss associated with media-retention screens.
2.0 Media Types A variety of media systems have been used, and several types have become standard in the industry. In general, the media types may be differentiated as either fixed or free-floating. Fixed media includes media that is woven into a rope or a hexagonal pattern. The fixed media is mounted on frames and remains stationary in the activated sludge basin. Free-floating media may consist of either cuboids of a sponge material or small plastic carrier elements resembling wagon wheels. Basically any type of fixed-film media can be used in an IFAS system. Early in the development of IFAS, sheets of plastic material and trickling filter media were used experimentally. The attributes that affect the potential usefulness of a particular media include its specific surface area (SSA), vulnerability to clogging, ability to control growth on the media, durability, installation requirements, and operator requirements. Several rotating biological contactor (RBC) plants have been upgraded by the addition of a return sludge flow and diffused aeration, which essentially converts the pure fixed-film RBC process to a hybrid of activated sludge and fixed film. Some
Hybrid Processes
plants have used a submerged biological contactor (SBC), which essentially is an RBC with a greater level of submergence. The RBCs typically operate with the media approximately 40% submerged, while SBCs operate at 70 to 90% submergence. The greater degree of submergence reduces the load imposed by the media and biomass on the shaft, which allows greater surface area to be installed on an SBC. The SBC can be driven by submerged aeration, which provides process air. If driven mechanically, the SBC can operate anoxically for denitrification.
2.1 Fixed-Media Systems Although various types of fixed-film media systems have been experimented with, the one type of fixed media that is prevalent in the market is the rope-type media. Rope media, also referred to as looped-cord or strand media, takes the form of a woven rope with protruding loops that provide a surface for the growth of biomass. The media is manufactured of a polyvinyl, polyester, or polyethylene material. Manufacturers of the media systems include Ringlace Products Inc. (Troutdale, Oregon), Brentwood Industries (Reading, Pennsylvania), Entex Technologies (Chapel Hill, North Carolina), Biomatrix Technologies (Providence, Rhode Island), and Eimco Water Technologies— Cleartec system (Montreal, Quebec, Canada). The Ringlace and Biomatrix products are strung vertically in racks that contain many strands of media. Brentwood Industries manufactures a mesh-type media, called AccuWeb. Entex Technologies manufactures BioWeb, which is woven into a mesh-like pattern. These media systems are hung in a frame. The Eimco Cleartec product also is manufactured in sheets, but with a different design intent than most of the other products. The Cleartec system is intended to promote the growth of a thin and more uniform biofilm, to promote nitrification. The Webitat system by Entex (BioWeb media) also is intended to promote the growth of nitrifiers. The other products support thicker biofilms, which can be aerobic, anoxic, and anaerobic, depending on the location of the media. Rope-type media systems have been used for carbonaceous BOD (CBOD) removal, nitrification, and, to a limited extent, denitrification. Generally, the media is installed in the aerobic zone of an activated sludge process, to enhance BOD removal and for nitrification. Some degree of denitrification occurs in the aerobic zone as the nitrates produced in the nitrifying biomass on the surface diffuse down into the anoxic layers of the biofilm. The amount of denitrification varies, depending on the bulk dissolved oxygen concentration and the amount of carbon available, but the amount of denitrification can be substantial.
263
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Biofilm Reactors
2.2 Free-Floating Media Systems 2.2.1 Plastic There have been several applications of a fixed plastic media, such as trickling filter media or caged packing material, in an activated sludge process. The problem typically experienced with this type of media is the potential for clogging within the media, as a result of rags and other coarse solids. Although there are some common and successful applications of fixed plastic media in industrial applications, the trend with municipal applications has been to use some form of free-floating plastic media, also referred to as plastic-carrier elements. There are several manufacturers of this type of media. Although they each have their own specific dimensions, they each loosely resemble a “wagon wheel.” The biomass grows on the surface, but is abraded from the outside surface of the media, leaving the active biomass on the inside of the wheel. Manufacturers include Veolia/ Kruger (AnoxKaldnes, Inc., Lund, Sweden), Headworks BIO (Houston, Texas), Infilco Degremont (Richmond, Virginia), Siemens/USFilter (Warrendale, Pennsylvania), and Entex (Chapel Hill, North Carolina). With a specific density slightly less than that of water, the media is distributed throughout the mixed liquor with the aid of aeration in the aerobic sections or with submersible-type mixers (e.g., slow speed and banana-blade type) in the anoxic sections. These media types work with both coarse-bubble aeration and fine-bubble aeration. The high-density polyethylene plastic, under normal operation, will not degrade or require regular replacement, especially if the media is manufactured with the inclusion of UV inhibitors. This type of media requires a retaining screen that typically is 1 mm smaller than the media’s smallest significant dimension. Screen types can be of the flat-panel type or flanged cylindrical type. The screen design should incorporate sufficient screen area to minimize headloss. An air knife may be required on some installations to continuously scour the screen. Upfront fine screening (1 or 2 mm smaller than the retaining screen) is recommended to keep inorganic debris from accumulating in the basin.
2.2.2 Sponge Sponge media is a free-floating media comprised of small cuboids that typically are made of a reticulated polyethylene foam. It has a specific gravity close to that of water and, with good mixing, the sponges are distributed throughout the mixed
Hybrid Processes
liquor. Screens are required at the upstream and downstream ends of the activated sludge basin to retain the sponge cuboids within the media zone. Because the media will tend to migrate with the flow toward the downstream screen, a pump is required to continuously return media to the upstream end of the media zone. Typically, an airlift pump is used. The amount of biomass growing on the sponge media will vary and must be controlled to prevent it from sinking. Typically, the airlift return pump system includes an impingement plate to control growth. As the media exits the discharge of the pipe, it is forced to impact a plate, which forces the mixed liquor and some of the attached biomass from the cuboid. Alternatively, a sponge cleaning pump can be used. Periodic operation of a submersible pump will scour biomass from the media and prevent excessive buildup of growth. An air knife is provided at the screen to continuously clean the screen and reduce the possibility of clogging. The main manufacturer of sponge-media systems is Mixing & Mass Transfer Technologies (State College, Pennsylvania).
3.0 History of Process The use of solid surfaces and fixed-film attachment media in aerated reactors is an old concept that originally was applied primarily for new designs, but, in recent years, has been used increasingly as a means of increasing the treatment capacity and nutrient-removal capabilities of existing activated sludge facilities. Some of the early work with submerged fixed-film media began with the development of the contact aeration process, which was used extensively in the United States in the 1930s and 1940s. The plant treatment trains consisted of primary clarification followed by two equal-sized aeration basins with intermediate and final clarification. Sludge return was not practiced. Asbestos sheets were suspended vertically in the aeration basins, such that they extended from 10 cm below the liquid surface to just above the horizontal pipe-grid aerators. The asbestos sheets were spaced 3.8 cm apart throughout the length of the aeration tanks and permitted a serpentine pathway for floating material. The aerators provided oxygen transfer and mixing between the asbestos sheets, and treatment was affected by the attached fixed-film growth on the sheets. Typical combined hydraulic retention times (HRTs) for the two aeration basins ranged from 1.7 to 3.0 hours. The National Research Council Report (1946) developed design criteria for the contact aeration process simultaneously with the development of design criteria for activated sludge and trickling filter processes. The contact aeration process eventually failed, because it had the cost of diffused aeration but produced a
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worse-quality effluent than activated sludge. Also, aeration costs were high, because the solid asbestos sheets restricted the horizontal diffusion of oxygen. The process essentially ceased to be used in the 1960s. Another major effort to use fixed-film media in aeration tanks originated in Japan during the 1960s. Kato and Sekikawa (1967) developed a process that they called fixed activated sludge (FAS) and applied it primarily to the treatment of industrial wastewaters, particularly those that were alkaline and deficient in nitrogen, which is typical of soft drink production and bottling. This process vertically suspended an open plastic matrix in an aerated reactor and was typically used without return sludge. More than 60 such installations were in place in Japan during the 1960s. Kato and Sekikawa (1967) demonstrated that the process used filamentous microorganisms, required less nitrogen than activated sludge, and could successfully treat wastewaters with influent pH values of 9 to 10. Randall et al. (1972) performed laboratory-scale experiments with the FAS system and demonstrated that it could successfully treat wastewaters that were nitrogen-deficient with pH values as low as 2.85. Although filamentous microorganisms were the dominant attached forms in the FAS system, the sloughed material settled readily. Although the process appeared to have a lot of promise, it was never used widely, except in the soft drink industry in Japan. Rope-type media originally was developed in Japan for the purpose of achieving greater levels of CBOD removal within the confines of an existing activated sludge basin. The intent was to build treatment units that were simple to operate and small in size for installation to remote areas. The product was then applied in Germany to upgrade treatment facilities for nitrification and subsequently in North America for both carbon removal and nitrification. Pilot-scale testing of the rope media began in the United States in the early 1990s (Sen et al., 1993), followed by full-scale installations. The Annapolis Water Reclamation Facility (Anne Arundel County, Maryland) (WRF) was one of the first full-scale applications of the media in North America, where it was used to implement biological nutrient removal (BNR). Free-floating plastic media originally was developed by Kaldnes-Miljoteknologi (Odegaard et al., 1994) in Norway for application in what they labeled as the KMT Moving Bed Biofilm Reactor. The media consisted of small cylindrical biofilm carrier elements made of polyethylene or polypropylene. The MBBR process using this media was applied successfully to the relatively dilute and cold wastewaters at several plants in Norway for carbon and nitrogen removal. Several treatment plants were built with considerable flexibility in their configuration and operation to
Hybrid Processes
provide an opportunity to research the capabilities of the media. An example is the Lillehammer plant in Norway. Since then, the media has been applied to a number of municipal and industrial treatment systems around the world. The media eventually was used in IFAS systems in the United States for carbon removal and nitrification. More recently, it has been used successfully in post-anoxic zones to enhance denitrification. Sponge media originally was developed in Europe in the late 1970s, and two basic systems emerged in the market. One system, known as the Captor process, was developed by Simon-Hartley from work done at the University of Manchester in the United Kingdom (Austin and Walker; Cooper, 1989). The concept was commercialized by Simon-Hartley in conjunction with several universities, Severn Trent (Coventry, United Kingdom), and the Water Research Center of Swindon, United Kingdom. Captor sponges are approximately 25 mm × 25 mm × 12 mm (1 in. × 1 in. × 0.5 in.). A second sponge-type system was developed by the Linde AG Corporation (Munich, Germany) in the mid-1970s and introduced commercially in Europe in the early 1980s. It then was introduced in both Europe and North America as the LINPOR System by the Lotepro Corporation (Mount Kisco, New York). The first full-scale applications were for CBOD removal. There are two alternatives available for nitrogen removal—LINPOR-CN (simultaneous carbon and nitrogen removal) and LINPOR-N (nitrification only). They differ primarily with respect to carbon loadings. Also, the LINPOR-N system is sometimes operated as a tertiary system for nitrification and is not followed by clarification. The LINPOR sponges are approximately cubical, with a side dimension of 10 to 12 mm (0.4 to 0.5 in.).
4.0 Application of INTEGRATED FIXED-FILM ACTIVATED SLUDGE Systems 4.1 Fixed Media 4.1.1 General Requirements The rope media is installed on frames that are installed in a fixed location in the reactors. The rope is looped around the top and bottom of the supporting bars, which are installed in the aluminum or stainless-steel frames. Some types of fixed media are installed as sheets of media with the top and bottom of the media, which resembles netting woven in a hexagonal pattern, supported by the bars in the frames. The frames are manufactured of aluminum and typically are supported off the floor,
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although, in some cases, they have been supported from above by structural members attached to the tank walls. The top of the frame typically is approximately 0.3 m (1 ft) below the water surface, and the bottom of the frames is approximately 0.6 m (2 ft) above the bottom of the tank, to provide clearance for the diffusers located below the frames. The frames should be anchored to the tank, because the forward velocity of the mixed liquor in the tank and agitation caused by aeration may be sufficient to move the frames. This may be especially true as a result of the added buoyancy caused by gas from denitrification trapped in the biofilm. The density of the rope media on the frames is a design condition. Once the required length of the media has been established, the density of the media, in terms of spacing of the rope on the individual supporting bars and the spacing of the bars in the frames, can be varied, within limits. As a starting point for design, the mean cell residence time (MCRT) in the suspended-growth mixed liquor should be greater than the minimum required to support the growth of nitrifiers at the minimum design temperature. If it is less than the minimum, nitrification eventually may be lost over the period of several sludge ages, as the biomass on the media sloughs off and is lost from the system. This has occurred in some rope-type media systems. However, there are other documented cases where the MCRT in the suspended growth was not adequate to support nitrification, but the plant was able to nitrify because of the media. This may be the result of recent evidence that a greater percentage of the biofilm on the fixed-film media is comprised of nitrifiers. Also, biomass containing nitrifiers tends to slough off of the media and seed the mixed liquor with nitrifying organisms. In addition, the process should be designed such that approximately one-half of the ammonia to be nitrified is oxidized by the suspended growth, and approximately one-half is oxidized by the fixed-film biomass. These design criteria essentially accomplish the same goal as stated previously, regarding the minimum MCRT in the suspended growth.
4.1.2 Growth on Media As discussed in the History of Process section, the use of rope-type media originated with a need to develop a compact and easy-to-operate process for secondary levels of wastewater treatment in rural areas. Although rope-type media systems certainly remove BOD, such as in a pre-anoxic zone or any aerobic reactor, currently they tend to be applied mostly to enhance nitrification. Because fixed-media systems are stationary, the type of growth on the media will depend on the nature of the substrate it is exposed to and therefore the location of the media in the tank. If a fixed media is
Hybrid Processes
located at the head of an activated sludge basin in a plug-flow reactor, the relatively high concentration of carbon will promote the growth of a thick biomass populated primarily by heterotrophic bacteria. This type of growth may not be desirable, if it is excessive and causes adjacent strands of media to stick together. Much of the biomass can become anaerobic and actually displace useful tank volume, which would be used more efficiently by the suspended-growth biomass. If the media is located at the end of a plug-flow reactor, it is possible that there would not be adequate substrate to promote the growth of a useful population of biomass. As discussed in the Process Design section, the location of the media and concentration of the various substrates in the bulk liquid are important design considerations. When the media is located properly, generally approximately in the middle twothirds of a plug-flow activated sludge basin, the growth of a nitrifying biomass will be encouraged. As discussed in the section on kinetics (Section 4.1.3), the ratio of carbon to nitrogen is an important consideration in determining the degree to which the growth of autotrophic nitrifying organisms will be favored over heterotrophic biomass. With rope-type media, a significant amount of denitrification can occur within the aerobic zone. The degree to which the ammonia oxidized to nitrate also is denitrified depends on the thickness of the biomass, carbon available, and dissolved oxygen concentration.
4.1.3 Kinetics As with any fixed-film system, the kinetics are dependent on the substrate concentration. If the goal of the media system is to enhance nitrification, then it is important to locate the media in a region of the activated sludge basin where it will be exposed to a concentration of ammonia that maximizes the nitrification rate. This can vary, depending on the media and the wastewater characteristics, but previous research has indicated that the optimum range is approximately 2 to 8 mg/L. If the media is exposed to higher concentrations of ammonia, there probably is also a high concentration of chemical oxygen demand (COD), which would tend to encourage the growth of heterotrophic organisms over the autotrophs that are responsible for nitrification. Because kinetics also are temperature-dependent, the optimum location also may vary seasonally. However, experience has shown that relocation of the media seasonally is not necessary once the media has been located approximately in the area of the tank where effective treatment is achieved. Adjustment in the MCRT can be made, within limits, to influence the amount of carbon and ammonia that is directed at the media zone of the process.
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4.1.4 Worm Growth Several rope-type media IFAS systems have experienced a bloom of a type of redworm population that feeds on the biomass on the media, thereby reducing or eliminating treatment. Limited operating experience under these conditions provides some insight to the cause for their bloom and means of treatment; however, the exact conditions that promote their growth are not well-documented or understood. The worms are obligate aerobes. Thus, high levels of dissolved oxygen favor their growth. They may be encouraged to grow when operating conditions change from a higher load and low dissolved oxygen condition to a lower load and higher dissolved oxygen condition. The high load condition could establish a fairly thick growth of biomass, which, once the load was reduced and the dissolved oxygen increased, would provide an ideal environment for the worms. Control of the worms consists of creating an anoxic condition in the area of the worm bloom, by turning the air off for several hours. In addition, the RAS should be chlorinated. This will kill the worms, but the treatment must be repeated to coincide with the egg cycle (approximately 2 weeks). A certain, limited population of redworms typically will exist in a well-functioning IFAS process. Their presence may reduce the effective sludge yield. They will be visible as small colonies, the size of a dime or quarter, located randomly on the media strands.
4.1.5 Media Breakage The breakage of media strands is not a problem during normal operation. However, excessively heavy growth or over-aeration can contribute to conditions that would stress the media strands. Media breakage is more likely to occur when the media rack is removed from a tank by an overhead crane for inspection or for maintenance of the aeration diffusers located below the media. Prevention requires the use of heavierduty media frames with adequate trusses on the frames that will not flex to the point that would cause the media to break.
4.1.6 Dissolved Oxygen Level Activated sludge processes typically operate with a dissolved oxygen level of 2.0 mg/L or slightly greater. Fixed-film processes benefit from higher levels of dissolved oxygen, because all substrates must penetrate the biofilm, and higher levels of substrate in the liquid provide a greater driving force into the biofilm; the greater the penetration of the substrate into the biofilm, the more biomass that is effectively doing the work. However, rope-type IFAS systems operate well at the same level of dissolved oxygen, or perhaps
Hybrid Processes
slightly higher levels, as pure suspended-growth systems. A significant amount of simultaneous nitrification/denitrification can occur at this dissolved oxygen level. Higher levels will decrease the amount of denitrification that occurs in the aerobic zone. The level of aeration not only controls the dissolved oxygen level, but also the degree of mixing and scouring of biomass. An adequate level of aeration is required to maintain mixing and drive the mixed liquor, which contains the dissolved oxygen, into and throughout the media section. A certain amount of biomass scour is desirable, but excessive turbulence will remove too much biomass and decrease the treatment efficiency. Both coarse- and fine-bubble diffusers have been used in rope-type media systems.
4.1.7 Mixing The media racks submerged within a suspended-growth reactor present an increased resistance to normal flow patterns, and the mixed liquor will tend to flow around the racks rather than penetrate completely through the media. The aeration pattern, whether using coarse- or fine-bubble diffusers, should be arranged to establish a cross-flow pattern that will circulate the mixed liquor into the media. Deflection baffles can be added to the edge of the aeration basin to force mixed liquor flowing along the sides to be diverted into the media. Also, if the media is located just downstream of a baffle wall, openings in the baffle wall can be arranged to distribute flow across the complete cross-section of the tank.
4.1.8 Access to Diffusers Access to the grid of aeration diffusers located below a media section is required periodically for maintenance. Access can be gained by either providing a means of relocating the media racks within the tank, such as by sliding them on rails, or by lifting the racks with an overhead crane. Because access to the diffusers is not required frequently, it may be more cost-effective and simpler to remove the media with a crane. In this case, lifting cables attached to the frames and tethered to the sides of the basin must be provided. Note that the media can deteriorate if subjected to UV radiation, so the media must be protected from exposure to sunlight when a tank is dewatered.
4.1.9 Odor Odors are not an issue unless a tank is dewatered and a media section with biomass is exposed to the air. Under these conditions, the odors can be quite severe. It is suggested that, before taking the tank offline, waste flow to the tank is stopped and aeration is continued for a period of time to “burn off” the active biomass. This will not eliminate any potential source of odors, but it may reduce the potential.
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4.2 Free-Floating Media—Sponge Media 4.2.1 General Requirements Sponge-type IFAS systems are installed within aerobic zones, with screens both upstream and downstream of the zone to retain the media. Ideally, the media zone is configured as a complete-mix reactor, such that mixing within the zone is maximized. The objective is to prevent migration of the sponge media toward the downstream screen. In a plug-flow reactor, the media zones can be established as a series of complete-mix zones within the reactor. However, the reality is that the combined forward flow of the influent, return sludge, and internal nitrate recycle will cause the media to migrate toward the screen. Thus, some type of media-return pump is required to recycle media from the area of the screen to the head of the media zone. Typically, an airlift pump is used for this purpose. The percent fill of media within the media cell typically is limited to approximately 35%. This constraint is imposed by the potential of greater quantities of media to plug the downstream screen and the amount of aeration energy required to suspend the media and keep it well-mixed. There is a considerable exchange of the suspended-growth mixed liquor with the interior volume of the sponge media. This interchange is promoted by the pore structure of the sponge media and the regular forcing-out of the liquid within the sponge by the impingement plate. Thus, from a practical design standpoint, the sponge media can be considered as a cuboid with a high concentration of mixed liquor. The equivalent biomass within the sponge cuboid has been estimated between 15 000 and 20 000 mg/L.
4.2.2 Screen Clogging The sponge media naturally will tend to migrate downstream toward the screen. A possible exception to this is where the aeration basin has been divided into media zones that resemble small, complete-mix basins, and sufficient mixing energy has been provided to ensure that the sponges are completely mixed throughout the zone, regardless of the flow through the basin. The operator should observe the in-basin screens on a daily basis. If there is an indication that the liquid level is building up on the upflow side of the screens, the operator should check the air knives, to make sure they are operating and injecting sufficient air, and the airlift recirculation pumps, to ensure they are operating satisfactorily. Also, the forward flow through the basin should be checked to make sure that it is not too high. Flows may be excessive, as a result of high wet-weather flows, poor distribution of
Hybrid Processes
flows through multiple basins, or excessive nitrate recycle rates. The velocity induced by the forward flow may present a constraint to the percent internal recycle rate for the nitrate recycle pumps and is a factor that should be considered during design. If these are within the normal range, the water level in the tank should be lowered 0.9 to 1.5 m (3 to 5 ft) for a few hours to check the screens for biofilm buildup. This may occur if the air knife was not operating properly, the roll pattern within the tank was too gentle, or if the sponges were starved of organic carbon for an excessive length of time. The latter, in combination with low mixing, leads to a buildup of long-tailed stalked ciliates, which appear as a “skin” on the sponges. When this skin sloughs off, the material can temporarily plug the screen. The screen then will have to be cleaned with a hose on the downstream surface. The intensity of the roll pattern and the airflow through the air knife then would need to be increased in the future. Consideration should be given to a bypass as a safety feature in the event of a clogged screen. The bypass could be a side overflow weir into an adjacent tank or a pipe in the basin that allows flow from an upstream zone to a downstream zone, thus bypassing the media zone.
4.3 Control of Biomass Growth Excessive growth on the sponge media can cause the sponges to sink to the bottom of the tank. This typically is managed by the airlift return pump, which includes an “impingement plate” at the discharge of the pipe. As the sponges impact the plate, mixed liquor and some attached solids are forced out of the sponge. Another approach is to use a submersible pump that operates periodically on a timer. The pump is located in the media zone, and, as sponges pass through the volute of the pump, turbulence removes some of the biomass. The operation of the pump should be adjusted to adequately control growth, but not to strip too much biomass from the media. Typically, the cleaning pump is run in one basin at a time, which increases the MLSS of the system by 100 to 300 mg/L. A schedule for operating the cleaning pump should be developed based on the capacity of the pump and the size of the basin.
4.3.1 Loss of Sponges The sponge media will abrade continuously against the walls of the aeration basin. Over a period of time, the cuboids become rounded, and surface area is lost. Typically, it is necessary to replace a few percent of the media, at least for the first few years. After that, as the sponges become rounded, there is no need to continue adding replacement sponges.
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4.3.2 Taking Tank Out-of-Service If the fill volume of a sponge-media-type IFAS system is relatively low, which would be in the range 20 to 25%, the sponges may be left in the basin as it is dewatered. After dewatering, the sponges could be pushed to the side of the tank to gain access to the diffusers. At 40%, it is recommended that the media be pumped from the cell to another tank that does not contain media or distributed to the basins remaining in operation. Distribution of the sponges and return of the media to the dewatered tank are managed by accounting for the sponges on a mass-balance basis. If the approximate concentration of the sponge media is known, then the number of sponges can be accounted for by metering the liters (gallons) pumped. The media should not be exposed to sunlight for extended periods of time, because UV radiation will cause the media to disintegrate.
4.3.3 Loss of Solids Should there be a washout of MLSS in a storm event, the operator can take advantage of the biofilm in the sponge system to replenish the MLSS. By squeezing biomass out of the media, the operator may be able to recover the process.
4.3.4 Air Distribution System Air is required not only for the process, but also to operate the air knife and airlift pump. If these are fed from a common manifold, a situation could develop where the system becomes unstable hydraulically. If the air supply to the air knife and airlift pumps is taken from a common manifold, which supplies air to the diffusers in all the reactors in the system, then even a small increase in head will cause the air to redistribute in the manifold to equalize pressure. Thus, at a time when the system needs more air, the system will deliver less. Therefore, it is recommended that consideration be given to an independent blower system—preferably a positive displacement blower—to supply air to the airlift pumps and air knives.
4.3.5 Plastic Media 4.3.5.1 General Requirements Plastic media typically is installed within the aerated zones of a reactor for CBOD removal and nitrification. However, plastic media also has been demonstrated to successfully enhance denitrification in anoxic zones—especially post-anoxic zones with supplemental carbon addition. Screens are required to retain the media both upstream and downstream of the media zone. The screens typically are made of stainless steel and are configured as
Hybrid Processes
either flat plates of mesh or wedge wire or as cylinders oriented horizontally and projecting into the direction of the flow. Plastic-media IFAS systems may be designed with a percent fill of up to 65%. The allowable fill fraction of plastic media is higher than allowed with other types of free-floating media, because plastic media has a lower potential to bind a downstream screen and because it is maintained more easily in a complete-mix condition within the reactor. Typically, a medium-bubble aeration system consisting of small pores drilled into a stainless-steel pipe is used instead of fine-bubble membrane diffusers. This type of air diffuser provides sufficient mixing energy, while also providing a high level of oxygen transfer. The oxygen-transfer efficiency may be comparable with that of finebubble membrane diffusers in suspended-growth systems, because the bubbles in the IFAS system tend to cling to the media, which increases the retention time of the bubble within the liquid. However, there is also evidence that the bubbles tend to coalesce on the media into larger bubbles with less efficient oxygen transfer. Thus, the gain in oxygen-transfer efficiency resulting from media may not be significant. The recommendations regarding the use of plastic media in MBBR systems (Chapter 5) are also relevant to this discussion.
4.3.5.2 Biomass Growth The operator should monitor the accumulation of growth inside the hollow plastic media. If growth appears to be excessive, the roll pattern may be too gentle, or the organic loading may be excessive. To rectify this, the organic loading may have to be reduced, or additional media may have to be installed. Also, the roll pattern could be increased, or media with a larger diameter could be installed in the section of the tank that is subjected to the highest organic load. The biomass growth can range from 5 to 30 mg total suspended solids (TSS/m2), depending on the loading, temperature, and operating conditions. Media located in upstream section of the system typically will have a thicker biofilm, with a greater portion of the biofilm attributed to heterotrophs. Media in downstream sections will have a thinner biofilm, with a higher proportion of autotrophs. Because the media is free-floating and completely mixed within the tank, the biofilm will adjust itself to the conditions in the bulk liquid. During periods of extended high loads, the biofilm will get thicker, while, during periods of extended low loads, the biofilm will get thinner. Typically, the whole aeration basin will contain the media, even if divided into several cells. Therefore, there is no need to relocate media to other sections of the basin in response to changing load patterns.
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4.3.5.3 Media Mixing The aeration pattern, whether using coarse- or fine-bubble diffusers, should be arranged to establish single or multiple roll patterns, to ensure that the media is mixed fully in the basin and that no dead zones are apparent.
4.3.5.4 Screens Retention screens are required at the effluent side of each of the media zones in the aeration basin to prevent media loss or migration from the aeration basin. The screen design has evolved from flat screens mounted on the face of the cell wall to cylindrical screens that are mounted horizontally facing into the flow. This type of design has proven to be effective in preventing clogging. The media itself tends to scrub the screens clean. Flat screens should be equipped with air knives to encourage media scouring and to avoid excessive headloss resulting from accumulations on the screen surface. Preventative maintenance schedules should include screen inspections and cleaning. As with sponge media, preliminary treatment should including screening, preferably with fine screens that have opening of 6 mm or less.
4.3.5.5 Foaming Foaming may occur during system startup. However, foaming may occur as a result of operational issues, especially when excessive airflow rates are used. Dissolved oxygen concentrations should be monitored, and airflow rates can be decreased if they are greater than 3.0 mg/L. Because aeration also serves to keep the media mixed, operators must remain mindful that substantial reductions in airflow may affect media movement. Suspended-growth parameters, such as MLSS concentration, food-to-microorganism ratio, and low influent flowrates also can contribute to foaming issues. These parameters should be monitored closely to ensure an optimized treatment scheme.
4.3.5.6 Media Replacement The breakage of media typically is not a problem, unless the media type is operated in open tanks and does not contain UV inhibitors. In this case, breakage may occur, and the media pieces may pass through the retaining screen and float to the surface in the downstream clarifier.
4.3.5.7 Taking Tank Out-of-Service If the fill volume of the plastic-media-type IFAS system is relatively low (20 to 25%), then the media may be left in the basin as it is dewatered. After dewatering, the media could be pushed to the side to gain access to the diffusers. At fill volumes greater than
Hybrid Processes
25%, it is recommended that the media be pumped from the media cell to another tank that does not contain media or distributed to the basins remaining in operation. Distribution of the media to other tanks during servicing and return of the media to the dewatered tank is managed by using a Secchi disk. The disk can be sunk into the basin during aeration. Aeration then is stopped, and the media depth is determined by lifting the Secchi disk. Levels are determined when the disk meets resistance from the media. Following servicing, media depths within the basins can be returned to initial levels based on the level readings.
4.3.5.8 Worm Growth Worm growth is not a problem for free-floating media systems, provided that there is constant and uniform mixing of the media to prevent any deadzones that potentially may allow the organisms to thrive. The organisms are not transported easily between carriers, because of their separation in space, and they cannot tolerate the turbulence of the biomass carrier elements.
4.3.5.9 Startup During initial installation, the media has a tendency to float on the water surface until being thoroughly “wetted” out, although greater airflow into the aeration basin will promote media mixing. Airflow rates may be reduced once the biofilm is established. Depending on the wastewater temperature, the media will show signs of performance in 2 to 4 weeks from startup. Foaming can occur during the initial weeks of startup. During this time, an antifoam chemical may be used to mitigate foaming issues, or, if equipped, the plant may use its foam abatement system. Excessive foaming typically ceases once the microbiology is established.
5.0 Process Design 5.1 Introduction The principle of an IFAS system is to enhance the removal of BOD and nitrogen over and above the removal that could have been achieved using the MLSS alone. The combination of removals in the biofilm and mixed liquor increases the complexity of the design of IFAS systems relative to activated sludge and MBBRs. Membrane bioreactors that have very low MLSS concentrations in municipal strength applications have little removal in the mixed liquor and can be modeled as pure biofilm systems. Integrated fixed-film activated sludge systems, however, must consider the interaction between the biofilm and the mixed liquor.
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5.2 Parameters Influencing Organics Removal in the Biofilm of Integrated Fixed-Film Activated Sludge Systems 5.2.1 Biofilm Flux Rates Biofilm flux rate is the rate of transport of a particular substrate or electron acceptor across the liquid–biofilm interface. The typical units of flux are in grams of substrate removed per square meter of biofilm surface per day (g/m2·d or kg/1000 m2·d). The biofilm surface is the surface area of biofilm that develops on the media; it is not the surface area of bare media. The flux rates can be in terms of various parameters, such as COD, dissolved oxygen, ammonium-nitrogen, oxidized nitrogen, volatile suspended solids (VSS), and inert solids. It is important to understand the factors that control the biofilm flux rates. The flux rate for a substrate, such as COD, increases with the concentration of substrate (COD) and the concentration of electron acceptors (e.g., dissolved oxygen and oxidized nitrogen forms) in the mixed liquor (bulk liquid outside the biofilm). It also is influenced by the biofilm thickness, biomass density in the biofilm, and thickness of the stagnant liquid layer. A lower thickness of the stagnant liquid layer, as may be observed at higher intensities of mixing or a more open media structure, can increase the concentrations in the biofilm and increase the flux rate. Additionally, a higher biofilm thickness increases the number of layers in which the substrate is used and can increase the flux rate, if electron acceptors are available in the deeper layers. A higher biomass density (mixed-liquor volatile suspended solids [MLVSS] of the biofilm, in milligrams per liter) also increases the COD flux rate.
5.2.2 Removals in Biofilm per Unit of Tank Volume In addition to the biofilm flux rate, the removal that takes place per unit of tank volume with media depends on the biofilm specific surface area (square meters of biofilm surface per cubic meters of tank volume) and the media fill fraction (mf). The removal per unit of tank volume (kg/m3·d) is as follows: Biofilm flux rate (kg/1000 m 2·d) × biofilm specific surface area at 100% fill (m2/m3) × media fill fraction (mf) The biofilm specific surface area at 100% fill depends on several factors. These include the following: • Type of media used. The specific surface area at a certain fill-volume fraction increases from fixed-bed media to sponge moving-bed media to certain types of plastic cylinders.
Hybrid Processes
• The thickness of the biofilm on the media increases with increase in organic substrate (soluble biodegradable COD) levels in the bulk liquid. For most media, the surface area decreases when the biofilm thickness increases above a certain optimal level (Figure 6.2). In those situations, the thickness can be decreased by more vigorous mixing to induce a higher rate of biofilm shear. • The extent of bare media surface that is covered by the biofilm. This can change with the external COD concentration. Cylinder dimensions: Outer diameter = 10 mm Length = 10 mm Thickness of annular ring = 1 mm Thickness of cross vane = 1 mm Hypothetical carrier particle has one vane per cylinder No fins/ridges/protruberance
(a)
10 mm
Thin biofilm
Thick biofilm
Vane
Biofilm surface area is the inner surface of the brown biofilm
(b)
Fin/ridge/ protruberance
Cylinder dimensions: Outer diameter = 10 mm Length = 10 mm Thickness of annular ring = 1 mm Thickness of cross vane = 1 mm Hypothetical carrier particle has one vane per cylinder Has several fins/ridges/ protruberances
10 mm
Thin biofilm
Vane
Thick biofilm
Biofilm surface area is the inner surface of the brown biofilm and outer surface of biofilm between fins/ridges/protruberances
Figure 6.2 Effect of biofilm thickness on locations of growth and surface area of biofilm (from Sen et al., 2007): (a) magnitude of difference in biofilm surface area for “thin” and “thick” biofilms on the same carrier particle of moving-bed or IFAS system, and (b) magnitude of difference in biofilm surface area for “thin” and “thick” biofilms on the same carrier particle of moving-bed or IFAS system.
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A summary of various media characteristics, including dimensions and the specific surface area, is provided in Table 6.1. It should be noted that the physical area potentially available for biofilm growth in a sponge media cuboid is potentially very high. It is reported that the LinPor media (Mixing & Mass Transfer Technologies) has a specific surface area of 2800 m2/m3 because of the open-pore structure, which makes 97% of the internal volume available. However, the specific surface area indicated in Table 6.1 is based on bench-scale kinetic studies, where the biofilm nitrification rate (kg/1000 m2·d) was calculated based on just the outside surface area of the sponge cuboid. If the surface area provided by the internal structure of the sponge cuboid was included, then the specific nitrification rate calculated would be lower. The media fill fraction (mf) is the fraction of activated sludge tank volume filled with media. At 100% fill, the value of mf equals 1.0. For fixed-bed media, the media fill fraction is the fraction of the tank floor and height covered by media frames. For Table 6.1 Summary of media characteristics. Media
Diameter (mm)
SSA* (m2/m3)
AnoxKaldnes (Kruger, Cary, North Carolina) K1
9
500
K2
15
350
K3
25
500
Biofilm Chip M
48
1200
Biofilm Chip P
45
900
Headworks (Houston, Texas) 22
402–680
Infilco Degremont (Richmond, Virginia) ActiveCell 515
22
515
ActiveCell 450
22
450
19
575
AgWise ABC4
14
600
AgWise ABC5
12
660
Entex Technologies (Chapel Hill, North Carolina) Bioportz Siemens/USFilter (Warrendale, Pennsylvania)
* SSA is for 100% fill volume.
Hybrid Processes
Table 6.2 Biofilm SSA of various types of media. Applied* SSA (m2/m3)
Minimum aerobic HRT Recommended (hours) at MLSS (mg/L) 12°C
Type of system
Media
Media fill volume percentage
Activated sludge
None
0
0
3000
7
IFAS—fixed bed
Bioweb, Accuweb
70 to 80
50 to 100
3000
5
IFAS—movingbed–sponge
Linpor, Captor
20 to 40
100 to 150
2500
4
IFAS—movingbed–plastic
K1 (Kaldnes), Entex, Hydroxyl
30 to 60
150 to 300
2500
4
MBBR—K1
K1 Kaldnes
40 to 67
200 to 335
<1000
3
*Based on SSA of media within range of recommended percent fill volumes (applied SSA = % fill × SSA at 100% fill).
moving-bed media, it is the fraction of the empty tank volume filled with the moving bed of biofilm carrier particles. For moving-bed media, the upper limit on fill fractions is determined by the sum total of recycle and influent flows through the portion of the reactor (cell) with media. The presence of recycle increases the forward flux, and this must be countered by the aerobic mixing and media recirculation. Because of the presence of nitrate and RAS recycles in secondary systems, the upper limits on media fill fractions for secondary systems typically are lower than those applied in tertiary systems (no RAS) and post-anoxic cells of enhanced nutrient removal (ENR) systems (no nitrate recycle). The specific surface area for various types of plastic media is provided in Table 6.1. Some general guidance on the media fill fraction and biofilm specific surface area for different media types is provided in Table 6.2.
5.3 Parameters Influencing Removals in the Mixed-Liquor Suspended Solids As with activated sludge systems, the removal of COD per unit of tank volume in the IFAS system increases with the increase in MLVSS and with the substrate
281
282
Biofilm Reactors
concentration (soluble biodegradable COD) and dissolved oxygen (or oxidized nitrogen) concentrations in the mixed liquor. For the same tank volume, an increase in the mixed-liquor MCRT (MLSS MCRT) increases the MLVSS. This increases the COD uptake rate (kg/m3·d) in the MLVSS, which reduces the quantity of COD that has to be removed by the biofilm. This decreases the biofilm surface area required.
5.4 Interaction Between the Mixed-Liquor Suspended Solids and the Biofilm In designing an IFAS system, it is important to understand the interaction between the biomass and interdependency of removals in the MLVSS and the biofilm. An increase in removal in the mixed liquor, as observed at a higher MLSS MCRT, decreases the COD concentrations at any location along the aerobic zone. This decreases the COD uptake rates (flux rates) for the biofilm at that location. For the same MLSS MCRT, an increase in temperature has a similar effect. Reducing the MLSS MCRT or the temperature has the opposite effect; the rate and the fraction of COD removal in the biofilm increases. Additionally, in designing an IFAS plant, one can take advantage of the available clarifier capacity to operate the system with a higher MLVSS and MLSS MCRT and reduce the amount of biofilm required. This may allow one to change the type of media required to a lower specific surface area. It also may allow use of a smaller volume compared with an activated sludge system and an MBBR. Other than the available clarifier capacity, one should evaluate the risk of entrapment of foam at the screens within the basin (Figures 6.3 and 6.4) when Nitrate Recycle
Anoxic With or Without Media
Aerobic Cell 1 With media
Aerobic Cell 2 With media
Aerobic Cell 3 With media
RAS Typical Layout of Plastic Carrier IFAS System Media does not move longitudinally within the aerobic zone, biofilm thickness and fraction of nitrifiers are different in different aerobic cells. Biofilm thickness is higher and fraction nitrifiers is lower in the first aerobic cell as compared to cells downstream.
Figure 6.3 Typical layout of plastic-carrier-media IFAS systems.
Hybrid Processes
Nitrate Recycle
Anoxic
Aerobic Zone with sponge media Air lift for Media Recycle Sponge impingement or squeezing device
U shaped Screen
RAS Typical Layout of Sponge Media IFAS System Media moves within the aerobic zone, therefore, the biofilm thickness and characteristic are the same at all location.;
Figure 6.4 Typical layout of sponge-media IFAS system.
Nitrate Recycle
Anoxic With or Without Media RAS Typical Layout of Cord or Web Fixed Bed IFAS System Media is held in place with frames. There are no in-basin screens. Biofilm thickness is controlled by maintaining a vigorous roll pattern through the media. The roll pattern is generated by arranging the diffuser grid or by installing supplemental diffusers below the media frames. For nitrification, the system is operated at a higher MLSS MCRT as compared to Sponge and Plastic Carrier Systems. It can be often operated at a higher MLSS as compared to Sponge and Plastic Carrier Systems.
Figure 6.5 Typical layout of cord or web fixed-bed-media IFAS systems.
the MLVSS is increased. If the plant has a propensity to accumulate foam (either seasonally or because of changes in the influent characteristics), one may want to limit the MLVSS and the MLSS MCRT and use a higher quantity of biofilm media instead. This is not a concern in fixed-bed systems, because they do not use in-basin screens (Figure 6.5).
283
284
Biofilm Reactors
5.5 Interaction Between Heterotrophs and Nitrifiers Understanding the interaction between heterotrophs and nitrifiers is one of the most interesting challenges of IFAS systems. As with activated sludge systems, factors that increase the soluble biodegradable COD levels in the bulk liquid (from 5 and 20 mg/L) result in a substantial increase in the growth rates of heterotrophs relative to nitrifiers. This decreases the fraction of nitrifiers in the biofilm and negatively affects the nitrification rates in the biofilm. In an activated sludge system, a reduction in MLSS MCRT close to the washout MCRT of nitrifiers decreases the fraction of nitrifiers in the MLVSS, until nitrification is lost. Installation of biofilm support media, which can carry biomass at longer MCRTs, will grow some nitrifiers in the biofilm. The fraction of nitrifiers in the biofilm will depend on the soluble biodegradable concentration and dissolved oxygen levels in the bulk liquid outside the biofilm. When nitrifiers are present in the biofilm of the IFAS system, the sloughing of the biofilm seeds the mixed liquor. At MLSS MCRTs close to the washout MCRT, this sloughing and seeding can result in a significant increase in the fraction of nitrifiers in the MLSS compared with an activated sludge system operating at the same MLSS MCRT. Therefore, the presence of biofilm media helps increase the nitrification rates per unit of tank volume, not only by an increase in rates in the biofilm, but also by increasing the rates in the mixed liquor. As long as nitrification in the biofilm is not compromised, it can reduce the time required to recover full nitrification compared with activated sludge systems (e.g., after winter wet-weather flows that reduce the temperature).
5.6 Design Tools/Procedures Two types of design procedures for IFAS systems have evolved over time. These are • Empirical methods based on the experience of each manufacturer or process supplier and • Process-kinetics-based models, which rely on different levels of integration of activated sludge and biofilm kinetics. The empirical methods are custom methods that are developed by each manufacturer. They may be based partially on activated sludge and biofilm kinetics. Typically, the contribution of the biofilm and its kinetics are simplified by the manufacturer or process supplier, based on observations made at their facilities.
Hybrid Processes
Models based on empirical methods should be applied with caution. Further, their application should be limited to the type of media and system for which they were developed and tested. For example, if nitrification was observed in a tertiary system with low influent COD, the rates and models should not be applied to a secondary system. The process kinetics methods apply different approaches to modeling a biofilm. These include the following:
(1) Biofilm rate models; (2) Semi-empirical biofilm models (which include additional Monod kinetics to incorporate the effect of bulk substrate concentrations on biofilm thickness and fraction of nitrifiers in the biofilm); (3) Single-layer biofilm with mass transport (with diffusion coefficients), in addition to Monod kinetics; and (4) Multi-layer biofilm models.
All four models can offer the requisite level of accuracy, if they are constructed in a manner that allows the user to calibrate the model and adapt it for the type of media being applied.
5.6.1 Empirical Methods 5.6.1.1 Equivalent-Sludge-Age Approach This method is applied by some suppliers of sponge-media systems. The quantity of biomass in the sponge media is added to the quantity in the mixed liquor. The concentration of biomass within an individual sponge cuboid is available from the manufacturer. Mixing & Mass Transfer Technologies has reported an equivalent biomass concentration of 18 000 mg/L within the volume of a sponge. Thus, if the percent fill volume of sponge media is 20% in an activated sludge basin with a mixed-liquor concentration of 3000 mg/L, then the combined biomass concentration is 3000 mg/L (1 to 0.20) + 18 000 mg/L (0.20). An equivalent sludge age (or MCRT) is calculated by dividing the total biomass (in the biofilm and mixed liquor) by the daily wasting rate. In this instance, the process supplier has developed specific curves to equate the equivalent MCRT to plant performance. It is recommended that the method is applied only for the sponge-media system in question. Sponge media have a different means of removal of excess biomass (sponge squeezing instead of sloughing). For the surface area of biofilm provided, they carry a relatively higher quantity of biomass in the biofilm relative to plastic
285
286
Biofilm Reactors
moving-bed media. This is because of the higher thickness and biofilm density. If the media is allowed to move along the length of the aerobic zone, the fraction of nitrifiers in the biofilm and the thickness are the same along the length of the aerobic zone. The method is reliable, as long as it is applied to the same conditions (and performance) as the system(s) from which the data were collected. It is akin to doing a scalable pilot study and applying the results to the design.
5.6.1.2 Quantity (Length or Web Surface Area) of Media Approach This approach may be used by some suppliers of fixed-bed media. The media is sold, either in terms of the length of media (e.g., Ringlace [Ringlace Products, Inc.]) or the surface of the web (e.g., Accuweb, Brentwood Industries, Reading, Pennsylvania, and Bioweb, Entex Technologies, Chapel Hill, North Carolina). The suppliers have measured rates on the plants and pilot studies they have operated. This information then is applied to other plants. Once again, one should be careful in the application of this method. It should be limited to the process conditions under which the data were collected. One should ask for evidence of rates measured or review the data from operating facilities to understand whether the improvements are because of nitrification rates in the biofilm or because of the creation of plug-flow kinetics and better SVI, often associated with the installation of fixed-bed media. The SVI can improve because of plug-flow kinetics and the nature of biomass sloughing off the biofilm.
5.6.2 Rates Based on Pilot Studies This approach uses a pilot study to represent the proposed full-scale application and offers the following possible advantages: • Opportunity to become familiar with the technology, • Potential to negotiate a performance guarantee based on results, • Provide data for in-depth calibration of process model, and • Refine the details of design. It is a good approach, as long as the pilot study data are representative of the conditions in the full-scale system. One should note that there are challenges in replicating the fixed-bed media system in pilot studies that do not accommodate the type of diffusers and the depth that would be used in full-scale systems (pilot studies can perform poorly).
Hybrid Processes
5.6.3 Biofilm Rate Model The following approach may be used for preliminary sizing. It is based on rates observed in full-scale and pilot studies and in calibrated process models.
5.6.3.1 Define Range of Flux Rates For a primary effluent (reactor influent) with a COD/total Kjeldahl nitrogen (TKN) ratio of 7.5:1 to 15:1, one may use the following rates at a mixed-liquor temperature of 15ºC and an aerobic zone dissolved oxygen of 3 mg/L: • Aerobic COD uptake rates = 0.5 to 5 kg/1000 m2·d and • Nitrification rates = 0.05 to 0.5 kg/1000 m2·d. To adjust rates for other temperatures, apply an Arrhenius temperature adjustment coefficient of 5% for every degree Celsius change in temperature. The actual value of flux rate for the media location and application is discussed in the following section.
5.6.3.2 Quantify Removal at Different Mixed-Liquor Suspended Solids Mean Cell Residence Times It is recommended that one use the following fraction of removals in the mixed liquor and biofilm to quantify the biofilm surface area required in IFAS systems. These are based on analyses performed with process kinetics models. These removals are at 15°C. • At a 2-day MLSS MCRT, 50% COD and 80% nitrification is on the biofilm. The rest is in the MLVSS. • At a 4-day MLSS MCRT, 25% COD and 50% nitrification is on the biofilm. • At an 8-day MLSS MCRT, 20% nitrification is on the biofilm. The MLSS MCRTs mentioned above may be increased by 3% for every degree Celsius increase in temperature. The quantity of ammonia to be nitrified on the media then can be estimated.
5.6.3.3 Select Flux Rates Based on Location Along Aerobic Zone The third step is to break up the aerobic zone into thirds. • Apply rates of 75, 50, and 25% of the maximum rate for COD uptake for first, second, and third thirds of the aerobic zone. • Apply rates of 25, 50, and 75% of the maximum rates for ammonium-nitrogen uptake for the first, second, and third thirds of the aerobic zone, when targeting an effluent ammonium-nitrogen concentration of 1 mg/L.
287
288
Biofilm Reactors
Nitrification rates in the middle and last third of the aerobic zone should be decreased in proportion to the ammonium-nitrogen concentration in the mixed liquor when the effluent quality required is less than 1 mg/L. For example, at an effluent ammonium-nitrogen concentration of 0.5 mg/L, the rate applied in the last third should be 0.75 × 0.5 / 1.0 = 0.375 of the rate above 1 mg/L.
5.6.3.4 Calculate the Quantity of Media Required Knowing the quantity of ammonia to be nitrified and the adjust rate, the quantity of media required then can be calculated. The surface area of biofilm required is the higher of the two required for COD removal and nitrification. The rates based on BOD generally would be one-half of the rates based on COD.
5.6.3.5 Additional Analysis to Finalize a Design It is recommended to use one of the software models discussed in Chapter 11. For IFAS systems, it also is recommended that one evaluate results with more than one type of software, to increase the confidence level in the predictions.
5.6.3.6 Application of Kinetics-Based Approach with Integrated Fixed-Film Activated Sludge Design Software A discussion of design equations used in the kinetics-based approach is presented in Chapter 11. The chapter also provides the results of application and evaluation of the software against actual long-term data from an IFAS system (Broomfield Wastewater Treatment Plant [WWTP], Colorado). The three types of software evaluated performed satisfactorily for COD removal, nitrification, and nitrogen removal. There were some differences in their ability to predict biofilm parameters, such as thickness and sludge production.
6.0 Case Studies 6.1 Annapolis Water Reclamation Facility, Anne Arundel County, Maryland The Annapolis WRF originally was a 38 000-m3/d (10-mgd) activated sludge plant designed to meet secondary treatment limits. The plant discharges to the Chesapeake Bay. It has undergone several upgrades to meet increasingly more stringent effluent limits requiring nutrient removal. The discharge permit requirements have been driven by the Chesapeake Bay initiatives (www.chesapeakebay.net), which first established goals for BNR levels of treatment or a discharge level of 8 mg/L total nitrogen and 2 mg/L total phosphorus. Then, nutrient limits based on a waste-load allocation were established, which would require effluent limits for total nitrogen of 6 mg/L and total
Hybrid Processes
phosphorus of 1.5 mg/L. Finally, ENR levels of treatment were required, which set the discharge limit at 3 mg/L total nitrogen and 0.3 mg/L total phosphorus. During this period of time, the Annapolis WRF was required to expand to 50 000 m3/d (13 mgd). The initial upgrade to BNR levels of treatment used an IFAS process, which is a system that incorporated fixed-film media within the conventional activated sludge process. The IFAS was selected, after considerable pilot study, because it proved to effectively increase the performance of the plant without having to build additional activated sludge basins. Thus, it was a cost-effective option for the County. When the Annapolis WRF had to expand to 50 000 m3/d and upgrade to ENR levels of treatment, there was no longer any alternative, except to build additional reactor volume, either with or without IFAS. At this point, the decision was made to abandon the media system and implement a more conventional process using preand post-anoxic zones in a strictly suspended-growth system. An IFAS system would have reduced the cost of the upgrade and expansion by requiring less reactor volume to be built, but the cost savings was not enough to overcome the desire to operate a plant with a more conventional HRT. Also, there was some concern regarding the future cost of replacing the existing media.
6.1.1 Original Wastewater Treatment Plant The original treatment plant, built in 1972, consisted of two parallel two-pass activated sludge basins in a horseshoe pattern followed by two circular clarifiers. Unit processes included headworks with screening and grit removal, primary clarification, activated sludge, chlorination, and dechlorination. The design capacity was 38 000 m3/d (10 mgd), with an effluent permit requiring secondary treatment. Basic design criteria are presented in Table 6.3.
6.1.2 Pilot Study (1993 to 1996) Chesapeake Bay initiatives (www.chesapeakebay.net) required the implementation of nutrient removal to an effluent level of 8 mg/L total nitrogen and 2 mg/L total phosphorus. Conventional methods requiring the construction of additional tanks were expensive, which led the County to investigate IFAS systems in a pilot study. Pilot studies were conducted in conjunction with Virginia Tech (Blacksburg, Virginia) at the plant and in the laboratory. The configuration of the full-scale pilot study is shown in Figure 6.6. For the first phase of the pilot study (1993), the Annapolis WRF was modified into two separate plants by separating the return sludge systems. The first plant (plant 1) was operated for nitrogen removal using suspended growth only. The second plant (plant 2) was converted to an IFAS system.
289
290
Biofilm Reactors
Table 6.3 Design criteria, original Annapolis WRF. Activated sludge tanks (2 tanks) Volume (m3)
4730 each
Length (m)
120.0 each pass
Width (m)
11.0
Depth (m)
3.6
Pre-anoxic fraction
0.27
Aerobic fraction
0.73
Final clarifiers (2 clarifiers) Diameter (m)
32
Side water depth (m)
3.0
RAS
Anoxic Primary Effluent
Nitrate Recycle
To Secondary Clarifiers
Aerobic
Anoxic
Aerobic
Optional Anoxic Zone w/ Step Feed Ringlace Racks
Figure 6.6 Pilot study configuration. The fixed-film media used in this IFAS installation was a rope-type media manufactured by Ringlace Products. Ringlace is a flexible polyvinyl chloride-type of material that is woven together into strands. The strands have loops that protrude out of it. The approximate diameter of the loop is 5 mm. The media is held by brackets installed on the top and bottom of modules. The modules are supported on aluminum frames. A total of 30 000 m of Ringlace were installed within 473 m3 of tank volume located in the aerobic zone of the second pass. Two lightweight test racks, each with 15 m (50 ft) of media, also were installed on cables at the front end of the Ringlace zone. The test racks were pulled out of the basin at regular intervals to examine the growth on the Ringlace. Modifications were made to implement the MLE process. Anoxic zones were constructed by portioning the front end of the aeration tank into four cells with mixers. Two of the anoxic cells also could be switched to aerobic cells. Nitrate recycle pumps were installed at the end of the aerobic zone to recycle nitrates from the end of the anoxic zone.
Hybrid Processes
The modifications to plant 1, which served as a control during certain phases of the study, were similar. However, fixed-film media was not installed in plant 1. The plant was operated for nitrogen removal by shutting down the air in the first 25 to 40% of the aeration basin and recycling nitrates from the end of the aerobic zone. From January through June 1993, the pilot plant was operated in the MLE mode. In July, the configuration was converted to step-feed with nitrate recycle and operated in that manner through December 1993. For the next phase of the pilot study (1994 to 1996), additional Ringlace media was installed such that approximately 63 000 m were added in the aerobic section of the first pass and 46 000 m were added in the aerobic section of the second pass. This increased the total amount of media in the IFAS system to 109 000 m.
6.1.3 Full-Scale Upgrade for Biological Nutrient Removal (1997 to 2000) The plant was upgraded based on the pilot study results. A schematic of the plant and the BNR process implemented in the activated sludge tanks using an IFAS system with rope media are shown in Figure 6.7. The 2 two-pass reactors were converted to 4 single-pass reactors operating in parallel, as shown in Figure 6.8. The sizes of the anoxic and aeration zones are shown in Table 6.4. The anoxic volume was 20% of the total volume. A total of 61 000 m (200 000 ft) of Ringlace media was installed in each reactor. Two new final clarifiers were added (39.6-m [130-ft] diameter). Construction began in 1997 and was completed in 2000. During most of this period, one of the reactors or one-half of the activated sludge basin volume was out-of-service for construction. Thus, treatment was maintained at an HRT that varied from 3.5 to 8 hours. The HRT, MCRT, and effluent ammonia levels are shown in Table 6.5.
6.1.3.1 Pilot Study Results of the full-scale pilot study have been reported in previous articles (Copithorn et al., 1995; Sen et al., 2000; Sen et al., 1994a; Sen et al., 1994b). Conclusions from these studies can be summarized as follows: • Nitrification and denitrification per unit volume of tank can be enhanced with biomass growing on rope-type (Ringlace) media installed in the aerobic zone of a single-sludge nitrogen removal system. • Nitrification rates were a function of the ammonium concentrations in individual segments (cells) with rope-type media, with optimum rates of 1.7 kg/d/1000 linear m of media observed at ammonium concentrations between 3 and 5 mg/L, and lower rates observed above and below this concentration.
291
292
Biofilm Reactors
Primary clarifiers City influent Bypass
Wet well
Biological treatment (1)
Screens
Grit removal
County influent Bypass
Shellfish protection holding pond Outfall
Secondary clarifiers Post - SO2 aeration
Junction box
Chlorine contact tank (2)
Figure 6.7 Schematic of BNR upgrade to IFAS. Primary Effluent To Secondary Clarifiers
Anoxic
Nitrate Recycle Aerobic
Switch Zone
Ringlace Racks
Figure 6.8 Process configuration for BNR upgrade. These rates are applicable in a temperature range 14 to 19°C, when the MCRT of suspended growth equals or exceeds the minimum required for retaining nitrifiers in a plug-flow system. • Under optimum conditions, the nitrification on the media was 60% of the total nitrification in the segment. These results were for a density of 0.8 linear m of Ringlace media per cubic meter of activated sludge tank volume.
Hybrid Processes
Table 6.4 Design criteria, BNR upgrade. Zone
Volume (%)
Volume (mil. gal [m3])
Anoxic 1.1
10
0.061 (230)
Anoxic 1.2
10
0.061 (230)
Anoxic 1.3 (switch)
10
0.061 (230)
Aerobic 1
26.5
0.162 (613)
Anoxic 2 (switch)
13.8
0.084 (318)
Aerobic 2
26.5
0.162 (613)
3.2
0.019 (72)
Deoxygenation Total
0.61 (2309)
Table 6.5 Operating MLSS MCRTs, HRTs, media, and average performance for IFAS.
Date
Phase
Season
MLSS MCRT (days)
1997 to 2000
BNR construction
Summer to fall
4.6
4.0
1092 1223
1.1
1997 to 2000
BNR construction
Winter to spring
5.2
4.6
1092 1223
1.5
2001 to 2005
BNR operations ENR construction
Summer to fall
6.2
8.1
244
0.35
2001 to 2005
BNR operations ENR construction
Winter to spring
6.5
8.8
244
0.9
HRT (hours)
Media in service (1000 m)
Effluent NH4N (mg/L)
• As much as 20% of the ammonium nitrified within the section with the media was denitrified while operating at dissolved oxygen levels in excess of 3.0 mg/L.
6.1.3.2 During Construction (1997 to 2000) For nearly the entire period of construction, as the 2 two-pass reactors were being converted to 4 single-pass reactors, the treatment plant met its effluent permit requirements, with only half of the activated sludge basins online. The basins kept in service initially were the pilot study basins with the Ringlace media. Table 6.6 shows the results of the IFAS process operated at an MCRT and HRT as low as 3.2 days and 3.2 hours, respectively. During construction, there were some occasional process upsets caused by taking units offline and interim startup issues, but, generally, BNR was
293
Table 6.6 Performance data. Influent flow Year 1995
294 1996
Date
BOD mgd ML/d mg/L
Raw TSS mg/L
TKN mg/L
Final effluent NH3-N BOD mg/L mg/L
TSS mg/L
NH3-N mg/L
TKN mg/L
TN mg/L Comment
Jan-95
7.5
28 388 232
218
31.7
19.0
13
12
0.4
2.8
10.2
Feb-95
7.6
28 766 208
197
26.9
16.8
13
11
0.4
2.7
9.8
Mar-95
7.6
28 766 240
236
31.5
18.0
11
9
0.2
1.7
6.8
Apr-95
7.5
28 388 207
219
31.4
17.2
15
9
0.4
2.4
11.0
May-95
7.8
29 523 209
227
31.6
15.6
6
6
0.3
2.4
9.9
Jun-95
7.5
28 388 196
233
25.7
14.0
8
6
0.2
1.3
7.3
Jul-95
7.8
29 523 178
224
22.0
12.0
5
4
0.2
1.0
8.1
Aug-95
8.2
31 037 213
303
20.6
16.0
5
8
1.7
2.9
8.3
Sep-95
8.2
31 037 228
261
19.4
16.3
11
9
6.4
7.1
12.3
Oct-96
8.4
31 794 228
230
21.1
16.3
5
9
0.2
1.6
7.2
Nov-95
8.3
31 416 213
220
20.8
15.0
7
8
0.2
1.5
7.6
Dec-95
7.7
29 145 188
166
23.8
16.6
10
8
0.6
2.7
8.5
Jan-96
5.8
21 877 172
173
36.8
15.7
16
15
0.8
3.3
8.8
Feb-96
5.5
20 893 190
198
22.0
16.4
14
13
1.3
4.6
11.0
Mar-96
5.5
20 780 194
228
23.6
17.2
18
12
11.9
15.0
18.5
Sludge treatment return flow
Apr-96
5.7
21 650 173
206
17.2
13.8
24
12
7.2
7.9
12.3
Marina waste dumped
May-96
5.8
21 764 174
226
12.8
9.1
16
11
3.9
4.9
12.8
1996
1997
295 1998
Jun-96
5.7
21 461 193
225
13.9
8.2
6
6
0.2
0.9
6.7
Jul-96
5.7
21 650 225
237
14.3
8.9
4
4
0.2
0.6
7.2
Aug-96
5.7
21 726 232
219
23.3
15.9
4
3
0.2
2.3
11.1
Sep-96
6.2
23 353 303
149
18.5
14.0
9
4
0.9
2.1
8.7
Oct-96
5.9
22 483 232
209
19.0
13.4
7
5
0.6
2.2
9.1
Nov-96
5.5
20 818 169
153
17.1
13.7
7
6
0.4
1.6
8.6
Dec-96
5.9
22 407 209
223
15.9
12.2
8
8
0.2
1.3
7.9
Jan-97
8.4
31 908 204
194
25.4
13.8
10
9
0.5
2.1
9.3
Feb-97
8.7
32 854 221
152
18.7
15.0
16
11
1.2
3.1
11.7
Mar-97
8.8
33 346 144
132
16.3
12.0
14
10
0.6
2.0
12.6
Apr-97
8.8
33 157 156
174
13.6
10.8
15
10
1.6
2.9
13.7
May-97
8.5
32 097 170
175
13.3
10.8
14
9
0.7
2.4
10.4
Jun-97
8.5
32 097 201
187
12.6
11.0
11
9
0.3
1.2
9.3
Jul-97
8.1
30 583 193
185
11.7
9.1
7
5
0.2
1.2
8.0
Aug-97
8.0
30 431 296
270
19.3
15.6
8
6
0.3
1.2
9.1
Sep-97
7.7
29 069 181
144
15.1
13.5
6
4
0.5
1.1
10.0
Oct-97
7.5
28 539 191
172
21.4
17.6
9
5
0.9
2.2
14.8
Nov-97
8.1
30 469 155
126
19.7
17.3
13
12
0.6
1.9
9.6
Dec-97
7.2
27 176 175
150
19.5
17.7
11
7
1.6
2.8
11.5
Jan-98
8.2
30 886 219
185
19.0
17.4
14
9
1.4
1.9
11.5
Feb-98
9.3
35 238 192
155
18.1
16.7
14
11
1.3
1.4
10.6
Apr-98
8.3
31 567 131
126
18.4
15.4
10
9
0.5
1.8
9.8
May-98
8.5
32 324 165
120
15.8
13.1
8
6
0.5
1.7
10.1
Table 6.6 Continued Influent flow Year
1998
296 1999
Raw
Final effluent
Date
BOD mgd ML/d mg/L
TSS mg/L
TKN mg/L
NH3-N BOD mg/L mg/L
TSS mg/L
NH3-N mg/L
TKN mg/L
TN mg/L Comment
Mar-98
9.4
35 428 170
153
15.4
12.9
12
12
0.8
2.1
7.4
Jun-98
8.0
30 242 192
172
16.6
14.1
8
7
0.6
1.5
9.6
Jul-98
7.8
29 409 159
183
17.8
13.8
6
5
0.4
1.1
8.2
Aug-98
7.6
28 842 170
168
14.5
13.4
4
4
0.2
1.0
8.4
Sep-98
7.4
27 933 212
186
18.2
12.2
9
8
1.1
3.2
9.9
Oct-98
7.0
26 457 237
184
20.1
14.4
15
8
3.1
4.5
11.6
Nov-98
6.7
25 170 252
206
23.6
17.4
16
5
4.8
6.6
12.8
Dec-98
6.3
23 846 251
256
29.6
20.8
11
4
7.6
10.1
15.2
Jan-99
6.5
24 474 291
221
25.3
19.0
8
7
0.4
1.8
10.1
Feb-99
6.5
24 659 215
209
23.4
16.6
8
7
0.4
1.8
10.1
Mar-99
6.7
25 262 228
215
19.0
15.0
6
4
0.8
2.8
10.6
Apr-99
6.6
25 145 201
194
18.0
13.0
6
3
1.0
2.4
10.1
May-99
6.6
24 993 193
193
8
3
0.7
1.1
8.4
Jun-99
6.8
25 700 232
210
6
3
0.6
1.5
9.6
18.8
10.8
Jul-99
6.7
25 360 224
183
13.9
10.1
21
2
6.5
7.8
13.8
Aug-99
6.5
24 567 212
217
20.3
11.0
14
3
2.9
6.6
14.1
Sep-99
8.2
31 093 286
252
18.5
8.7
10
3
0.8
1.7
9.2
Oct-99
7.7
29 258 225
199
17.7
10.7
4
2
0.2
1.1
8.0
Nov-99
7.3
27 682 247
196
24.4
15.2
3
2
0.2
0.9
8.0
Dec-99
7.1
26 926 258
230
22.8
14.6
4
2
0.2
1.2
6.8
Blower shut down to retain solids
2000
297 2001
Jan-00
6.8
25 749 254
230
32.9
14.6
7
3
0.6
2.8
8.9
Feb-00
7.1
26 874 289
220
30.7
19.4
7
3
0.2
2.3
8.2
Mar-00
7.2
27 252 231
202
27.5
14
2
6
Apr-00
7.7
29 145 308
263
29.7
16.9
7
3
0.2
1.2
6.6
May-00
7.4
28 009 231
213
29.9
14.8
4
2
1.2
5.5
Jun-00
7.1
26 874 270
170
33.4
16.2
3
1
0.2
1.2
6.4
Jul-00
7.6
28 766 245
219
23.9
18.5
3
2
0.2
1.0
5.2
Aug-00
7.6
28 766 225
178
24.3
16
3
2
0
1.1
4.7
Sep-00
7.7
29 145 287
200
27.5
15.8
3
2
0.3
1.5
6.0
Oct-00
6.7
25 360 192
183
26.9
17.0
6
3
7.8
10.7
16.0
Nov-00
6.5
24 603 241
190
28.9
19.9
6
3
0
2.0
8.7
Dec-00
6.6
24 981 299
158
30.3
19.8
5
2
0.3
1.7
7.7
Jan-01
6.3
23 846 198
144
28.3
20.9
8
3
0.4
1.6
6.5
Feb-01
6.4
24 224 223
155
27.3
18.2
9
3
0.4
1.7
6.5
Mar-01
6.8
25 738 233
166
26.9
16.4
5
3
0.2
1.1
5.5
Apr-01
6.9
26 117 261
148
28.6
14.0
4
2
0.2
1.5
6.2
May-01
6.9
26 117 256
184
23.1
13.2
3
2
0.2
1.3
5.5
Jun-01
7.2
27 252 263
193
19.4
10.1
2
1
0.2
1.2
5.4
Jul-01
7.2
27 252 277
180
16.6
10.3
2
1
0.2
1.3
6.2
Aug-01
7.1
26 874 238
173
20.1
10.7
2
1
0.2
1.4
6.6
Sep-01
6.8
25 738 297
194
21.9
13.9
7
2
0.4
2.0
7.2
Operational problem causes upset
Table 6.6 Continued Influent flow
Raw
Final effluent
Year
Date
BOD mgd ML/d mg/L
TSS mg/L
TKN mg/L
NH3-N BOD mg/L mg/L
TSS mg/L
NH3-N mg/L
TKN mg/L
TN mg/L Comment
2001
Oct-01
6.5
24 603 252
168
23.4
12.0
5
3
0.2
2.1
7.2
Nov-01
6.2
23 467 306
190
21.5
13.0
10
2
0.3
2.4
7.8
Dec-01
6.0
22 710 257
181
28.6
11.1
8
2
0.3
3.0
8.0
2002
298 2003
Jan-02
6.0
22 710 337
201
32.6
18.4
6
2
0.2
1.5
6.5
Feb-02
5.9
22 332 240
134
27.5
13.5
5
3
0.2
1.7
6.3
Mar-02
6.0
22 710 345
297
42.6
19.1
7
3
0.2
1.8
6.3
Apr-02
6.3
23 846 316
161
27.0
18.9
5
2
0.2
1.8
6.0
May-02
6.4
24 224 335
173
37.5
20.6
4
2
0.2
1.5
5.8
Jun-02
6.6
24 981 292
223
29.5
21.6
3
2
0.2
0.7
5.8
Jul-02
6.7
25 360 296
209
26.2
17.2
4
2
0.2
1.2
7.2
Aug-02
6.8
25 738 217
158
22.6
14.7
6
2
0.2
1.7
7.5
Sep-02
6.8
25 738 315
188
30.8
19.5
6
3
0.2
2.0
7.2
Oct-02
7.0
26 495 241
181
34.9
18.4
14
5
1.0
3.9
13.1
Nov-02
6.9
26 117 326
224
28.0
22.8
11
4
0.5
2.0
7.8
Dec-02
6.9
26 117 357
195
39.4
28.2
14
4
0.5
2.6
7.5
Jan-03
6.9
26 117 285
140
30.5
26.5
9
4
0.4
1.3
6.6
Feb-03
7.3
27 631 222
111
32.3
22.6
16
7
0.7
4.3
9.6
Mar-03
8.2
31 037 218
172
36.6
20.2
9
3
0.2
1.6
5.2
Apr-03
7.7
29 145 212
178
34.5
23.5
6
2
0.2
1.7
6.3
May-03
8.0
30 280 223
208
33.2
24.6
7
3
2.0
3.7
11.0
Jun-03
8.9
33 687 188
202
26.8
23.3
6
3
0.2
1.3
6.0
Jul-03
7.9
29 902 216
202
25.4
16.7
5
2
0.2
1.1
5.7
1.1 0.2
5.7
1.3 0.2
6.0
3.7
2 5 29 902 216 Jul-03
7.9
202
25.4
16.7
3 6 33 687 188 Jun-03
8.9
202
26.8
23.3
3
7
24.6
33.2
30 280 223
208
6.1.3.3 Post-Construction (2000 to 2003) From the time of the BNR plant startup (2000) through the beginning of construction for the ENR upgrade and expansion (July 2003), the plant maintained consistent nitrification and nitrogen removal; the data are presented in Table 6.6. The BNR upgrade implemented hydraulic wasting by pumping mixed liquor to the gravity thickeners. This provided excellent control over the MCRT.
6.2 Westerly Wastewater Treatment Plant, Westerly, Rhode Island 6.2.1 Introduction
8.0
May-03
Hybrid Processes
maintained, and effluent total nitrogen remained below 8 mg/L, on an annual average. Monthly average influent and effluent data are shown in Table 6.6. Explanations are added for those months where the ammonium-nitrogen was higher than 2 mg/L. These were associated with process upsets from marina waste hauling containing formaldehyde, blower malfunction, or other construction-related issues.
2.0
11.0
After the discharge permit for the Westerly (R. I.) WWTP was revised to include seasonal limits for ammonia and total nitrogen, the facility’s biological treatment process needed a significant overhaul. Before the upgrade, the activated sludge plant was designed to remove CBOD only. To comply with the revised permit requirements, the plant was modified to include an IFAS process, in conjunction with an anoxic reactor capable of removing total nitrogen.
6.2.2 Description of Original Facilities Owned by the Town of Westerly and operated by Aquarion Operating Services Company (Auburn, New Hampshire), the plant has a monthly design average daily flow of 12 500 m3/d (3.3 mgd). The plans for the upgrade were designed by the Beta Group Inc. (Lincoln, Rhode Island) and implemented by the Hart Engineering Corporation (Smithfield, Rhode Island). Before construction began on the process modifications in March 2002, the facility used the following treatment unit processes: influent screening by manual bar rack, grit removal, primary clarification, mechanical aeration, secondary clarification, and disinfection by sodium hypochlorite.
6.2.3 Description of Upgrade The IFAS system installed in each aeration tank is a LINPOR-CN System provided by Mixing & Mass Transfer Technologies. This system uses suspended porous flexible support media—15-mm (0.6-in.) polyurethane cubes—in combination with a freely
299
300
Biofilm Reactors
suspended biomass. The media increases the amount of biomass the aeration basin can support and the resulting solids retention time (SRT). Being able to attain additional SRT in a reduced volume achieves nitrification and allows the use of some of the existing basin as an anoxic denitrification zone. Structural and process equipment modifications were made to the two existing aeration tanks to provide an appropriate configuration for removing nitrogen. The upstream cell of each aeration tank was converted to a baffled anoxic zone with submersible mixers. A single fiberglass-reinforced plastic baffle and support structure were installed in each anoxic zone to minimize flow short-circuiting and improve mixer performance. Two submersible stainless-steel mixers operate in each anoxic zone (see Figure 6.9). The other two cells of each aeration tank were converted to aerobic reactors containing the IFAS carrier media. Tubular, membrane-type, fine-pore, air-diffuser grids were installed in the aerobic zones. The diffusers were designed in a tapered configuration, so that the highest densities of air are released by the diffusers at the location where demand is greatest—that is, the feed end of the reactors (see Figure 6.10). The porous support media is prevented from flowing out of the reactors and into the final clarifiers by a perforated, stainless-steel retention screen installed ahead of the effluent weirs. Extending the full depth of each reactor, the screens have circular perforations measuring 8 mm (0.3 in.) in diameter and spaced 10 mm (0.4 in.) apart on center. A row of coarse-bubble diffusers—known as an air knife—installed at the base of both sides of each screen prevents media from impinging on the screens. The air knife periodically scours the screens to enable mixed liquor to pass freely through the screens and to the secondary clarifiers. Anoxic zone
Aerobic zone Media recycle Process air
Media retention screen Final clarification To disinfection
Primary effluent
Nitrate recycle RAS
WAS
Figure 6.9 Westerly, Rhode Island WWTF IFAS system flow diagram (WAS = waste activated sludge).
Flow
Tank effluent (Typ.)
Media retention screen (Typ.)
Airlift pump discharge hoco (Typ. for 4).
T.O. wall el. 20.33
Airlift pump (Typ. for 4)
FRP baffle (Typ.) Spray nozzle (Typ.)
Flow
Primary effluent (Typ.) RAS (Typ.) Process air from blowers Airlift discharge (Typ.)
301
Flow
Flygt mixer (Typ.)
Tubular fine bubble diffuser (Typ.)
Nitrate recycle pump (Typ. for2) Flow
Figure 6.10 Westerly, Rhode Island, WWTF IFAS system reactor train plan.
302
Biofilm Reactors
Airlift pumps within the aerobic zone perform the following two functions:
(1) Flexing of the foam media to control the solids concentration and (2) Ensuring that the media is distributed uniformly throughout the aerobic reactors.
Air for this service is supplied by the process aeration blowers. Nitrified mixed liquor is pumped from the end of the aerobic zone to the anoxic zone by submersible, mixer-type axial pumps. These pumps are installed downstream of the media-retention screen and do not contact the support media. The flowrate of these pumps is controlled by variable-frequency drives based on a manually programmed speed set-point. Sensors installed in each reactor monitor the water level. If a sensor detects a high level of wastewater in a reactor, it will shut off that reactor’s nitrate-recycle pump. Process air is supplied to the IFAS system by two multistage, centrifugal air blowers; one duty unit; and one backup to provide 100% redundancy. The system operates under average design-load conditions, with one blower in operation at approximately 122 m3/min (4300 scfm) and 49.6 kPa (7.2 psig). The output of the blowers is controlled by variable-frequency electric drives controlled by a programmable logic controller (PLC) tied to the monitored dissolved oxygen level. Airflow to the two aerobic zones is based on the monitored dissolved oxygen level, and the blower speed (output) is controlled based on the resulting air manifold pressure. The dissolved oxygen control set-points in the system are approximately 2.0 mg/L. The blowers are protected locally by an alarm system that detects vibrations and power surges. This alarm system will report problems to the system’s main PLC and send alarms to operators via local telemetry and dialer systems.
6.2.4 Design Criteria Table 6.7 presents the discharge limits, and Table 6.8 provides the design criteria for the Westerly IFAS system upgrade design.
6.2.5 Performance of the Integrated Fixed-Film Activated Sludge System Initially commissioned in July 2003, the IFAS system at Westerly began to demonstrate nitrification, removal of total nitrogen, and stable operation within 4 weeks. Compliance with the revised permit limits was achieved soon after startup. Performance data from the period October 2003 through March 2007 show consistent removal of total nitrogen through the cold-weather months (see Table 6.9).
Table 6.7 Summary of discharge permit limits. Parameter
Discharge limit (average monthly)
BOD
30 mg/L
TSS
30 mg/L
Total nitrogen
15 mg/L, June to October
Ammonia
5.5 mg/L, June to October 30.9 mg/L, November to May
Fecal coliform
Most probable number per 100 mL = 200
Total residual chlorine
65 µg/L
Table 6.8 Design criteria for the Westerly LINPOR-CN system. Design flow Average day
13 000 m3/d (3.3 mgd)
Maximum day
21 000 m3/d (5.5 mgd)
Peak hour
30 000 m3/d (7.8 mgd)
Design loads to secondary treatment CBOD5 average
2110 kg/d (4651 lb/d)
TSS average
1311 kg/d (2890 lb/d) 375 kg/d (826 lb/d) 249 kg/d (550 lb/d)
TKN average NH3 average CBOD5 maximum day
4526 kg/d (9979 lb/d)
TSS maximum day
2781 kg/d (6131 lb/d) 596 kg/d (1314 lb/d) 400 kg/d (883 lb/d)
TKN maximum day NH3 maximum day Anoxic zones Tank length × Width × Sidewater depth
21 m × 5.2 m × 4.3 m (68 ft × 17 ft × 14 ft)
Total anoxic volume of two trains
916 m3 (0.242 mil. gal)
HRT (excluding RAS)
1.8 hours
Linpor reactors Tank length × Width × Sidewater depth
21 m × 10 m × 4.3 m (68 ft × 34 ft × 14 ft)
Total aerobic volume of two trains
1832 m3 (0.484 mil. gal)
SRT
6.6 days
Food-to-microorganism ratio, average
0.25/day
MLSS suspended, average
2500 mg/L
MLSS fixed, average
18 000 mg/L
303
Table 6.9 Summary of operating data.
Parameter
4Q03
1Q04
2Q04
3Q04 4Q04
1Q05 2Q05
3Q05
4Q05 (a)
1Q06 (a)
2Q06
3Q06
4Q06
1Q07 (a)
42 monthly average
Influent Flow,
mgd
2.147
2.147
2.906
1.952
2.530
3.107
3.021
1.892
2.994
3.071
3.098
2.504
2.725
2.977
2.648
BOD5 ,
mg/L
250
207
202
293
202
168
195
292
200
189
251
227
223
171
219
BOD infl,
lb/d
4547
3655
4640
4827
4207
4439
4601
4599
4994
4749
6120
4696
4862
4127
4647
TSS,
mg/L
256
159
154
204
159
130
248
283
182
149
278
193
166
146
193
SS infl ,
lb/d
4409
2834
3341
3318
3312
3338
5691
4446
4553
3672
6315
4073
3705
3376
4027
NH3,
mg/L
13.8
16.0
7.9
6.3
8.2
10.4
TKN,
mg/L
27.0
30.0
15.0
12.8
19.0
20.8
Temp., ave.
°C
18.1
13.5
17.4
22.7
18.6
13.8
17.7
23.7
18.7
14.5
18.3
22.7
18.5
14.2
18.0
Temp., min.
°C
15.1
11.1
14.8
20.0
15.8
11.8
14.8
20.8
15.3
12.3
16.2
20.9
15.5
11.7
15.4
BOD5,
mg/L
4.6
4.6
5.8
5.6
10.3
7.1
4.9
5.9
7.4
7.0
6.4
6.2
5.9
6.5
6.3
TSS,
mg/L
12.0
11.3
10.1
9.4
13.9
10.5
14.2
10.2
18.9
17.1
13.3
12.1
12.8
16.2
13.0
NH3,
mg/L
0.5
0.5
0.7
0.9
0.9
2.3
1.7
0.6
0.8
1.0
0.6
0.4
1.6
2.3
1.0
TKN,
mg/L
2.2
2.1
2.0
2.6
3.0
4.9
3.7
2.5
2.4
2.6
2.4
2.4
3.7
5.1
3.0
TN,
mg/L
4.2
5.4
4.3
5.1
5.6
6.5
5.0
6.4
4.5
4.3
3.8
4.8
5.5
7.1
5.2
Q-RAS,
mgd
1.70
1.66
1.62
1.67
1.67
1.63
1.68
1.67
1.62
1.78
1.60
1.64
1.65
1.73
1.67
Q-IMLR,
mgd
2.12
2.12
2.12
2.12
2.05
2.02
2.02
2.02
1.68
1.01
2.02
2.02
2.02
0.34
1.83
Effluent
Reactors
D.T. oxic, (Q),
hr
5.50
5.41
4.20
5.95
4.66
3.74
4.01
6.15
3.89
3.91
3.95
4.68
4.33
4.01
4.60
D.T. oxic, (Q+R+IMLR),
hr
1.95
1.96
1.76
2.02
1.86
1.72
1.74
2.08
1.85
1.99
1.75
1.89
1.82
2.37
1.91
D.T. anox, (Q)
hr
2.75
2.71
2.10
2.98
2.33
1.87
2.01
3.07
1.94
1.96
1.98
2.34
2.16
2.00
2.30
D.T. anox, (Q+R+IMLR)
hr
0.98
0.98
0.88
1.01
0.93
0.86
0.87
1.04
0.93
1.00
0.87
0.94
0.91
1.18
0.96
MLSS free,
mg/L
2866
2753
2815
3078
2650
2645
3362
3216
2939
2751
2842
2947
2913
2777
2897
MLSS fixed,
mg/L
6452
8171
10633 10877 10368 8733
11360 11514 9344
10210
12977
11897
11789
11671 10428
81
77
72
63
60
65
64
55
46
54
57
57
63
62
63
SVI, WAS ,
mgd
0.023
0.031
0.018
0.028
0.034
0.052
0.032
0.032
0.027
0.025
0.026
0.026
0.022
0.014
0.028
RSS,
mg/L
5684
5835
6633
5548
5232
6208
7310
5737
6918
6618
6947
5850
6407
6463
6242
WAS,
lb/d
1108
1520
964
1284
1460
2713
1944
1540
1587
1377
1490
1268
1149
785
1442
ESS,
lb/d
213
202
243
153
292
273
373
159
472
448
346
251
299
396
294
SRT, free
days
7.6
5.2
7.6
7.1
5.0
2.9
4.8
6.1
4.6
5.0
5.2
6.6
6.6
7.9
5.9
SRT, free + fixed
days
11.8
9.0
14.8
13.4
9.8
5.2
8.8
11.6
8.3
9.6
11.2
13.2
13.3
16.6
11.2
WAS, lb/lb BODrem.*
0.39
0.66
0.34
0.43
0.57
0.99
0.68
0.52
0.51
0.46
0.40
0.44
0.38
0.31
0.50
%SDN
68.0% 30.7% 66.7% 51.6% 59.1% 72.3% 80.4% 15.0% 73.0% 80.6%
77.0%
63.3%
69.5% 77.4% 63.2%
Notes: (a) Quarter s when IMLR was shutdown to test simultaneous denitrification (SDN) performance IMLR = internal mixed liquor recycle D.T. = detention time SDN = simultaneous denitrification
306
Biofilm Reactors
These 3.5-year performance data validate the cold-weather data collected during the LINPOR-CN system pilot-testing in 2002. The IFAS system also has shown excellent performance during wet-weather, highflow events. Because the media-retention screens prevent the fixed biomass from washing out of the system, the high flows had a negligible effect on the process, even in April 2004, when record wet weather caused the influent flowrate to surpass the facility’s design peak-hour flow. Because the fixed biomass of approximately 10 000 mg/L was maintained in the reactors, the system did not require reseeding or excessive pumping of return sludge. During the 42-month data period, as shown in Table 6.9, there were 6 months with monthly average flows exceeding the plant design. Notably, the system removes total nitrogen to levels below 6 mg/L, with nitrate-recycle rates significantly less than the literature values of 2 to 4 times the reactor influent rates. This system operates at a total fixed nitrate-recycle rate of approximately 7570 m3/d (2 mgd), equating to a nitrate-recycle rate of less than one-half of the rate and the pumping-related energy requirements of some conventional nitrogen-removal systems. This reduction in the nitrate-recycle rate results from the simultaneous nitrification and denitrification that occurs in the aerobic reactors. The simultaneous reactions are possible because the interior of the support media can act as small anoxic environments, in which nitrate in the aerobic reactor is denitrified without having to be returned to the anoxic zone by the nitrate-recycle pumps. The IFAS system has shown that effluent ammonia and total nitrogen concentrations below the discharge limits can be met consistently, even in wet-weather and winter conditions.
6.2.6 Operational Issues Between October 2003 and April 2004, some process upsets occurred, in which excessive foaming was experienced. The foaming was attributed to fats, oils, and grease introduced by way of septage added at the facility’s headworks. Although these process upsets did not affect effluent quality, modifications were made at the facility headworks to control the flow of septage and eliminate the foaming. The system has required no more attention than the conventional activated sludge process previously operated at the facility. The operations staff also report that, since the system was placed in service, the facility’s SVI values have improved significantly, resulting in more consistent removal and processing of solids at the facility.
Hybrid Processes
6.2.7 Costs The Westerly IFAS system was part of a $7 million plantwide upgrade. The installed cost of the IFAS system (blowers, piping, media, media return system, screens, instrumentation and control, internal mixed liquor recycle pumps, anoxic mixers, and support services) is estimated at $2 million. Operating costs at the facility are not broken out by unit process; however, the plant operator indicates they are similar to other conventional nitrogen-removal, activated sludge systems.
6.3 Broomfield Wastewater Treatment Plant, Broomfield, Colorado 6.3.1 Introduction The City and County of Broomfield WWTP was expanded in 1988. The secondary treatment processes consisted of a roughing biofilter (trickling filter) followed by activated sludge. The discharge limits were 25 mg/L cBOD and 30 mg/L TSS at a flowrate of 20 000 m3/d (5.4 mgd). As a result of population increases and new regulations with respect to NH3-N, the City again needed to upgrade the facility. Also, the City intended to reuse a large portion of its wastewater as irrigation water. This led to a requirement for lower total phosphorus and total nitrogen levels. Table 6.10 shows the effluent requirements that the City’s WWTP would have to meet for the reuse water storage standards. The City and County of Broomfield evaluated six treatment alternatives based on • Future expansion, • Similar treatment process, Table 6.10 Effluent requirements. Parameter
Effluent requirement (mg/L)
BOD
<10
TSS
<10
NH3-N
Summer
<1.5
Winter
<3.0
Total inorganic nitrogen
<10
Total phosphorus
<1.0
307
308
Biofilm Reactors
• Land usage (as the selected approach needed to leave room at the site for eventual expansion to approximately 61 000 m3/d [16 mgd]), and • Overall cost. An IFAS-type process using free-floating plastic media was selected. The first process train was converted during the summer months of 2002. Wastewater was treated through the process train in September 2002, and media first was added in November 2002. The second process train was converted in the summer of 2003, with final media addition by fall 2003.
6.3.2 Full-Scale Plant Results The full-scale design criteria for the system were based on winter wastewater temperatures and the maximum month primary effluent concentrations. Specific design criteria for the upgraded plant are shown in Table 6.11. A flow diagram of the treatment facility is shown in Figure 6.11. Tables 6.12 and 6.13, respectively, show the monthly averages for the main influent and effluent characteristics of the facility. The overall flow to the facility over the past 3 years has been fairly consistent, ranging from approximately 15 000 to 19 000 m3/d (4 to 5 mgd). It should be noted that only the influent ammonia is provided. Organic nitrogen is not shown, because the facility does not perform TKN analysis. Thus, the overall nitrification load is actually higher. As can be seen from the data, the plant has operated consistently and reliably. The plant has maintained low effluent ammonia levels, while operating at an aerobic SRT of 3 to 4 days, which is typical of a plant designed only for cBOD removal. In addition, the plant has achieved low effluent ammonia levels, with a mixed-liquor temperature of as low as 14°C. This performance is illustrated in Figure 6.12. The stability of the IFAS process, with respect to the aerobic SRT, was demonstrated during the Broomfield WWTP IFAS project.13. Reliable nitrification was maintained at low temperatures and at suspended-growth SRTs that would typically cause the nitrifiers to wash out of the system. During the cold-weather periods, the attached-growth SRT tended to increase, thereby compensating for the reduced nitrification capacity in the suspended-growth biomass.
6.4 Colony Wastewater Treatment Plant, Colony, Texas 6.4.1 Introduction and Background The City of The Colony, Texas, is a growing suburban community located north of the City of Dallas, adjacent to Lake Lewisville. The existing 13 000-m3/d (3.4-mgd)
Table 6.11 Specific design criteria for the upgraded plant. Flow
Average month Maximum month
25 000 m3/d (6.7 mgd) Summer Winter
30 000 m3/d (8.0 mgd) 25 000 m3/d (6.7 mgd)
TSS
Maximum month
97.5 mg/L (6504 lb/d)
BOD
Maximum month
145.8 mg/L (9725 lb/d)
Soluble BOD
90.0 mg/L (6005 lb/d)
NH3-N
Maximum month
37.2 mg/L (2480 lb/d)
TKN
Maximum month
40.8 mg/L (2724 lb/d)
NO3-N
Maximum month
5.2 mg/L (350 lb/d)
MLSS
3500 mg/L
SRT
4.7 days (suspendedgrowth)
Temperature
13 to 25ºC Mixed Liquor Recycle
Primary Effluent
To Secondary Clarifiers
FEQ Return
Mixed Liquor Recycle RAS from Clarifiers
Flow Junction/ Splitter Box
Anaerobic and Anoxic Basins (mixed liquor only)
IFAS Aeration Basins (media and mixed liquor)
Figure 6.11 Full-scale flow diagram (FEQ = flow equalization). 309
Table 6.12 Influent data. Raw influent characteristics
Month, year
Influent flow (mgd)
310
TSS mg/L
Total BOD5 mg/L
NH3-N mg/L
NO3-N mg/L
NO2-N mg/L
Alkalinity Total PO4-P [as CaCO3) Temp. P mg/L mg/L mg/L (°C)
July-03
4.73
349.70
185.30
30.71
2.57
2.17
7.02
1.39
246.00
20.63
August-03
4.85
321.81
195.35
32.15
2.74
2.42
6.62
1.66
238.00
21.61
September-03
4.76
337.59
202.32
33.61
2.81
2 17
7.38
1.39
245.20
21.11
0ctober-03
4.48
464.50
221.38
35.16
3.44
2.75
8.38
1.78
244.25
19.98
November-03
4.35
337.33
218.90
33.85
3.81
2.80
7.32
1.27
230.75
17.80
December-03
4.09
322.13
216.83
33.76
3.80
2.77
7.12
1.79
229.40
15.93
January-04
4.28
308.10
206.86
34.88
4.18
2.86
7.36
2.64
204.25
14.62
February-04
4.34
236.71
202.38
35.88
4.36
3.18
7.35
1.57
214.50
14.21
March-04
4.25
376.22
217.35
37.00
4.09
3.03
7.72
2.27
225.40
15.32
April-04
4.52
406.29
223.05
34.83
3.44
2.71
8.56
1.81
234.00
16.18
May-04
4.59
422.77
221.23
36.39
2.99
2.31
8.39
0.93
241.00
17.91
June-04
4.58
421.05
187.73
32.92
2.23
2.10
7.70
1.10
245.00
19.50
July-04
4.90
420.32
188.24
28.41
1.96
2.19
7.12
0.82
242.00
20.77
August-04
5.06
370.22
192.05
29.66
1.72
1.81
6.91
0.92
244.25
21.51
September -04
4.93
368.82
224.00
32.27
1.72
2.04
6.95
1.18
243.33
21.02
October-04
4.93
299.29
226.62
34.48
2.21
1.84
7.66
1.37
249.00
19.80
November-04
4.56
309.82
200.45
36.48
2.57
3.01
7.67
1.86
242.20
17.11
December-04
4.49
315.23
217.14
36.52
2.97
3.45
7.29
1.98
254.50
15.25
January-05
4.37
374.32
201.27
40.43
3.17
3.45
8.24
3.87
297.00
14.27
311
February-05
4.38
368.21
214.53
40.59
3.25
2.63
9.66
5.05
283.00
14.13
March-05
4.41
380.12
199.60
39.07
3.25
2.88
8.61
3.67
284.00
14.87
April-05
4.75
323.20
169.05
44.30
2.00
1.72
7.90
4.29
274.50
15.85
May-05
4.89
375.57
189.57
34.29
1.87
2.38
9.97
4.67
301.00
17.03
June-05
5.18
372.09
203.32
39.46
0.06
0.12
9.86
4.75
303.40
18.59
July-05
4.83
379.29
197.57
31.08
0.11
0.15
9.50
3.79
254.50
20.61
August-05
5.25
428.22
204.57
29.46
0.06
0.08
8.74
3.10
246.60
21.58
September-05
5.07
419.41
210.82
30.10
7.58
3.08
251.25
21.26
October-05
5.53
344.74
185.65
31.12
7.00
2.89
262.50
19.75
November-05
5.10
292.09
213.36
34.89
9.04
4.93
296.80
18.06
December-05
4.93
285.26
220.43
37.88
10.56
6.77
273.50
15.68
January-06
4.75
375.76
232.29
38.27
12.38
7.22
255.75
15.16
February-06
4.49
373.37
218.10
35.50
8.69
3.51
250.75
14.09
March-06
4.31
313.02
221.91
41.88
9.06
4.25
256.60
14.35
April-06
4.47
375.76
229.17
39.02
9.98
5.11
258.50
15.83
Table 6.13 Effluent data. Final effluent characteristics
312
Month, year
TSS mg/L
Total BOD5mg/L
NH3-N mg/L
NOX-N mg/L
Total P mg/L
PO4-P mg/L
Temp. (°C)
July-03
2.37
2.58
0.25
7.16
0.15
0.01
20.63
August-03
2.74
1.60
0.17
6.80
0.74
0.48
21.61
September-03
2.55
2.23
0.18
5.78
1.69
0.99
21.11
October-03
2.96
2.66
0.25
6.33
1.53
0.90
19.98
November-03
2.63
1.62
0.22
7.70
1.12
0.56
17.80
December-03
4.11
1.88
0.25
8.62
1.42
1.01
15.93
January-04
5.73
2.75
0.77
10.48
1.37
1.15
14.62
February-04
4.96
2.33
0.95
9.28
1.17
0.55
14.21
March-04
3.25
2.24
1.02
9.33
2.15
1.45
15.32
April-04
5.58
2.36
0.32
6.78
1.25
0.80
16.18
May-04
4.00
1.85
0.23
7.20
0.18
0.01
17.91
June-04
3.64
1.72
0.16
7.20
0.29
0.08
19.50
July-04
3.70
1.78
0.26
6.15
0.24
0.03
20.77
August-04
2.85
1.91
0.16
5.79
0.23
0.05
21.51
September-04
3.12
1.62
0.17
5.77
0.16
0.03
21.02
October-04
2.22
1.68
0.14
5.93
0.58
0.38
19.80
November-04
3.33
1.67
0.16
7.55
0.48
0.29
17.11
December-04
4.48
2.13
0.17
7.65
0.78
0.46
15.25
January-05
4.93
2.19
0.17
8.21
1.04
0.71
14.27
313
February-05
5.91
2.66
0.25
8.26
2.63
2.03
14.13
March-05
4.85
2.21
0.25
7.20
1.18
0.87
14.87
April-05
4.92
2.34
0.41
8.03
0.79
0.46
15.85
May-05
4.08
2.51
0.26
8.37
1.48
1.19
17.03
June-05
2.57
2.43
0.14
11.56
1.33
1.11
18.59
July-05
2.54
2.10
0.18
10.89
1.54
1.50
20.81
August-05
3.32
2.10
0.16
11.70
1.17
0.92
21.58
September-05
3.09
1.88
0.11
10.78
0.44
0.38
21.26
October-05
3.76
1.93
0.10
11.42
1.31
1.34
19.75
November-05
4.87
3.33
0.19
12.85
2.43
2.06
18.06
December-05
6.20
3.57
0.21
10.80
1.81
1.18
15.68
January-06
5.87
3.94
1.21
1.83
1.31
15.16
February-06
5.36
3.58
0.04
11.20
1.79
1.42
14.09
March-06
8.17
4.33
0.52
11.78
3.12
2.38
14.35
April-06
7.30
3.00
0.18
12.28
2.05
1.68
15.83
Biofilm Reactors
Aerobic SRT Aerobic MLSS Effluent NH3-N SVI
7.0 6.0
180 160 140 120
5.0
100
4.0
80
3.0
60 40
1.0
20
0.0
0
Ju
lSe 03 pN 03 ov -0 Ja 3 nM 04 ar M 04 ay -0 Ju 4 l-0 Se 4 pN 04 ov Ja 04 nM 05 ar M 05 ay -0 Ju 5 lSe 05 pN 05 ov Ja 05 nM 06 ar -0 6
2.0
SVI (mL/L)
8.0 MLSS (g/L); Aerobic SRT (d); NH3-N (mg/L)
314
Date
Figure 6.12 Monthly operating data. process at The Colony could consistently meet the original design intent for BOD (20 mg/L) and TSS removal (20 mg/L). In fact, the plant was routinely capable of meeting its permit limitations of 10 mg/L BOD and 15 mg/L TSS at lower average daily flow conditions. Historical influent loadings and the effluent discharge performance for 2004 are shown in Table 6.14. However, the existing contact stabilization plant was not designed to provide nitrification to meet the new 3-mg/L NH3-N effluent limits. The plant was able to nitrify ammonia, to some extent, during warm weather, but not to the lower levels in the new discharge permit. Additionally, the rapid residential development in the City of The Colony quickly was overloading the existing treatment units, causing reduced treatment efficiency. Therefore, the City was in need of substantial increases in their system capacity. Under the conditions of the 3-mg/L ammonia discharge limit, the City was faced with two choices—downrate the existing facility to approximately 9100 m3/d (2.4 mgd) to achieve ammonia removal in the available plant basin volume, or expand the plant to achieve ammonia removal and increase capacity. Obviously, with the capacity expansion needs, a downrating of the facility was not possible. Therefore, expansion of the plant was needed. The evaluation of the upgrade alternatives recommended that the contact stabilization basins be modified to a plug-flow arrangement, which included the use of an IFAS system.
Hybrid Processes
Table 6.14 Colony, Texas, 2004 historical influent loadings and effluent discharge. Parameter
Influent loadings
Effluent discharge
Average annual flow
8400 m /d (2.22 mgd)
8400 m3/d (2.22 mgd)
BOD
236 mg/L
6.5 mg/L
TSS
324 mg/L
4.2 mg/L
Ammonia-nitrogen
30 mg/L
9.4 mg/L
3
The conversion to the plug-flow basins with IFAS would improve nitrification of ammonia, allow the plant to return to compliance with the discharge limits, and allow for capacity expansion of the facility. As an added bonus, the implementation of the IFAS system did not require additional basin construction, resulting in a significant cost savings over conventional treatment. In making these changes, the City was able to meet the ammonia discharge requirements of the permit and expand the plant under these conditions, from a downrated 9100-m3/d (2.4-mgd) capacity to a new 170 000-m3/d (4.5-mgd) treatment capacity (essentially doubling the plant capacity); improve effluent quality; and meet new permit requirements—all without the construction of new basins. Both fixed-rope-type media and free-floating plastic-media systems were considered. The free-floating media would require a medium-bubble aeration system and screens to retain the media. In The Colony, an existing membrane diffused air system already was in place and was to be reused in the design. In addition, challenging hydraulic conditions made the use of retention sieves more restrictive. A fixed-media system was selected to accommodate the existing fine-bubble aeration system, improve blower efficiency, and avoid the headloss associated with retention screens for floating media. The selected media system, BioWeb (Entex Technologies), uses polyester fabric media (Figure 6.13) on steel frames with dedicated coarse-bubble mixers underneath the media, to ensure good mixing and oxygen transfer, provide a thin biofilm, and control any potential redworm predators. The existing configuration is a “bull’s-eye” plant, with the aeration basin in an exterior annulus and the final clarifier in the center of the “bull’s-eye.” The aeration basins consisted of two separate treatment trains—train A and train B. Train A is
315
316
Biofilm Reactors
Figure 6.13 Polyester fabric media.
sized for approximately 60% of the plant capacity, with train B sized for approximately 40%. Modifications to the plant would include the removal of the divider walls between the contact, reaeration, and digester zones and the creation of a new plug-flow configuration. Solids no longer would be digested aerobically and instead would be wasted directly to a sludge holding tank for subsequent dewatering and disposal in the municipal landfill. The new plug-flow design would include two swing zones intended to function as anaerobic zones for biological phosphorus removal and the addition of IFAS units in the aerated zones. An effluent splitter was added to each aeration basin to split flow to the central final clarifier and the existing exterior final clarifiers. Figures 6.14 and 6.15 provide diagrams of the existing and proposed process configurations.
6.4.2 Changing Design Conditions On the day The Colony’s City Council was to award the contract for construction of the plant improvements, the state regulatory agency notified the City that the plant also would be receiving an effluent phosphorus limit, probably 1 mg/L. The primary reason for the new limit was the proposed increase in permitted discharge flow from the plant to approximately 170 000 m3/d (4.5 mgd). Under these flow conditions, the State determined that the location of the outfall would not be capable of sustaining adequate dissolved oxygen and water quality, as a result of the additional nutrient load. Therefore, a reduction in total phosphorus loading was necessary to achieve the increased capacity.
Hybrid Processes
Reaeration Zone
Final Clarifier
Aerobic Digester
Final Clarifier
Contact Zone
Anaerobic #2
Effluent Splitter
Anaerobic #1
Figure 6.14 Existing and proposed aeration basin configuration.
Clarifier Anaerobic 0 – 0.5 hr
Anaerobic 0 – 0.5 hr
Aerobic 4 – 6 hr
RAS (50 – 100%) Anoxic Carbon 0.5-1 hr WAS Containing P
Figure 6.15 Proposed process flow diagram. The City immediately implemented a modification to the design to include new anaerobic zones. These new zones were designed with both coarse-air diffusers and mixer systems, to allow them to operate as swing zones, and added to the contract through a change order. The ability to change the design to a new modified Johannesburg process configuration was enhanced greatly by the use of IFAS. The additional treatment and nitrification capacity provided by the IFAS units allowed for additional partitioning of the existing basin volume to provide dedicated anaerobic zones. Although the new phosphorus limit is required to meet the future 170 000-m3/d (4.5-mgd) flow condition, the implementation of the limit is phased. Therefore, The
317
318
Biofilm Reactors
Colony has time to fine-tune and adjust the new process appropriately to consistently meet the new 1-mg/L limit. Effluent phosphorus levels have shown the ability to be maintained below 1 mg/L during construction. The plant staff is currently finetuning the process to achieve consistent phosphorus removal as the plant’s nitrification increases.
6.4.3 Plant Construction and Operation During construction, each aeration train needed to be out-of-service for several months to perform the modifications and convert each to a plug-flow configuration. This meant that all of the plant flow had to be treated by only one train at a time, while the other train was being modified. In the initial phase, in September 2005, the existing contact and reaeration zones in train A were retrofit with the IFAS media as a temporary treatment solution, while train B was out-of-service for permanent modification. The installation of the IFAS units in the basin was simple. This “temporary” installation of 10 media modules took only 1 day for construction (Figure 6.16). This included time to dewater the basin, add the frames, and refill the tank. The tank was brought back into service within a 48-hour period. For approximately 5 months, the entire plant flow was treated in this small hybrid contact stabilization IFAS reactor. In February 2006, aeration train B was started up in full plug-flow integrated fixed-film activated sludge mode, with 10 new IFAS media modules. During this startup period, the plant experienced several challenges, including failures in the existing fine-air-diffuser grid and problems in activated sludge wasting. While these problems were not related to the IFAS media installation, they hampered the ability of train B to consistently achieve reduced ammonia concentrations below the 3-mg/L limit. Average daily flow during this period was approximately 8520 m 3/d (2.25 mgd), and influent BOD loadings averaged approximately 530 kg/m3 (33 lb/1000 cu ft) of the aeration basin volume. At this flow and loading, train B was loaded effectively at a near-maximum hydraulic condition and at a significant biological loading for average daily flows. However, even with the construction issues identified above, the plant complied with its BOD and TSS limits, while putting 100% of the plant flow through only 40% of the total plant reactor volume. Figure 6.17 shows the 7-day average effluent BOD concentrations from the plant upon startup of train B. In September 2006, aeration train A was started in full plug-flow IFAS mode and experienced none of the diffuser and wasting issues that train B suffered. During this time, train B was taken out-of-service to allow for repairs of the damaged diffuser
Hybrid Processes
Figure 6.16 Installation of 10 IFAS media modules in existing basin. Effluent CBOD
10.00
7 -0 Ju l
7 M
ay
-0
7 M
ar
-0
07 Ja
n-
06 ov N
06 pSe
6 -0 Ju l
ay M
-0 ar M
06
0.00 6
Effluent CBOD (mg/L)
20.00
Month Effluent CBOD
Figure 6.17 Effluent cBOD concentration after train B startup.
grid, and 100% of the plant flow was placed through train A, which represents 60% of the total aeration volume. The average daily flow during this period was approx imately 8820 m3/d (2.33 mgd), and influent BOD loadings averaged approximately 500 kg/m3 (31 lb BOD/1000 cu ft) of the basin volume. Immediate improvement in BOD and ammonia removal was noted, and the plant achieved compliance with the 3 mg/L ammonia limit within a few days after startup.
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Figure 6.18 Effluent ammonia concentration (7-day average). During this period, train A was run at near 85% of maximum design flow capacity and was able to consistently meet effluent discharge requirements of 10, 15, and 3 mg/L of BOD, TSS, and NH3-N, respectively. This was especially significant because of the lower temperatures during these winter months. Figure 6.18 shows the 7-day average effluent ammonia concentrations before, during, and after plant construction. The data clearly demonstrate the dramatic change in effluent ammonia quality upon implementation of the improvements in train A. This was maintained until the end of March 2007, when the facility was faced with another challenging event.
6.4.4 System Flexibility At the end of March 2007, the WWTP once again was beset with a process challenge. The existing diffuser grid broke at one of the diffuser joints, causing a significant release of air under one of the IFAS units. The dramatic increase in air and buoyancy resulted in one IFAS unit working loose from its anchor bolts. Train A was taken out-of-service, on an emergency basis, as a result. To further complicate matters, a failure in the RAS return valves resulted in automatic wasting of the MLSS in train B to below 500 mg/L. As a result, the plant lost nitrification during this period. However, it is important to note that, even with extremely low suspended biomass concentrations, the plant was able to maintain compliance with BOD and TSS limits. The biomass on the IFAS remained intact during this period, unaffected by the error in wasting, and provided the majority of treatment. This allowed for the plant to maintain permit compliance with BOD and TSS, while reestablishing autotrophs
Hybrid Processes
in the remaining basin. The IFAS system functioned as intended during this critical period, mitigating what could have been a severe excursion beyond the allowable effluent limits.
6.4.5 Redworm Predation During the period from February to September 2006, train B was operating alone, while train A was undergoing permanent modification. A coupon of IFAS media was installed during this period to allow for observation of biomass growth on the IFAS systems. This small coupon was removed periodically for visual observation and analysis under the microscope. As part of this observation, it was noted that the coupon contained redworm predators . Plant staff had indicated that they routinely would have blooms of redworms during the summer months, even before the installation of IFAS media, so the observation of redworms in the mixed liquor was not considered unusual. In an effort to control the growth of the worms and assist in control of the biomass thickness, the IFAS manufacturer provided a modification to the existing media modules to include a patented coarse-bubble system, which would allow for additional air-scour of the modules on a periodic basis. The system was connected to the diffused air header using simple solenoid valves and timer systems to control the release of air. The coarse-bubble diffuser below the units would release a periodic burst of air, alternating through the individual IFAS units in the basin and then repeating. The retrofit was installed in both trains A and B. Even with the appearance of the redworms and their growth on the media, there has been no indication of any effect on the health or effectiveness of the IFAS biomass, as evidenced by the performance of the units during the emergency shutdown in April 2007. The redworms have not had any other deleterious effects on plant performance. They typically are settled in the clarifier underflow, where the lack of oxygen causes the majority to perish. Those that do escape in the clarifier effluent are captured in the plant’s cloth disk filters, with no negative effect on filter performance.
7.0 References Austin, E. P.; Walker, I. Publication of Ashbrook-Simon-Hartley (Houston, Texas). Copithorn, R.; Schwinn, D. E. S.; Mitta, P. R.; Sen, D. (1995) Evaluation of FullScale Design Factors in Integrated Fixed Film Activated Sludge Processes
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for Biological Nitrogen and Phosphorus Removal. Proceedings of the New and Emerging Environmental Technologies and Products Conference for Wastewater Treatment and Stormwater Collection, Toronto, Canada, June 4–7; Water Environment Federation: Alexandria, Virginia. Cooper, P. (1989) Demonstration and Evaluation of the CAPTOR Process for Sewage Treatment; EPA-600/S2–88-060, Risk Reduction for Engineering Laboratory; U.S. Environmental Protection Agency: Cincinnati, Ohio. Kato, K.; Sekikawa, Y. (1967) FAS (Fixed Activated Sludge) Process for Industrial Waste Treatment. Proceedings of the 22nd Purdue Industrial Waste Conference, West Lafayette, Indiana, May 2–4; Purdue University: West Lafayette, Indiana. Odegaard, H.; Rusten, B.; Westrum, T. (1994) A New Moving Bed Biofilm Reactor—Application and Results. Proceedings of the 2nd International Specialized Conference on Biofilm Reactors, Paris, France, Sept 29–Oct 1; International Association on Water Quality: London, United Kingdom, 221–229. Randall, C. W.; Edwards, H. R.; King, P. H. (1972) Microbial Process for Acidic Low-Nitrogen Wastes. J. Water Pollut. Control Fed., 44, 401–413. Sen, D.; Copithorn, R.; Randall, C.; Jones, R.; Phago, D.; Rusten, B. (2000) Investigation of Hybrid Systems for Enhanced Nutrient Control, Project 96-CTS-4; Water Environment Research Foundation: Alexandria, Virginia. Sen, D.; Farren, G. D.; Copithorn, R. R.; Randall, C. W. (1993) Full-Scale Evaluation of Nitrification and Denitrificaton Fixed Film Media (Ringlace) for Design of Single Sludge System. Proceedings of the 66th Annual Water Environment Federation Technical Exposition and Conference, Anaheim, California, Oct 3–7; Water Environment Federation: Alexandria, Virginia, 137–148. Sen, D.; Mitta, P.; Randall, C. W. (1994a) Performance of Fixed Film Media Integrated in Activated Sludge Reactors to Enhance Nitrogen Removal. Water Sci. Technol., 30, 13–24. Sen, D.; Randall, C. W.; Copithorn, R. R.; Huhtamaki, M.; Farren, G.; Flournoy, W. (2007) Understanding the Importance of Aerobic Mixing, Biofilm Thickness Control and Modeling on the Success or Failure of IFAS Systems for Biological Nutrient Removal. Water Practice, 1 (5), 1–18.
Hybrid Processes
Sen, D.; Randall, C. W.; Jenson, K.; Farren, G. D.; Copithorn, R. R.; Young, T. A.; Brink, W. P. (1994b) Design Parameters for Integrated Fixed Film Activated Sludge (IFAS) Processes to Enhance Biological Nitrogen Removal. Proceedings of the 67th Annual Water Environment Federation Technical Exposition and Conference, Chicago, Illinois, Oct 15–19; Water Environment Federation: Alexandria, Virginia, 713–724.
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Chapter 7
Biological Filters 1.0 Introduction 2.0 Descriptions of BIOLOGICALLY ACTIVE FILTER Reactors and Equipment
3.1 Mineral Media
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3.2 Random Plastic Media 343 3.3 Modular Plastic Media 345 329
2.1 Brief History of Biologically Active Filters 329 2.2 Downflow Biologically Active Filter with Sunken Media 331
4.0 Backwashing and Air-Scouring
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5.0 BIOLOGICALLY ACTIVE FILTER Process Design 349 5.1 Process Design for Secondary Treatment
351
2.3 Upflow Biologically Active Filter with Sunken Media 334
5.1.1 Volumetric Biochemical Oxygen Demand Loading 351
2.4 Upflow Biologically Active Filter with Floating Media 335
5.1.2 Hydraulic Loading
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2.5 Moving-Bed, Continuous Backwash Filters 337
5.1.3 Backwashing
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5.1.4 Design Example: Design of a Submerged, Upflow Biological Aerated Filter System for Secondary Treatment (No Nitrification) 353
2.6 Non-Backwashing, Open-Structure Media Filters 339 3.0 Media for Use in BIOLOGICALLY ACTIVE FILTERS
341 325
5.1.5 Solution
353 (continued)
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5.2 Process Design for Nitrification
355
5.2.1 Influence of Hydraulic Filtration Rates 355
5.4.6 Tertiary Denitrification Typical Operations Issues and Corrective Actions 370
5.2.2 Effect of Process Air Velocity 357
5.4.6.1 Excess Backwashing 370
5.2.3 Dependence on Loading Conditions
5.4.6.2 Gas (Nitrogen) Accumulation 371
5.2.4 Temperature Effects
359 360
5.2.5 Design Example: Design of a Submerged, Upflow Biological Aerated Filter System for Nitrification Following Secondary Treatment 360 5.2.6 Solution
361
5.3 Process Design for Combined Nitrification and Denitrification 362 5.4 Process Design for Tertiary Denitrification 365 5.4.1 Volumetric Mass Loading
365
5.4.2 Half-Order Kinetic Model 367 5.4.3 Hydraulic Loading 368
5.4.6.3 Solids Breakthrough 371 5.4.6.4 Nitrate/Nitrite Breakthrough 371 5.4.6.5 Carbon Breakthrough 372 5.4.6.6 Phosphorus Management 372 5.4.6.7 Operation During Peak Flow Events372 5.5 Phosphorus Removal Considerations for Biologically Active Filter Processes 373 6.0 Design Considerations
374
5.4.4 Solids Removal and Sludge Production 368
6.1 Preliminary and Primary Treatment 374
5.4.5 Supplemental Carbon Requirements 369
6.2 Backwash Handling Facilities
374
(continued)
Biological Filters
6.3 Biologically Active Filter Process Aeration 375 6.3.1 Oxygen-Transfer Efficiency 375 6.3.2 Process Air Distribution Systems 377 6.3.3 Process Air Control 377 6.4 Supplemental Carbon Feed Requirements 378 7.0 Biologically active filter Case Studies 379 7.1 Chemically Enhanced Primary Treatment Followed by Two-Stage Biologically Active Filter for Total Nitrogen Removal: VEAS Wastewater Treatment Plant, Oslo, Norway 379
7.2 Chemically Enhanced Primary Treatment Followed by Three-Stage Biologically Active Filter for Total Nitrogen Removal: Siene Centre Wastewater Treatment Plant, Colombes, France 381 7.3 Total Nitrogen Removal in a Single-Stage Biologically Active Filter: Frederikshavn Wastewater Treatment Plant, Denmark 384 7.4 Nitrification and Denitrification: West Warwick, Rhode Island 387 7.5 Post-Denitrification Sand Filters: Havelock, North Carolina 389 8.0 REFERENCES
391
1.0 Introduction Biological wastewater treatment and suspended-solids removal are carried out in biologically active filters (BAFs), under either aerobic or anoxic conditions. In a BAF, the media acts simultaneously to support the growth of biomass and as a filtration medium to retain filtered solids. Accumulated solids are removed from the BAF through backwashing. There is a direct interaction between the media characteristics and the process, because the configuration (sunken media or floating media) and flow and backwash regimes depend on the media specific gravity. Media may be natural mineral, structured plastic, or random plastic. The BAF reactors can be used for carbon oxidation (biochemical oxygen demand [BOD] removal) only, combined BOD removal and nitrification, combined nitrification and dentrification, tertiary nitrification, and tertiary denitrification. Once the raw wastewater has undergone screening, grit removal, and primary treatment, the BAF
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process can comprise full secondary treatment for a facility or can be constructed for operation in parallel to an existing secondary treatment process. Using BAF as a tertiary treatment process for nitrification and/or denitrification as an upgrade to existing secondary processes is common. A typical process flow diagram for four different BAF options is provided in Figure 7.1. Historically, the acronym BAF has meant biological aerated filter, and the term typically has been used to refer to an aerated biofilter used for secondary treatment and nitrification. However, the acronym BAF is being expanded herein to cover all (a)
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Figure 7.1 BAF process flow diagrams.
Effluent
Backwash Water
Sludge
Biological Filters
biologically active filters, including those that operate under anoxic conditions for denitrification (these processes historically have been referred to as denitrification filters). The BAF reactors can be characterized into groups, according to their media configurations and flow regime, as follows: • Downflow BAF, with media heavier than water. This general category includes both the Biocarbone reactors (OTV, Paris, France) commercially marketed in the 1980s for secondary and tertiary treatment and packed-bed tertiary denitrification reactors, such as Tetra Denite filters (Severn Trent Water Purification, Inc., Ft. Washington, Pennsylvania). These BAFs are backwashed using an intermittent countercurrent flow regime. • Upflow BAF, with media heavier than water. This includes BAF reactors for secondary and tertiary treatment that use expanded clay and other mineral media, such as the Degremont Biofor (Degremont, Rueil-Malmaisson Cedex, France). These BAFs are backwashed using an intermittant co-current flow regime. • BAF with floating media. This includes BAF with polystyrene, polypropylene, or polyethylene media, such as the Kruger Biostyr (Veolia, Paris, France). These BAFs are backwashed using an intermittant counter-current flow regime. • Continuous backwashing filters. These filters operate in an upflow mode and consist of media heavier than water that continuously moves downward, countercurrent to the wastewater flow. Media is directed continuously to a center airlift, where it is scoured, rinsed, and returned to the top of the media bed. • Non-backwashing, submerged filters. These processes consist of submerged, static media and often are referred to as submerged aerated filters (SAFs), although there has been recent work in applying this technology under anoxic conditions for denitrification. Solids are intended to be carried through the reactor and removed through a dedicated solids separation process.
2.0 Descriptions of BIOLOGICALLY ACTIVE FILTER Reactors and Equipment 2.1 Brief History of Biologically Active Filters Initial developments of fixed-film submerged bioreactors were for nitrate removal under anoxic conditions. The first applications were in the United States (Chen, 1980); in France, with biological denitrification (Biodenit, [OTV/Veolia, Paris, France]); and
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in Germany, with floating filters for denitrification (Preussag Denipor, [Hanover, Germany]). The application of sand filters for nitrogen removal preceded the first aerobic applications, as the first full-scale static-bed denitrification filters were commissioned in Florida in the late 1970s (Pickard et al., 1985). The BAF and similar processes appeared in the literature at various times (Canadian Patent, cited by Smith and Hardy, 1992), and, in the United States, an early application of the technology for wastewater treatment at a paper mill was reported (Brown, 1992). Larger municipal application first occurred in France, where the process was developed in the late 1970s, based on Italian patents, up to the startup of the first commercial Biocarbone (OTV/Veolia, Paris, France) in Soissons, north of Paris, in 1982. In parallel, the process also was applied in drinking water, blowing air through a downflow filter bed to enhance nitrification (Sibony, 1982). While early BAFs used a sunken mineral media, BAFs using plastic media were developed during the late 1980s in France (Rogalla and Bourbigot, 1990) and in the United Kingdom (Meaney, 2007; Whitaker et al., 1993). Although most of the BAF technologies developed, to date, are of the intermittent backwash type, various attempts have been made to develop BAFs that backwash continuously, including the following: • Stephenson (1996) described a device called a recirculating plastic media biological aerated filter (REBAF), which used plastic medium that was removed from the top of the reactor with a screw conveyor and recycled to the bottom after washing. • Continuously backwashing sand filters originally were developed in Scandinavia and have found wide application in smaller plants for tertiary polishing. They first were adapted as anoxic biological reactors for denitrification in the late 1980s. These moving-bed filters also have been modified by injecting air to the media bed to support operation for tertiary nitrification. The impetus for more widespread application of BAF technology in Europe came with the European Directive on the Treatment of Urban Wastewaters (European Union, 1991), which resulted in a need to upgrade wastewater treatment facilities on restricted sites. In Europe, BAF was adopted quickly by the private operators in France, who each had developed their own versions, and in the United Kingdom after the privatization of the water industry in England and Wales in 1989. The private business models in France and the United Kingdom, driven by the search for more efficient technologies and lower cost, favored procurement
Biological Filters
procedures that allowed turn-key bidding and competition between many suppliers. The majority of BAF plants can be found in those two countries. More recently, a focus on nutrient control on constrained treatment plant sites has resulted in a number of new installations in the United States. Open structure, non-backwashing biological filters (also known as submerged aerated filters [SAFs]) were developed during the same timeframe as BAF configurations, with intermittent and continuous backwashing. A non-backwashing BAF process patented in 1984 comprised mixing the effluent with air and circulating the mixture through a bed of structured plastic media before passing the effluent on for further treatment for suspended-solids removal (Tolley Process Engineering Ltd., 1981). Initially, this type of plant was used for high-rate roughing treatment of high-strength wastewater and industrial effluents in the United Kingdom (Churchley, Jarvis and Pickett, 1990), and with aeration using fixed biofilm carriers in Germany (Schlegel and Teichgraber, 2000). One type of system, using modular plastic media, was developed in Europe, predominantly in Switzerland and Germany (Ryhiner et al., 1992), and often was applied to nitrification of secondary treated effluent on a large scale. A granular BAF mostly was used to capture the solids downstream, either in aerated mode or for denitrification. The SAF process was developed further through the 1990s, by a number of manufacturers using a variety of media and techniques, for example, random or structured plastic medium, and different types of trickling filter packings, as reflected by the number of patents for various improvements (e.g., Froud, 1994).
2.2 Downflow Biologically Active Filter with Sunken Media The general process arrangement of a downflow BAF is shown in Figure 7.2. Air is sparged into the lower zone of the downflow submerged granular bed of expanded shale to produce good oxygen-transfer efficiency by countercurrent gas–liquid flow and a circuitous flow path through the media. Wastewater is fed to the top of the packed granular media, where the combined effects of airflow and headloss distribute the flow across the bed. Most of the larger total suspended solids (TSS) particles are entrapped near the top of the media. Therefore, fine screening is not necessary, as the larger particles are retained quickly and removed from the reactor by each backwash, without coming in contact with the filter nozzles at the bottom of the filter. As the media begins to foul, the headloss through the filter increases, which results in an increase in the water depth above the media. Countercurrent backwashing of the filter removes accumulated solids and excess biofilm growth. The
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Raw Water
Treated Water
Granular Medium Sludge Backwash Water Sludge Treatment Air Process
Air Scour
Figure 7.2 BioCarbone® biological filter (courtesy of OTV).
backwashing schedule is programmed based on headloss and time schedules, trying to avoid backwashes during peak flow periods. The OTV Biocarbone system was a commercially available downflow BAF that was installed in more than 100 plants throughout the world during the early 1980s. The municipal plant sizes varied from 2000 m3/d (0.5 mgd) to 80 000 m3/d (21 mgd), and there also were a number of smaller plants in Japan for building recycle or industrial applications. These plants are capable of providing secondary treatment and advanced wastewater treatment, including biological nutrient removal. While the process performance of this type of BAF reactor was improved over previous practice, the countercurrent air and water flow limited its application for BOD removal and nitrification. Air would become entrapped in the accumulated solids on the surface and at the top of the media. Headloss then could increase unpredictably, and backwashing would be necessary. Some remedies were applied, such as intermittent aeration at higher rates to expand the bed, or “mini-backwashes” to expel the excess solids from the surface. However, for secondary treatment applications, the downflow BAF ultimately was replaced by upflow configurations, which could operate at higher hydraulic rates and handle wider hydraulic variations. A downflow BAF configuration with sunken media was developed successfully, however, for tertiary denitrification applications (Figure 7.3) by Tetra Process Technologies (now Severn Trent, Coventry, United Kingdom). The Denite process configuration has been used since the late 1970s for meeting stringent total nitrogen
Biological Filters
Filtration
Backwashing
Filtrate Backwash filter Backwash water
Figure 7.3 Downflow denitrification filter (courtesy of Severn Trent). limits, while providing a filtered effluent. Methanol or another carbon source is added to the influent wastewater to provide substrate for denitrification. A typical installation includes 1.8 m (6 ft) of 2- to 3-mm sand media over 457 mm (18 in.) of graded support gravel. F.B. Leopold Company (Zelienople, Pennsylvania) also offers a similar denitrification BAF configuration. Several conventional deep-bed filter installations have been retrofitted over the years for this application. These installations vary, in terms of the type of underdrain support used, but it is important that the underdrain design be compatible with biofilm growth. For example, early testing showed that nozzle-type underdrains were prone to fouling (Pickard et al., 1985). In a downflow denitrification BAF, the backwash cycle typically consists of a brief air-scour followed by an air-water backwash and water rinse cycle. Design backwash water and air-scour flowrates typically are 15 m3/m2·h (6 gpm/sq ft) and 90 m3/m2·h (5 cfm/sq ft), respectively. The filter influent and backwash piping are similar to that of conventional filters. Backwash water usage typically is 2 to 3% of the average flow being treated. Nitrogen gas accumulates within the media and is released by pumping backwash water up through the media bed for a short duration. The denitrification capacity between nitrogen-release cycles typically ranges from 0.25 to 0.5 kg NOx-N/m2 (0.05 to 0.10 lbs NOx-N/sq ft) (McCarty, 2008).
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2.3 Upflow Biologically Active Filter with Sunken Media The upflow mode of BAF operation through a sunken granular bed has been used in more than 185 installations worldwide. The Degremont Biofor reactor is a commercially available upflow BAF and may be used for BOD removal, nitrification, and denitrification. Its general process arrangement is shown in Figure 7.4. Solids are trapped mostly in the lower part of the media bed during normal operation and are backwashed, as required, by increasing the hydraulic rate and applying scour air. As the backwash consists of co-current scour air and backwash water, the accumulated solids travel up through the media bed before being released at the top. Three types of media can be used in the Biofor BAF, depending on the application. The media consists of expanded clay or expanded shale, either in the form of spherical grains (with an effective size of 3.5 or 4.5 mm) or as angular grains (with an effective size of 2.7 mm). The media form a submerged, fixed bed in the bottom of the reactor, typically a height of 3 to 4 m (9.8 to 13.1 ft), with approximately 1 m (3.3 ft) of freeboard zone above the media. The clean surface area of grains is approximately 1640 m2/m3 (500 sq ft/cu ft). Influent water is introduced to the bed through a filter plenum and nozzle air/water distribution system. The nozzles are installed in a false floor located approximately 1m (3.3 ft) above the filter floor. The influent flow must be fine-screened to prevent blockage of the nozzles. Backwash water and scour air is introduced through the same plenum/nozzle system. Process air is introduced through separate air diffusers located in the media bed above the inlet nozzles. To reduce the quantity of backwash water and the risk of media loss, the backwash starts with a drain down. The duration of the drain down is affected by the level of solids accumulation and determines the need for more vigorous backwashing. The Treated Water
Process Air or C-Source Raw Water Air Scour
O2/N2
O2/N2
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Sludge
Sludge Treatment
Figure 7.4 Biofor® upflow biological filter (courtesy of Degremont).
Biological Filters
co-current backwash then consists of an air-scour to break up the media, followed by an air/water wash, and finally by a water rinse. During backwashing, the solids are pushed from the bottom of the bed and transferred into the freeboard zone between top media level and the discharge, and on to waste. Backwash water quantities equivalent to several times the volume of the freeboard zone are needed to reduce solids levels to discharge limits, as they are released directly from the filter. The effect of the discharge of these solids depends on the actual treatment objectives and number of BAF cells. This can be addressed by incorporating a “filter-to-waste” step at the end of the backwash cycle. Alternately, increasing the total volume of flushing water may be necessary to improve the effluent quality following backwashing (Michelet et al., 2005). A key issue with backwash of sunken media systems is the potential for “boils” during backwashing. For even backwashing, the water flow must be well-distributed across the plan area of the BAF, and therefore the headloss across the distribution system must be greater than the headloss through the bed. If the bed becomes blocked because of high loads or insufficient backwash, then its headloss becomes the controlling factor. The flow will short-circuit through the line of least resistance. This will result in a “boil” or violent eruption of the flow through the point of least resistance. Similar short-circuits and boils also can occur if the nozzles are blocked. These “boils” during backwashing can result in excessive media loss.
2.4 Upflow Biologically Active Filter with Floating Media These processes use a floating bed of media to provide biological surface area and filtration. This process first was used in industrial filtration and drinking water denitrification (Roennefahrt, 1986). Later, coarse-bubble aeration diffusers were introduced at the bottom of the media to enhance the contact of air, water, and biomass (Rogalla and Bourbigot, 1990). While the latter, the Biostyr process, uses very lightweight expanded polystyrene, a process using recycled polypropylene with a specific gravity slightly lower than 1—the Biobead (Brightwater F.L.I., Hertfordshire, United Kingdom)—has found large application in the United Kingdom. The Veolia Biostyr (Paris, France) unit (see Figure 7.5) is a reactor that is partially filled with small (2- to 6-mm) polystyrene beads. Process objectives determine selection of the bead size; larger beads can be loaded more heavily, and smaller beads typically achieve higher process performance. The beads, which are lighter than water, form a floating bed in the upper portion of the reactor, typically a height of 3 to 4 m (9.8 to 13.1 ft), with approximately 1.5 m (4.9 ft) of free zone below the bed.
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Feed Channel Backwash Storage
Aeration Grid
Polystyrene Media
Anoxic Zone
Sludge Removal
Figure 7.5 Biostyr® process arrangement for nitrification and denitrification (courtesy of Veolia). The top of the bed is restrained by a ceiling fitted with filtration nozzles to evenly collect the treated wastewater. The clean surface area of spherical beads is 1000 to 1400 m2/m3 (300 to 425 sq ft/cu ft). In the bottom of the reactor, influent is distributed by troughs formed in the base of the cells. The troughs are covered with plates, which have gaps at intervals to allow the flow to enter the cells and backwash wastewater to be collected. There is no need for a filter underdrain, as the media does not require support. Process air is distributed through diffusers located along the bottom of the reactor or within an aeration grid within the media bed; the latter are used if an anoxic zone is required for nitrogen removal. Only treated wastewater comes in contact with nozzles. Backwashing consists of countercurrent air-scour and backwash water flow. Solids are removed through the shortest pathway at the bottom of the reactor. Since 1993, when Toettrup et al. (1994) presented the first information regarding the process status, the Biostyr process has been applied widely in European wastewater markets and a number of installations in the United States. The Biobead BAF is very similar to Biostyr, except that the media is larger and heavier, using polypropylene or polyethylene with a specific gravity of approximately 0.95. Wastewater flow enters at the bottom of the reactor, and the flow is distributed by a grid or by a specially designed distribution system. For small cells, a simple system using one central feed and distribution plate arrangement is sufficient. For larger cells (larger than 5.5 m × 5.5 m [18 ft × 18 ft]), a more sophisticated arrangement is required, such as horizontal slots staggered across the cells. The slot size needs to be designed carefully to give even distribution, especially at low flows (Cantwell and Mosey, 1999).
Biological Filters
To prevent the media loss, a metal grid is fixed near the top of the reactor. Process air is supplied from a grid located below or within the media bed. By placing the process air grid within the media, it is possible to achieve some solids removal at the bottom of the bed. Because of the media specific gravity, it is relatively easy to release the accumulated suspended solids during backwashing, requiring only a relatively low head. Typically, backwashing consists of a combination of partial drainage, airscour, and countercurrent water flush. The dirty backwash water is removed over the outlet weir or through a bottom drain. Recovery of solids retention following backwash, until the bed is sufficiently packed again, may take some time, during which the effluent from the reactor may be recirculated through the plant. Upflow floating BAFs also may require a certain number of mini-backwashes (typically 4 to 8 and, in extreme cases, more than 10), to bump the filter, remove some solids, and lower the headloss to achieve a complete filtration cycle of 24 or 48 hours (time between two “normal” backwashes). The requirement for mini-backwashes plus normal backwashes can generate significant backwash wastewater. During demonstration testing in San Diego, California, of an upflow BAF with floating media for BOD removal, a floating media BAF generated a volume of backwash wastewater between 10.3 and 13.9% of the influent flow, compared with a sunken media BAF, which produced between 7.4 and 7.9% (Newman et al., 2005).
2.5 Moving-Bed, Continuous Backwash Filters Moving-bed, continuous backwash filters operate in an upflow mode and consist of media heavier than water that continuously moves downward, countercurrent to the wastewater flow. These filters are used widely for tertiary solids and turbidity removal, but also have been applied to separate-stage nitrification and denitrification. For nitrifying systems, air or pure oxygen is added; for denitrifying systems, a source of readily biodegradable carbon substrate, such as methanol, is added. Two commercially offered systems using this technology are Parkson DynaSand filters (Parkson Corporation, Fort Lauderdale, Florida) and Paques Astrasand filters (Balk, Netherlands). The filter cells are supplied as 4.65-m2 (50-sq ft) modules, with center airlift assembly. The effective media depth typically is 2 m, and sand media size generally ranges from approximately 1 to 1.6 mm. Moving-bed filters backwash continuously at a low rate; the treatment process is not interrupted by intermittant backwash cleaning cycles. A typical unit is shown in Figure 7.6. Influent wastewater enters the filter bed through radials located at the bottom of the filter. The flow moves up through the downward-moving sand bed, and effluent flows over a weir at the top of the filter. The media, with the accumulated
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Central reject compartment (H) Rejects (L)
Feed (influent) (A) Top of airlift pump (G) Filtrate weir (J)
Reject weir (K) Sand washer (L) Effluent (E)
Downward moving sand bed (D) Downward feed (B) Feed radials (C)
Bottom of airlift pump (F)
Figure 7.6 Schematic of moving-bed denitrification filter (courtesy of Parkson Corporation).
solids, is drawn downward to the bottom cone of the filter. Compressed air is introduced through an airlift tube extending to the conical bottom of the filter and rises upward with a velocity of greater than 3 m/s (10 ft/s), creating an airlift pump that lifts the sand at the bottom of the filter up the center column. The turbulent upward flow in the airlift provides scrubbing action that effectively separates solids from the media before discharging to the filter wash box. There is a constant upward flow of liquid into the wash box (backwash water) controlled by the wash box discharge weir. This discharge weir elevation is lower than the filtrate weir elevation, thus ensuring a constant flow over the backwash weir. Accordingly, the rate of backwash water remains reasonably constant and is independent of the filter forward flow. The media will drop to the surface of the filter, while
Biological Filters
the lighter suspended solids will be washed out with the backwash liquid. Movingbed filter manufacturers typically set the reject weir to provide a wash-water flowrate of approximately 54 to 65 m3/d (10 to 12 gpm) per filter module. This is the equivalent to a wash-water rate of approximately 10% of the forward flow at an average filter loading rate of 117 m3/m2·d (2 gpm/sq ft). The backwash frequency is quantified by the bed turnover rate. The media turnover rate is controlled by the airlift. If used for solids removal only, moving-bed filter media turnover rates range from 305 to 460 mm/h or 4 to 6 bed turnovers per day. To maintain sufficient biomass in the filter for denitrification, the bed turnover rate must be reduced to approximately 1 to 3 turnovers per day, or 100 to 250 mm/h.
2.6 Non-Backwashing, Open-Structure Media Filters These processes consist of submerged, static media to support the growth of biofilm for BOD removal, nitrification, or denitrification, but solids are intended to be carried through the reactor. This type of BAF is commonly referred to as a submerged aerated filter (SAF). If suspended-solids removal is required beyond adsorption and capture in the biofilm, it is carried out in a separate downstream process. A diagram of a typical simple SAF system is shown in Figure 7.7. The system arrangement depends on the supplier and duty of the SAF.
Treated wastewater effluent
Media
Influent wastewater
Aerators (diffusers or pipes)
Figure 7.7 Schematic of non-backwashing, open-structure media filter.
339
340
Biofilm Reactors
The media must be open in structure to prevent blockage by accumulated solids. Structured plastic media systems must include provisions for retaining the media. Mineral media has a high specific gravity and is unlikely to be dislodged in normal use. In the United Kingdom, blast furnace slag is readily available and is used for both carbonaceous and nitrification applications. Influent wastewater typically is introduced at the bottom of the reactor. Larger systems generally have more than one cell in series or have a more sophisticated flow-distribution system, similar to upflow BAFs, or use the air and the headloss to distribute the fluids when used in a downflow mode (Sulzer Biopur [Aker-Kvaerner, Stockton-on-Tees, United Kingdom]). Proper distribution of water and air has to be ensured to scour all parts of the media and prevent anaerobiosis in the outer zones (Cooper-Smith and Schofield, 2004; Frankl, 2004). For SAFs using mineral media, the air and influent distribution systems are combined with a floor system designed to support the heavier media. This system configuration is shown in Figure 7.8. Influent wastewater enters via a central channel in the base of the reactor, which is covered by plates. The plates are covered by rows of specially designed concrete underdrain blocks, fitted with interlocking plastic jackets. The media height typically is less than 4 m (13 ft.), but may be increased to reduce the
Figure 7.8 Mineral media upflow SAF with block underdrain (courtesy of Severn Trent).
Biological Filters
total cross-sectional area of the filters (Rogalla et al, 2005). For upflow SAF, treated effluent is discharged over a weir or trough at the top of the cell. For downflow SAF, piping or channels in the bottom collect the treated effluent. In addition to process needs, aeration is required to prevent media blockage. The air can be supplied either by a grid consisting of perforated distribution pipes, or by diffusers installed at the bottom of the SAF. Though improved air distribution may be achieved by using diffusers, some studies have shown that, in packed-beds, airbubble coalescence results in little to no oxygen-transfer advantage (Hodkinson et al., 1998) with diffusers. In some wastewater treatment plants (WWTPs), aeration scouring was insufficient to keep the media clean, resulting in deteriorating performance. Therefore, it is good design practice to maintain a backwash or backflush option, where air and water flows can be increased to scour the media periodically. Jet aeration, in which air is injected to a moving stream of water, typically via a venturi, and dispersed into fine bubbles, also has been coupled with a shallow SAF plant for small communities (Daude and Stephenson, 2004).
3.0 Media for Use in BIOLOGICALLY ACTIVE FILTERS Media is the heart of the BAF process, and there are several parameters that will influence the selection, including specific gravity, hardness, abrasion resistance, surface roughness, shape, granular size, uniformity coefficient, availability in large quantities within reasonable delivery times, and cost. Media selection is integral to treatment objectives, flow and backwashing regimes, and the specific process equipment manufacturer. The media acts simultaneously to support the growth of biomass and as a filtration medium to retain filtered solids. On the other hand, it has to be able to release the retained solids and biomass when backwashed, while being sufficiently robust to resist breakdown resulting from attrition during backwashing. Media can be categorized as mineral, structured plastic, or random plastic. In most cases, mineral media is denser than water, while plastic media is buoyant. Commercially available BAF systems and their media are listed in Table 7.1.
3.1 Mineral Media Mineral media are granular in form. There are various standards applied to natural media for use in sand filters (e.g., British Standards Institution, 1983). These apply
341
Table 7.1 Commercially available BAF reactor systems and media. Supplier
Flow regime
Media
Specific gravity
Size (mm)
Astrasand
Paques/ Siemens
Upflow, moving-bed
Sand
>2.5
1 to 1.6
Biobead
Brightwater F.L.I.
Upflow
Polyethylene
0.95
Biocarbone
OTV/Veolia
Downflow
Expanded shale
1.6
2 to 6
Biofor
Degremont
Upflow
Expanded clay
1.5 to 1.6
2.7, 3.5, and 4.5
Biolest
Stereau (Gayancourt, France)
Upflow
Pumice/ pouzzolane
1.2
Biopur
Sulzer/Aker Kvaerner
Downflow
Polyethylene
Biostyr
Kruger/ Veolia
Upflow
Polystyrene
0.04 to 0.05
2 to 6
1000 to 1400
Colox
Severn Trent
Upflow
Sand
2.6
2 to 3
656
Denite
Severn Trent
Downflow
Sand
2.6
2 to 3
656
Dynasand
Parkson
Upflow, moving-bed
Sand
2.6
1 to 1.6
Eliminite
FB Leopold
Downflow
Sand
2.6
2
SAF
Severn Trent
Up/down
Slag
2 to 2.5
28 to 40
Washed gravel
2.6
19 to 38
342
Process
Specific surface area (m2/m3)
1400 to 1600
Structured
240
Biological Filters
mainly to the finer media used in potable water filters, or in tertiary anoxic filters. The German DIN standard (Deutsches Institut für Normung, Berlin, Germany) covers larger-size gravels. Typically, mineral media have a specific gravity greater than 1, and sink in water. A floating medium made from foamed clay has been described, but has not been used in a full-scale application (Moore et al., 1999). Some expanded clays can have a density less than water, but tend to be brittle and absorb water over time. The media needs to maintain its structural integrity and form despite abrasion during backwashing. Most mineral media are resistant to chemical attack by constituents present in normal wastewaters. Particle size, shape, and size distribution must be considered carefully during design. The smaller the particle size, the larger the surface area available for growth of the biofilm, which then promotes higher levels of treatment. However, with smaller media, the flow channels between the particles will be reduced in size, causing increased headloss. If the particles are all of similar size, with an even particle-size distribution, the voids between the particles tend to be maximized, allowing a higher storage volume for accumulated solids and longer intervals between backwashes. With expanded clay (Lytag; York, United Kingdom) as the media in downflow pilot columns, smaller grain sizes generally provided higher levels of treatment and filtration efficiency, but at shorter filter cycle times (Kent et al., 2000; Smith et al., 1999). Often mineral media are graded, with coarser material near the bottom of the cell to prevent the fine medium from entering the underdrain. One upflow BAF system uses a multilayer concept with large media (40 mm size) at the bottom of the reactor and smaller media above this, ending in a fine media (2.5 mm size) for filtration at the top third of the bed (Brewer et al., 1997). The objective of this system is to avoid primary settlement, but only a few small plants with this media configuration are operating in France and Italy (Rogalla, 2004).
3.2 Random Plastic Media The BAF reactors with floating media typically use random media in the form of beads, with the key characteristic being grain size, which influences available surface area. The media material typically is either polypropylene or polystyrene, sometimes made from recycled plastic. The beads may be produced with an artificially roughened surface, to provide a better support for the micro-organisms. Because the beads are manufactured products, there is some control over properties, such as specific gravity. The lighter the beads, the more they compact at the top of the reactor and the greater the headloss and solids removal. However, beads
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Biofilm Reactors
with a specific gravity close to water (0.95) can become coated with biomass, increasing their specific gravity to the point of sinking. This leads to channeling and shortcircuiting and loss of media in the drain-down step of the backwash. The random media beads can be made in a range of sizes, with diameters from approximately 2 to 6 mm. Figure 7.9 shows the theoretical surface per cubic meter of carrier, depending on the grain size and the influence of various factors, such as the form factor, F, and the porosity, ε. The form factor is taken as 1 for ideal spherical particles and as 1.5 in practice for most materials. Epsilon is related to the density of the bulk media and the particle, as follows: Porosity, ε = 1–
Bulk density Particle density
(7.1)
For an ideal round particle of 3-mm grain size, and a porosity of 0.45, the specific surface area would be 1000 m2/m3.
3,000 Specific surface area (m2/m3)
344
F = 1 (ideal ball shaped carriers) F = 1.5 Simplified formula eps. = 0.3 Simplified formula eps. = 0.4
2,000
1,000
0
0
1
2
3
4
5
6
7
8
9
10
Effective carrier diameter da = (dfg*dcg)0.5 [mm]
Figure 7.9 Carrier diameter versus media specific surface area (Rother, 2005).
Biological Filters
With expanded media, such as polystyrene, care must be taken to select a particle density sufficient to withstand the compression during the filtration cycle. While there is no evidence of breakdown of the bead-type plastic medium, and the plastics are resistant to most chemicals present in domestic wastewater, they are vulnerable to organic solvents and petroleum spirit, which could be present in higher concentrations because of a particular industrial effluent.
3.3 Modular Plastic Media Modular plastic media are similar to those used in high-rate trickling filters, designed to provide a high void space, and free passage for flow and air. Modular media consists of flat and corrugated sheets bonded together to form rectangular modules or blocks, with the appearance of a honeycomb. There are two types of block—vertical flow (with all of the flow channels running top to bottom) and crossflow (flow channels at an angle to the vertical), with the latter typically used in BAF applications. The modular media can be made in a number of configurations to give various surface areas. Modular media with the largest surface area have the smallest flow channels and are used for tertiary treatment, particularly nitrification. However, the cost of the media is proportional to the surface area, and, for high surface area, random media may be more economical. As plastic media generally are lighter than water and are kept submerged within BAF (and SAF) reactors, the modules must be held down firmly by a structure at the top of the BAF cell, to prevent dislodging during operation and backwashing.
4.0 Backwashing and Air-Scouring Backwashing filters maximize capture and run times and guarantee proper effluent quality. Proper backwashing requires filter-bed expansion and rigorous scouring, followed by efficient rinsing. Poor filter cleaning will result in shortened filter runs, accumulation of solids, and deteriorating performance. Accumulation of solids and media (mudballing) produces short-circuiting of water flow and can result in excessive media loss. Feed characteristics and type of treatment provided by the BAF affect solids production and frequency of backwashing. For wastewaters with high suspended-solids concentrations, a significant portion of solids is removed by filtration. Inert solids will be retained within the media, until removed by backwashing, but biological solids may be degraded, depending on retention time. Inorganic salts of iron or aluminum,
345
346
Biofilm Reactors
which may be added to the influent for phosphorus removal, will form precipitates within the media bed and increase backwash frequency. Solids growth for tertiary BAF systems typically is low, so backwashing is relatively infrequent (one backwash per 36 to 48 hours). When solids content is low, foam caused by detergent in the wastewater may be a problem, because the scour aeration is concentrated in a small surface. Foam also can be an issue during process startup. Netting is recommended across the surface of the cells to keep the foam from blowing about the site. Reactor characteristics and media type influence backwash frequency. More openly structured media capture fewer solids. This reduces the backwash frequency, but the effluent wastewater may contain higher suspended-solids concentrations. Fine mineral media, such as silica sand, typically have the best solids retention characteristics, but tend to require frequent backwashing. Increasing flow velocity, particularly in upward-flow systems, may distribute the biomass accumulation more evenly through the reactor and prevent blockage and premature backwashing. The media beds in floating media filters, with a density close to that of water, have been shown to expand as solids accumulate, enabling them to hold solids without increasing headloss. Intense backwashing regimes have been developed to clean rapid gravity filters used in potable water treatment (Fitzpatrick, 2001). The bed typically is fluidized to allow the grains to separate and move freely and to remove as much accumulated material as possible. While backwashing regimes for BAFs were developed from these procedures, the removal of all of biological solids is not necessary or desirable, as the attached biomass is necessary for the proper functioning of the BAF process. Fluidization is avoided in BAFs; instead, the removal of excess biomass and accumulated solids is achieved during backwash by intense media contact and air-scouring in a slightly expanded media bed. Examples of backwashes and the influence of air and water velocities are illustrated in Figures 7.10 and 7.11 (IWA, 2006). In a small pilot column (diameter = 300 mm [12 in.] and media height = 1.8 m [5.9 ft], using a particle size of 4.5 mm with a bulk density of 0.83), backwash air velocities of 25, 32, and 40 m/h (1.4, 1.8, and 2.2 cfm/sq ft) were tested with a water flush of 8 m/h (3.3 gpm/sq ft). Only the two higher velocities generated sufficient turbulence for media backwash, as evidenced by the significant difference in backwash solids concentrations from the parallel columns. When an airflow rate of 36 m/h was applied and the water flush was varied between 5, 11, and 18 m/h, only the higher water flush velocities were sufficient to displace the solids within the 15 minutes duration of the backwash.
Biological Filters
Total solids content/mg.L–1
q = 10-12L/(s*m2) 7000
q = 8-10L/(s*m2)
q = 6-8L/(s*m2)
6000 5000 4000 3000 2000 1000 0
0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 Time of backwashing/min
Figure 7.10 Relation between total solids of backwash liquor and time for different backwash airflow rates (courtesy of IWA, 2006). Table 7.2 provides a comparison of typical BAF backwashing requirements, and the Descriptions of Biologically Active Filter Reactors and Equipment section described backwashing for each type of BAF configuration. Final backwashing requirements and duration typically are developed in collaboration with the BAF manufacturer. For example, the backwash sequence for an upflow sunken media BAF typically includes drain-down, air-scour, air- and water-scour (may include cycling between air-only and air/water scour), a water-only rinse, and filter-to-waste, when
Total solids content/mg.L–1
q = 4-6L/(s*m2) 7000
q=
q = 2-4L/(s*m2)
1-2L/(s*m2)
6000 5000 4000 3000 2000 1000 0
0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 Time of backwashing/min
Figure 7.11 Relationship between total solids of backwash liquor and time for different backwash water flowrates (courtesy of IWA, 2006).
347
Table 7.2 Summary of BAF backwashing requirements (Degremont, 2008; Kruger, 2008; Parkson, 2004; Slack, 2004). Total backwash water volume per cella
Total backwash wastewater volume per cellb
50
9.2 m3/m2 (225 gal/sq ft)
12 m3/m2 (293 gal/sq ft)
97 (5.3)
25
9.2 m3/m2 (225 gal/sq ft)
10 m3/m2 (245 gal/sq ft)
55 (22.5)
12 (0.65)
16
2.5 m3/m3 mediae (18.7 gal/cu ft media)
2.5 m3/m3 mediae (18.7 gal/cu ft media)
55 (22.5)
12 (0.65)
5
1.5 m3/m3 mediae (11.2 gal/cu ft media)
1.5 m3/m3 mediae (11.2 gal/cu ft media)
Downflow, sunken media
15 (6)
90 (5)
20 to 25
3.75 to 5 m3/m2 (90 to 120 gal/sq ft)
3.75 to 5 m3/m2 (90 to 120 gal/sq ft)
Upflow, movingbedf
0.5 to 0.6 (0.2 to 0.24)
Continuous through airlift
Continuous
55 to 67 m3/d (14 400 to 17 300 gpd)
55 to 67 m3/d (14 400 to 17 300 gpd)
Upflow, sunken media normal backwash Energetic backwashc Upflow, floating media normal backwash
Backwash water rate (m/h) (gpm/sq ft)
Air-scour rate (m/h) (scfm/sq ft)
Total duration (minutesa)
20 (8.2)
97 (5.3)
30 (12.3)
Mini-backwashd
Energetic backwash once every 1 to 2 months, depending on trend in “clean bed” headloss following normal backwash.
a
Mini-backwash applied as interim measure when pollutant load exceeds design load.
b
Backwash duration reflects total duration of the typical backwash cycle, which includes valve cycle time and pumping and nonpumping steps. The duration of each step is adjustable via programmable logic controller and supervisory control and data acquisition control systems.
c
Total backwash wastewater volume includes drain and filter-to-waste steps, where applicable.
d
Backwash volume requirements for upflow floating media BAFs typically are based on media volume rather than cell area, because depths vary.
e
Continuous backwash filter backwashing is based on a standard 4.65-m2 cell and a typical weir setting for reject flow of approximately 2.3 to 2.8 m3/h/cell (10 to 12 gpm/cu ft/cell).
f
Biological Filters
the backwashed cell initially is placed back in operation. Thus, backwash water is delivered only to the BAF cell for a portion of the total duration. Media, hydraulic, and organic loading rates and treatment objectives influence the frequency and duration of each step. Backwashing parameters are often adjusted during facility commissioning and long-term operation. The BAF backwash facilities and equipment typically include effluent clearwell; backwash water pumps (for co-current backwashing systems only); air-scour blowers; a backwash waste equalization tank; backwash waste-return pumps; and all of the automatic valves, instruments, and controls required for automatic initiation and sequencing of the backwash procedure. The equipment and facilities must be sized adequately to handle the air and water rates and volumes necessary for effective backwashing. During backwashing, the effluent flow from the BAF will be reduced, so care must be taken that no effluent violations occur if an existing treatment train operating in parallel with a BAF system relies on blending the two effluents. In addition, this flow variation of BAF system effluent must be accounted for in the design and operation of any downstream treatment process, such as UV disinfection. In multiple-stage BAF systems, where there could be several different BAF cell sizes and different backwash water- and air-scour requirements for each cell size, backwash facilities and equipment sizing typically is based on requirements for the largest cell, to avoid separate sets of backwash equipment for each BAF stage. Final effluent taken directly from the effluent channel of the last stage of a multistage BAF system or via a final clearwell typically is used for backwash water. If an interstage clearwell is used, the ability to maintain minimum flow to the downstream BAF cells must be considered in the clearwell design.
5.0 BIOLOGICALLY ACTIVE FILTER Process Design Several factors influence the process design for BAF systems. As discussed earlier, mass-transfer limitations into the biofilm often limit substrate removal performance, and the media-specific surface area available for biofilm attachment and substrate flux affect the biofilm reactor design. Several physical conditions within BAF systems also significantly affect performance, including oxygen availability and airflow velocity, filtration velocity, media-packing density, and backwash efficiency. These factors all affect external mass transfer and, indirectly, penetration into the biofilm. Because of the importance of these parameters and perhaps because of uncertainty of the
349
350
Biofilm Reactors
actual media-specific surface area, BAF performance results typically are expressed as a function of substrate volumetric loading rates, rather than surface area. Deterministic modeling of BAFs based on kinetic expressions is complicated, as biofilms are complex and highly dynamic structures. Therefore, uncertainty of prediction persists because of the many variables involved, which affect the degree of soluble and particulate substrate diffusion, rate of biomass growth, biofilm density, and type and quantity of microorganisms in the biofilm. The filtering capability of BAFs makes the already difficult task of quantifying the degree of particle hydrolysis even more important in BAFs than in activated sludge systems or other biological processes with downstream solids separation. However, although continued development and calibration of biofilm models is needed, these models provide an excellent tool for evaluation and development of more tailored designs. Parameters governing treatment capacity of BAF are as follows: • Substrate loading (volumetric loading rates, in terms of kg BOD/m3∙d or kg N/m3∙d), which will determine the media volume. Guidelines for design loading rates have been compiled from the literature and are a function of wastewater characterization, substrate flux, temperature, and physical conditions in the biofilm reactor, as discussed above. Design guidance is provided based on flow regime and typical media and backwashing practices in use for different BAF reactors. • Filtration rate, or total volume of wastewater applied per area of media per unit time (m3/m2∙d), also is used to determine the filter surface area. Filtration velocity affects system headloss, solids capture, and air and water distribution within the media, diffusion, and detention time. • Solids holding capacity, which will determine the backwash frequency. Design requirements will vary, depending on the wastewater characteristics, wastewater temperature, required effluent quality, and the media specific surface area provided. Variations in these parameters tend to limit the ability to tailor a design when using “typical” volumetric loading rates and filtration rate guidelines. Therefore, process designs typically are developed through a combination of loading rate criteria, manufacturer’s proprietary models, and, increasingly, process simulation models. This section focuses on basic BAF process design considerations for secondary treatment (BOD removal), nitrification, combined nitrification and denitrification, post-denitrification, and phosphorus removal considerations. In all cases, careful
Biological Filters
attention must be paid to the influent wastewater characteristics and variations entering the BAF processes and the influence of upstream processes. To avoid repetition within these process design subsections, additional key components, such as the aeration system design and the need for supplemental carbon substrate, are addressed within the Design Considerations section, later in this chapter.
5.1 Process Design for Secondary Treatment This section reviews criteria for BAFs designed for secondary treatment (e.g., carbon oxidation [BOD removal] and suspended solids removal). Design guidelines for nitrifying and denitrifying BAFs are covered in later sections of this chapter.
5.1.1 Volumetric Biochemical Oxygen Demand Loading Volumetric BOD loading rates vary widely in the literature for upflow BAFs designed for secondary treatment, ranging from 1.5 to 6 kg/m3·d (94 to 375 lb/d/1000 cu ft). Generally, BAFs designed for secondary treatment have BOD and TSS loading rates that are 2 to 3 times higher than tertiary BAFs. Information regarding aeration design is provided under the Design Considerations section later in this chapter.
5.1.2 Hydraulic Loading Average and peak hydraulic loading rates for secondary treatment systems typically range from 4 to 7 m/h (1.5 to 3.0 gpm/sq ft) and 10 to 20 m/h (4 to 8 gpm/sq ft), respectively. Because BAFs for secondary treatment typically are placed immediately downstream of primary clarification, the applied volumetric mass loading rate is almost always the limiting design parameter for these systems. Typical design volumetric mass and hydraulic loading rates for BAFs designed for secondary treatment are provided in Table 7.3. For simultaneous secondary treatment and nitrification, the carbon loading at lower temperatures needs to be limited to less than 2.5 kg BOD/m3·d (156 lb/d/1000 cu ft) (Rogalla et al, 1990). Simultaneously, a total Kjeldahl nitrogen loading removal rate of 0.4 kg N/m3·d can be obtained.
5.1.3 Backwashing The backwash frequency for BAFs designed for secondary treatment is related to the applied organic and TSS load, degree of particle hydrolysis taking place within the media, biomass yield, and solids retention capacity of the media. Because of the higher biomass yield of heterotrophic bacteria and higher applied TSS loadings, BAFs for secondary treatment (BOD removal) need to be backwashed at least once
351
352
Biofilm Reactors
Table 7.3 Typical BAF loading rates for secondary treatment.* Applied volumetric loading (kg/m3∙d) (lb/d/1000 cu ft)
Hydraulic loading (m3/m2∙h) (gpm/ cu ft)
Upflow sunken or floating media, backwashing (Degremont, 2007; Kruger, 2008)
BOD = 1.5 to 6 (94 to 370) TSS = 0.8 to 3.5 (50 to 220)
3 to 16 (1.2 to 6.6)
BOD = 65 to 90% TSS = 65 to 90%
Upflow, sunken media (German Association for Water, Wastewater and Waste, 1997)
10
Upflow, floating media (German Association for Water, Wastewater and Waste, 1997)
8
Submerged, nonbackwashing (McCarty, 2008)
BOD = 0.8 to 1.5 (50 to 94) at 20°C
2 to 12 (0.8 to 5) at 20°C
BOD = 85 to 95%
Type of BAF
Removal efficiency (%)
* Design loading rates depend on specific wastewater characteristics and level of treatment required.
per day. More frequent backwashing results in less hydrolysis of particulate BOD, which, in turn, results in lower oxygen demand and higher backwash waste solids quantities. Phipps and Love (2001) calculated biomass observed yields in the range 0.43 to 0.48 mg biomass as COD generated per mg substrate COD consumed in a fullscale Biofor (Degremont) treating conventional primary clarifier effluent for carbon removal (backwash frequency of once per day). They also determined that 40 to 46% of applied particles underwent hydrolysis. The main limitation of any BAF remains the solids storage capacity. The volume that can be accumulated between backwashes is between 2.5 and 4 kg TSS/m3·d (Degremont, 2007). The BAF backwash water typically contains 500 to 1500 mg/L suspended solids, but this varies with the type of treatment being provided, cycle time, and quantity of backwash water used. The BOD removal produces biomass from growth of the
Biological Filters
microorganisms, which convert degradable material into new cells, carbon dioxide, and water—similar to that of activated sludge. Sludge production generally is 0.7 to 1 kg solids/kg BOD removed. In a two-stage SAF/BAF in Aberdeen (Scotland, United Kingdom), the German ATV equation to predict solids from activated sludge was applied successfully (Jolly, 2004; German Association for Water, Wastewater and Waste [ATV-DVWK], 2000).
5.1.4 Design Example: Design of a Submerged, Upflow Biological Aerated Filter System for Secondary Treatment (No Nitrification) Determine the total volume of BAF media, total BAF reactor filtration area, and number of BAF cells required to achieve BOD5 and TSS removal efficiencies (EBOD and ETSS) of at least 90% when treating domestic wastewater. Determine the BAF backwash wastewater volume and solids concentration. Assume that the following conditions apply for this example:
(1) Influent (including returns) maximum monthly flowrate = Q0 = 3950 m3/h (25 mgd) (2) Influent (including returns) flow peaking factor (P.F.) = 2.8 (3) BOD5t after primary settling = CBOD5t = 220 mg/L (4) TSS after primary settling = CTSS = 150 mg/L (5) BAF media height = HM = 4 m (13.1 ft) (6) BAF effluent used as backwash water (6) BAF backwash return flow equalized and combined with other return flows at head of plant
5.1.5 Solution The solution can be derived as follows: (1) Calculate the BOD5t and TSS load to the BAF system. BOD5t load = 24(Q0)(BOD5t)/1000 = 24(3950)(220)/1000 = 20 856 kg/d (45 883 lb/d) TSS load = 24(Q0)(TSS)/1000 = 24(3950)(150)/1000 = 14 220 kg/d (31 284 lb/d) (2) Assume maximum volumetric applied loading rates. BOD5t = 3 kg/m3·d (185 lb/d/1000 cu ft) for 90% removal efficiency TSS = 1.6 kg/m3·d (99 lb/d/1000 cu ft) for 90% removal efficiency
353
354
Biofilm Reactors
(3)
Calculate the total BAF media volume (VM) required. V1BOD = 20 866/3 = 6955 m3 (245 512 cu ft) V2TSS = 14 220/1.6 = 8888 m3 (313 746 cu ft) Based on these calculations, the TSS load limits the BAF capacity and governs the sizing.
(4) Calculate the total BAF filtration area (A) required based on volumetric loading. Avol = V/HM = 8888/4 = 2222 m2 (23 909 sq ft)
(5) Calculate the total BAF filtration area based on a maximum hydraulic loading rate of 20 m/h. Ahyd = (3950)(P.F.)/20 = (3950)(2.8)/20 = 553 m2 << 2222 m2, Avol is limiting
(6) Select the standard cell size, Acell. Assume 144-m2 (1550-sq ft) standard cell sizes provided by BAF manufacturers.
(7) Calculate the number of BAF cells required assuming 1 backwash per cell per 24 hours. n = 2222/144 = 15.4 N = 15.4 + 15.4/24 = 16 BAF cells Note: Depending on the initial capacity needs compared with the design capacity, the designer should consider incorporating a redundant BAF unit for reliability and ease of maintenance. (8) Check the BAF media solids retention capacity. Assume 2.5 kg/m3·cycle solids retention capacity Total media retention capacity = (2.5)(16)(144)(4) = 23 040 kg/cycle Biomass yield = Y = 0.7 – 1 kg TSS/kg BOD removed. Assume Y = 1.0; Solids production = (Y)(BOD5t load)(EBOD) = (1.0)(20 856)(0.90) = 18 770 kg/d = 782 kg/h; Backwash frequency = (23 040)/(782) = 29 hours.
(9) Check maximum hydraulic loading rate with one cell in backwash and one cell out-of-service. (Q0)(P.F.)/(N–2)(Acell)= (3950)(2.8)/(16 – 2)(144) = 5.5 m/h <20 m/h (10) Calculate BAF backwash wastewater volume and solids concentration based on one backwash per day.
Biological Filters
Volume of backwash wastewater produced per media volume, VolBW (from Table 7.2) Assume VolBW = 3 m3/m3 media Volume of backwash wastewater produced per backwash, VBW VBW = (VolBW)(HM)(Acell) = (3)(4)(144) = 1728 m3 Backwash wastewater solids concentration, CBW CBW = (Y)(BOD5t load)(EBOD)/(N)(VBW) CBW = (1.0)(20 856)(0.9)(1/16)(1/1728) = 679 kg/m3 = 679 mg/L This set of calculations represents an initial estimate of the BAF facility s izing. Development of the final design typically is an iterative process between the design engineer and the process equipment manufacturers being considered. Refinements typically are made by incorporating a combination of the manufacturer’s experience and more detailed process modeling results.
5.2 Process Design for Nitrification Temperature, effluent requirements, fluid velocities (air and water), and loading influence nitrification capacity. As discussed in the Descriptions of Biologically Active Filter Reactors and Equipment section, biochemical transformation processes occurring in the biofilm are dependent on substrates diffusing in and out of the biofilm. Reaction rates or level of treatment achieved are defined by the rate-limited substrate. Bulk-phase ammonia, alkalinity, oxygen, and COD concentrations affect nitrification. As COD loadings increase, oxygen tends to become the rate-limited substrate. Competition for oxygen intensifies, and the heterotrophic respiration at the outer layers of the biofilm will reduce the availability of oxygen for nitrification in the deeper layers (Wanner and Gujer, 1985). Rogalla et al. (1990) found that nitrification tends to decrease when biodegradable COD loadings approach 4 kg/m3∙d. The influence of C:N on nitrification is illustrated in Figure 7.12 (Rother, 2005).
5.2.1 Influence of Hydraulic Filtration Rates In BAFs, increasing fluid velocities serve to increase the external mass transfer, leading to higher nitrification rates (Tschui et al., 1993). The results are summarized in Table 7.4, which shows a particularly noticeable effect with upflow floating media. Under constant volumetric loading rates of 1.3 to 1.4 kg NH3-N/m3∙d and 0.65 ± 0.2 kg CBOD5/m3∙d, Husovitz et al. (1999) observed a 17% increase in ammonia mass
355
1.0 0.9 0.8
rV,NH4-N [kg NH4-N/(m³*d)]
0.7 0.6 0.5 0.4
tCOD/NH 4-N
0.3 0.2
BOD 5/NH4-N
0.1 0.0 0
2
4
6
8
10 12 14 C/N in feed water
16
18
20
22
COD/N raw wastewater
COD/N C-reduced
COD/N floc. raw wastewater
BOD/N raw wastewater
BOD/N C-reduced
BOD/N floc. raw wastewater
Figure 7.12 Nitrification rate for differently pretreated raw wastewaters as a function of C:N (temperature adjusted to 12°C; expanded clay; fluid velocity, vw = 8 to 8.5 m/h; and air velocity, vG = 20 m/h) (Rother, 2005).
Table 7.4 Effect of water velocity on nitrification rate for three types of biofilters (results taken from Tschui et al., 1993). Type of system
Water velocity (m/h)
Nitrification rate (g N/m3·d)
Upflow, floating medium
4 6 8
1300 1650 1700
Downflow, sunken medium
2 4 6
650 750 850
Upflow, modular plastic medium
6 8 10 12 14
200 250 300 380 400*
* For this filter, there was no increase in rate beyond 400 g N/m3·d.
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removal (up to 1.26 kg NH3-N/m3∙d) as the hydraulic loading rate was increased from 5.1 to 15.8 m/h. If temperature, effluent quality, and removal efficiency are not the limiting factors, then ammonia removal of 80 to 90% at ammonia loads between 2.5 and 2.9 kg/m3∙d can be achieved (Peladan, et al., 1996, 1997). On a full-scale demonstration cell (surface 144 m2) at 22°C, 91% removal of NH3-N at loadings up to 2.3 kg/m3∙d was observed (Pujol et al., 1994). Tests at the VEAS plant in Oslo, Norway, on expanded clay showed the influence of grain size and shape. Table 7.5 summarizes the average results for nitrification rates with crushed and round media. It can be seen that, with 30% of crushed media, the larger grain size is as effective as the smaller one, with only regularshaped beads, and nitrification efficiency rises with the degree of crushed media. A summary of typical loading rates for nitrification applications is provided in Table 7.6.
5.2.2 Effect of Process Air Velocity Similar to water velocity, the increase of process air velocity has been shown to improve nitrification rates (Tschui et al., 1993, 1994), as higher process air velocities increase turbulence and increases external mass transfer, indirectly allowing better penetration of substrate into the depth of the biofilm.
Table 7.5 Influence of grain size and shape on performance from testing in Oslo, Norway (Filtralite, 2008). Parameter
Media 1
Media 2
Media 3
Media 4
Media 5
Media 6
Degree of crushed material (%)
0
0
30
30
60
100
Grading: effective media size (mm)
5.74
3.47
5.29
3.49
3.04
3.05
Relative average nitrification rate as a percentage of the observed rate of media 6
48%
66%
65%
80%
94%
100%
357
Table 7.6 Typical BAF loading rates for nitrification.* Applied volumetric loading (kg/m3∙d) (lb/d/1000 cu ft)
Hydraulic loading (m3/m2∙h) (gpm/cu ft)
Removal efficiency (%)
Upflow, sunken, or floating media, backwashing (Degremont, 2007; Kruger, 2008) following primary treatment
BOD = <1.5 to 3 (<94 to 188)
3 to 12 (1.2 to 5)
BOD = 70 to 90%
Upflow, sunken, or floating media, backwashing (Degremont, 2007; Kruger, 2008) following secondary treatment
BOD = <1 to 2 (<62 to 125)
Type of BAF
TSS = <1.0 to 1.6 (<62 to 100)
TSS = 65 to 85%
NH3-N = <0.4 to 0.6 (<31 to 62) at 10°C; <1.0 to 1.6 (< 62 to 100) at 20 °C
NH3-N = 65 to 75%
3 to 20 (1.2 to 8.2)
BOD = 40 to 75%
358
TSS = <1.0 to 1.6 (<62 to 100)
TSS = 40 to 75%
NH3-N = <0.5 to 1.0 (<31 to 62) at 10°C; <1.0 to 1.6 (<62 to 100) at 20°C
NH3-N = 75 to 95%
Upflow, floating media, backwashing (German Association for Water, Wastewater and Waste, 1997) following secondary treatment
NH3-N = 1.5 (94)
Upflow, sunken media, backwashing (German Association for Water, Wastewater and Waste, 1997) following secondary treatment
NH3-N = 1.2 (75)
Submerged, non-backwashing (McCarty, 2008) following secondary treatment
NH3-N = 0.2 to 0.9 (12 to 56) at 20°C
2 to 12 (0.8 to 5) at 20°C
NH3-N = 85 to 95%
* The design loading rates depend on specific wastewater characteristics, upstream treatment processes, and level of treatment required.
Biological Filters
5.2.3 Dependence on Loading Conditions When BAFs are operated for significant periods under reduced ammonia-loading conditions, the inventory of biomass also will decrease. An example of this was shown in testing by Tschui et al. (1994), where volumetric ammonia removal rates decreased by 30% after a transition from operation under non-NH3-N-limiting conditions to lower volumetric NH3-N loading rates. This is an important consideration for separate-stage nitrification BAF applications, in which some nitrification can occur in the main secondary plant during the summer months. The BAF will need to be able to treat higher ammonia loads when the temperature drops and upstream nitrification decreases. Nitrification performance also depends on long-term loading of the reactor. In the media, excess biomass accumulates, which can be available when peaks are applied, because they typically are associated with higher velocities or concentrations, which allow deeper penetration of the substrates into the biofilm. Figure 7.13 illustrates that the maximum instantaneous removal rate can be twice the average applied load, up to the maximum capacity of the reactor.
1.2
instantaneous removed load [kg NH4-N/(m3*d)]
1
0.8 0.6 Nicolavcic (2002) Le Tallec et al. (1997)
0.4
0.2 0
0
0.1
0.2
0.3
0.4
daily applied load [kg NH4
0.5
0.6
0.7
0.8
-N/(m3*d)]
Figure 7.13 Maximum nitrification versus long-term average loading (from Rother [2005], adapted from Le Tallec et al. [1997]and Nicolavcic [2002]).
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5.2.4 Temperature Effects Lower operating temperatures have a significant effect on nitrification. Comparing three types of tertiary nitrifying BAFs (downflow mineral media and upflow plastic media—either floating or modular), the following long-term temperature dependency of nitrification was established (Tschui et al., 1994):
rV , NH4 − N (T ) = rV , NH
4
− N ( T = 10 o C )
* expkT (T −10)
(7.2)
Where rV , NH
rV , NH4 − N (T ) = volumetric nitrification rate at temperature, T (°C); o = volumetric nitrification rate at T = 10°C; and 4 − N ( T = 10 C ) kT = temperature coefficient (Arrhenius factor) = 0.03/°C.
These authors found that the temperature coefficient was 0.03/°C for all three types of media studied (corresponding to an Arrhenius factor of 1.04). Nitrifying organisms have a relatively low solids yield (approximately 0.05 kg/kg N removed [Downing, Painter, and Knowles, 1964; Downing, Tomlinson, and Truesdale, 1964]). Most of the sludge produced by tertiary BAFs is derived from the suspended solids removed by filtration. A portion of these solids undergo hydrolysis and heterotrophic degradation. The net sludge production is approximately 0.5 to 0.8 kg/kg solids removed.
5.2.5 Design Example: Design of a Submerged, Upflow Biological Aerated Filter System for Nitrification Following Secondary Treatment Determine the total volume of BAF media, total BAF reactor filtration area, and number of BAF cells required to achieve a nitrification efficiency (EN-NH4) of at least 85% year-round, when treating secondary effluent. Determine the BAF backwash wastewater volume and solids concentration. Assume that the following conditions apply for this example:
(1) Influent (including returns) maximum monthly flowrate = Q0 = 3950 m3/h (2) Influent (including returns) flow peaking factor = P.F. = 2.8 (3) Influent minimum wastewater temperature = 12°C (4) Secondary effluent BOD5t = CBOD5t = 25 mg/L (5) Secondary effluent TSS = CTSS = 30 mg/L (6) Secondary effluent N-NH3 = CN-NH3 = 38 mg/L (7) BAF media height = HM = 3 m (8) BAF effluent used as backwash water (9) BAF backwash return flow equalized and combined with other return flows at head of plant
Biological Filters
5.2.6 Solution The solution can be derived as follows: (1) Calculate BOD5t , TSS, and ammonia loads to the BAF system. BOD5t load = 24(Q0)(BOD5t)/1000 = 24(3950)(25)/1000 = 2370 kg/d TSS load = 24(Q0)(TSS)/1000 = 24(3950)(30)/1000 = 2844 kg/d Ammonia load = 24(Q0)(N-NH3)/1000 = 24(3950)(38)/1000 = 3602 kg/d (2) Assume volumetric applied loading rates. BOD5t = 3 kg/m3·d TSS = 1.6 kg/m3·d N-NH4 = 0.8 kg/m3·d (3)
Calculate the total BAF media volume (VM) required. V1BOD = 2370/3 = 790 m3 V2TSS = 2844/1.6 = 1778 m3 V3N-NH3 = 3602/0.8 = 4503 m3
(4) Calculate the total BAF filtration area (A) required based on volumetric loading. Avol. = V/HM = 4503/3 = 1501 m2
(5) Calculate the total BAF filtration area based on maximum hydraulic loading rate of 20 m/h. Ahyd. = (3950)(P.F.)/20 = (3950)(2.8)/20 = 553 m2 << 1501 m2, Avol is limiting
(6) Select standard cell size, Acell. Assume 100-m2 standard cell sizes provided by BAF manufacturers.
(7) Calculate the number of BAF cells required assuming 1 backwash per cell per 48 hours (see Table 7.3). n = 1501/100 = 15 N = 15 + 15/48 => 16 BAF cells Note: Depending on the initial capacity needs compared with the design capacity, the designer should consider incorporating a redundant BAF unit for reliability and ease of maintenance.
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Biofilm Reactors
(8) Check BAF media solids retention capacity. Assume 2.5 kg/m3·cycle solids retention capacity Total media retention capacity = (2.5)(16)(100)(3) = 12 000 kg/cycle Biomass yield = Y = 0.5 – 0.8 kg TSS/kg BOD removed. Assume Y = 0.8 Solids production = (Y)(BOD5t load) = (0.8)(2370) = 1896 kg/d = 79 kg/h Backwash frequency = (12 000)/(79) = 152 hours; 48 hours maximum
(9) Check maximum hydraulic loading rate with one cell in backwash and one cell out-of-service. (Q0)(P.F.)/(N–2)(Acell)= (3950)(2.8)/(16 – 2)(100) = 7.9 m/h < 20 m/h (10) Calculate the BAF backwash wastewater volume and solids concentration. Volume of backwash wastewater produced per media volume, VolBW (from Table 7.2) Assume VolBW = 3 m3/m3 media Volume of backwash wastewater produced per backwash, VBW VBW = (VolBW)(HM)(Acell) = (3)(3)(100) = 900 m3 Backwash wastewater solids concentration, CBW CBW = (Y)(BOD5t load)/(N)(VBW) CBW = (0.8)(2370)(1/16)(1/900) = 0.132 kg/m3 = 132 mg/L This set of calculations represents an initial estimate of the BAF facility sizing. Development of the final design typically is an iterative process between the design engineer and the process equipment manufacturers being considered. Refinements typically are made by incorporating a combination of the manufacturer’s experience and more detailed process modeling results.
5.3 Process Design for Combined Nitrification and Denitrification Nitrogen removal can be accomplished either by oxidizing the ammonia in a first stage followed by reducing the nitrate in a second stage, where an external carbon source is added (referred to as post-denitrification); or by recycling the nitrified effluent to a denitrification stage before nitrification (pre-denitrification). In pre-denitrification, the nitrified effluent is recycled to an anoxic reactor located upstream of the nitrifying reactor. In some upflow, floating-media BAF configurations, a portion of the nitrified effluent may be recycled to an anoxic zone in the bottom part of the media (Payraudeau and Tallec, 2003). Schematics of these pre-denitrification options are shown in Figure 7.14.
Biological Filters
(a)
(b)
FI C-source
FI C-source
N
FI C-source DN
N
DN
DN
recycle
O2
recycle
O2
Figure 7.14 BAF configurations for combined nitrification and denitrification: (a) pre-denitrification/nitrification and optional post-denitrification, and (b) pre-denitrification/nitrification in one reactor.
It is noted that reaction rates will be affected by the combination of biomass in a combined nitrification and denitrification reactor. Figure 7.15 (Rother, 2005) compares nitrification rates in a filter specialized in nitrification with the aerobic zone of a combined system. For this particular media (expanded clay), nitrification could reach 1 kg N/m3·d for a separate reactor, but was limited to approximately 0.6 kg N/m3·d for a combined system.
1.2
rV,NH4-N [kg/(m3N*d)]
1.0 0.8 0.6 0.4 Combined system
0.2
Separated system 0.0 0.0
0.4
0.8 loading rate [kg NH4
1.2
1.6
-N/(m3
N*d)]
Figure 7.15 Nitrification with flocculated raw wastewater (T = 17 to 22°C, BOD5/NH4-N = 2, vw = 8.5 m/h, vG = 20 m/h, and no temperature adjustment) (Rother, 2005).
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364
Biofilm Reactors
Another limitation of the combined nitrification and denitrification is the recycle rate, which, with nitrates, also contains dissolved oxygen. A part of the anoxic zone will be used to consume this oxygen, also depleting available carbon for denitrification in the process. This limits the denitrification rates to approximately 1 kg N/ m3·d (Ninassi et al., 1998), when, with external carbon addition, a denitrification rate up to 3 times higher can be achieved. The economic optimum of the internal recycle seems to be approximately 200% (Karschunke and Sieker, 1997), which corresponds to nitrogen removal of approximately 70%, once nitrogen uptake into the biomass is considered. If sufficient carbon is available and the anoxic zone is large enough, then nitrogen removal is proportional to the recycle, but with a diminishing return. Recycle typically is limited to ratios of 3 or total nitrogen removals of 75% because of the excessive hydraulic load and oxygen recycle. The minimum nitrate concentration achievable, assuming complete nitrification and denitrification and ignoring nitrogen uptake because of cell synthesis, is given by the following:
no 3 − n eFF =
nH 3 − n InF R+ 1
(7.3)
Where NH3-NINF = ammonia concentration in influent, mg/L; NO3-NEFF = nitrate concentration in effluent, mg/L; and R (= QR/Qin) = recirculation ratio. Recycling treated wastewater has the advantage of increasing upflow velocity in both pre-denitrification and nitrification reactors, which increases the reaction rate. Recycling of the treated wastewater has the advantage that it increases the upflow velocity in both the pre-denitrification and nitrification reactors, which increases the reaction rate. Ryhiner et al. (1992) tested a pre-denitrification configuration using submerged structured media BAFs with a final polishing filter to ensure low nitrogen and suspended solids. Approximately 60 to 70% nitrate-nitrogen (NO3-N) removal was achieved in the pre-denitrification reactor at loading rates ranging from 0.1 to 0.6 kg NO3-N/m3∙d. During a 1-year study, Pujol and Tarallo (2000) achieved approximately 68% NO3-N removal (or 0.9 kg NO3-N/m3∙d), with typical wastewater feed to the pre-denitrification BAF and up to 90% removal when additional substrate (methanol) was added. A range of recycle rates (150 to 350%) and corresponding filtration velocities (9.4 to 16.9 m3/m2∙h) also were tested. Design guidance for predenitrification BAF systems is provided in Table 7.7.
Biological Filters
Table 7.7 Typical biologically active filter loading rates (BAF) for pre-denitrification.*
Type of BAF
Applied volumetric loading (kg/m3∙d) (lb/d/1000 cu ft)
Hydraulic l oading (m3/m2∙h) (gpm/cu ft)
Removal efficiency (%)
Upflow, sunken media (Degremont, 2007) separate BAF stages (pre-denitrification + nitrification)
N-NO3 = 1 to 1.2 (62 to 75)
10 to 30 (4 to 12)
NO3-N = 75 to 85%
Upflow, floating media (Ninassi et al., 1998) combined anoxic/ aerated BAF stage
1 to 1.2 (62 to 75)
12 to 21.5 (4.9 to 8.8)
NO3-N = 70% without supplemental carbon; 85% with supplemental carbon
* Design and performance are dependent on wastewater characteristics, upstream treatment processes, effluent goals, and readily biodegradable carbon substrate.
5.4 Process Design for Tertiary Denitrification This section focuses on basic design criteria for post-denitrification BAFs, including hydraulic loading, volumetric mass loading, half-order nitrogen removal kinetics, solids loading, temperature, supplemental carbon requirements, and nitrogen-release cycle requirements.
5.4.1 Volumetric Mass Loading The volumetric loading (and removal) rates vary widely, with citations ranging from 0.2 to 4.8 kg/m3 (15 to 300 lb/1000 cu ft) (Degremont, 2007; Metcalf & Eddy, 2003; U.S. EPA, 1993; WEF, 1998). Many post-denitrification filters are preceded by an activated sludge biological nutrient removal process. Because some denitrification is achieved upstream, filter influent NOx-N concentrations typically are less than 10 mg/L. In these cases, hydraulic considerations generally govern the post-denitrification design, and many
365
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Biofilm Reactors
installations are operating at mass loading rates of approximately 0.3 to 0.6 kg/m3·d (20 to 40 lb/d/1000 cu ft). The range of loading rates for upflow post-denitrification BAF reactors tends to be higher than for post-denitrification sand filters, because they are not designed for the same level of TSS removal and are not as limited by hydraulics. A summary of typical volumetric and hydraulic loading criteria for different types of denitrification filters is provided in Table 7.8. Table 7.8 Typical biologically active filters (BAF) loading rates for post-denitrification.*
Type of BAF
Applied volumetric loading (kg/m3∙d) (lb/d/1000 cu ft)
Hydraulic loading (m3/m2∙h) (gpm/cu ft)
Removal efficiency (%)
Downflow, sunken media (Slack, 2004; U.S. EPA, 1993)
NO3-N = 0.3 to 3.2 (20 to 200)
4.8 to 8.4 (2 to 3.5) average 12 to 18 (5 to 7.5) peak
NO3-N = 75 to 95%
Upflow, sunken media (Degremont, 2007)
NO3-N = 0.8 to 5 (50 to 300)
10 to 35 (4 to14)
NO3-N = 75 to 95%
Upflow, sunken media (German Association for Water, Wastewater and Waste, 1997)
2 (125)
Upflow, floating media (German Association for Water, Wastewater and Waste, 1997)
1.2 to 1.5 (75 to 94)
Moving bed, continuous backwash (deBarbadillo et al., 2005)
NO3-N = 0.3 to 2 (20 to 120)
4.8 to 5.6 (2 to 4) average 13.4 (6) peak
NO3-N = 75 to 95%
* Selected loading rates are dependent on treatment objectives, upstream processes, wastewater characteristics, and carbon source.
Biological Filters
5.4.2 Half-Order Kinetic Model Harremoes (1976) suggested that denitrification filter kinetics are dependent on the diffusion of substrate into pores in the biofilm. Zero-order heterogeneous reactions in a pore that is penetrated only partially by the substrate result in bulk half-order reaction kinetics for nitrate concentration. The half-order reaction kinetics result in the following expressions for a plug-flow system (Harremoes, 1976; Hultman et al, 1994): rdn = or
dS = − k 1 . SB,no3 -n dt 2
SB,no3 -n − Sin,no3 -n
(7.4)
1 = − . k1 . t 2 2
Where rDN = denitrification rate per unit volume of filter (mg/L∙min or g/m3∙min), SB,NO3-N = bulk-liquid nitrate-nitrogen concentration (effluent) (mg/L), Sin,NO3-N = nitrate-nitrogen concentration in the influent stream (mg/L), k 1 = half-order reaction coefficient per unit volume of filter 2
[(mg/L)1/2∙min, or (g/m3)1/2·min], h = filter empty bed retention time (minutes), qA h = filter media height (m), and qA = filter surface hydraulic loading rate (m3/m2∙min). t=
Substituting ST = 0.87 mg/L dissolved oxygen + 2.47 mg/L NO3-N + 1.53 mg/L NO2-N for SB and Sin results in the following (McCarty et al., 1969):
or
SB,t − Sin,t = − SB,t Sin,t
= 1−
1 h . k1 . 2 2 qA a. h Sin,t
(7.5)
(7.6)
Where a=
k . . 2 qA
Profile sampling conducted at different media depths by Harremoes (1976), Hultman et al. (1994), and Janning et al. (1995) has shown that the half-order kinetic model reasonably correlates to observed values. Estimated half-order kinetic
367
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Biofilm Reactors
constants for several post-denitrification BAFs ranging from 0.09 to 0.9 are summarized in deBarbadillo et al (2005). Because of relatively limited information on halforder rate constants at full-scale facilities, several examples also have been included. The rate constant is calculated from influent and effluent data under the assumption that plug-flow conditions were achieved, as demonstrated in earlier work.
5.4.3 Hydraulic Loading There is a significant difference in design hydraulic loading rates for upflow postdenitrification BAFs and post-denitrification BAFs that serve as the final filtration step (generally downflow BAF with sunken media or moving-bed continuous backwash filters). Average hydraulic loading rates for denitrification sand filters typically range from 4 to 9 m/h (1.5 to 3 gpm/sq ft), and peak hour rates typically are limited to 18 m/h (7.5 gpm/sq ft), with one cell out-of-service for backwashing. When upflow BAFs are used for post-denitrification applications, the larger media size allows for higher hydraulic loading rates of 10 to 35 m/h, but the solids retention capability may be reduced. Hydraulic loading rates also affect contact time within the filter. Original denitrification filter design curves (Savage, 1983) related the percent NO3-N removed to empty bed detention time (EBDT). Additional data from pilot- and full-scale systems superimposed onto these curves suggested that NOx-N removals of 90% could be achieved at an EBDT as low as 10 minutes, at temperatures ranging from 13 to 21°C, as shown in Figure 7.16 (deBarbadillo et al, 2005).
5.4.4 Solids Removal and Sludge Production In addition to removal of solids from the filter influent wastewater, biomass is produced within the filter, and both sources must be considered in determination of backwashing requirements. Generally, a biomass yield coefficient of 0.4 g biomass COD produced/g methanol COD consumed (approximately 0.4 g volatile suspended solids/g methanol consumed) is adequate for post-denitrification systems using methanol as the carbon source. When estimating solids quantities, it typically is assumed that approximately 10% of the biodegradable solids removed and produced undergo hydrolysis. Post-denitrification sand filters (downflow sunken media and moving-bed) typically serve as the final solids separation step and can produce effluent with an average TSS concentration of 5 mg/L or lower. The downflow, sunken media filter manufacturers have found that as much as 9.8 kg/m2 (2 lb/sq ft) solids can be captured reliably between backwashes (Slack, 2004). Typical design procedures account
Biological Filters
100
Percent NO3-N Removed, %
80 Data Points from Savage (1983) Water Temperature 16 to 19 C (Savage, 1983)
60
Water Temperature 11 to 15 C (Savage, 1983) Blue Plains Downflow Denitrification Filter Pilot, 5 ft media depth, temperature 16 to 21 C (calculated from Bailey et al, 1998)
40
Confidential Downflow Denitrification Filter, 6.6 ft media depth, temperature 30 C De Groot Lucht STP, UCB Denitrification Filter, 11.8 ft media depth, temperature 9 to 15 C De Groot Lucht STP, UCB Denitrification Filter, 11.8 ft media depth, temperature 16 to 24 C
20
Hagerstown WWTP UCB Denitrification Filter Pilot, 6.67 ft media depth, temperature 13 to 17 C
0 0
20
40
60
80
Empty Bed Detention Time, minutes
Figure 7.16 Denitrification filter design curves using EBDT (from Savage [1983], with additional data points [deBarbadillo et al., 2005]) (ft × 0.3048 = m). for backwash frequencies based on 4.9 kg/m2 (1 lb/sq ft). Data from a moving-bed, continuous backwashing denitrification filter pilot test in Hagerstown, Maryland, suggest that solids loading rates to the filters should be limited to 2.45 kg/m2 (0.5 lb/ sq ft) or lower, to maintain an average effluent TSS concentration of 5 mg/L or lower (Schauer et al., 2006). This solids loading rate limitation is specific to denitrification mode under the conditions tested and at a media recirculation rate of approximately 2 turnovers/d. Other BAF configurations, including upflow BAF with sunken media and upflow BAF with floating media, have installations that achieve final effluent TSS as low as 5 mg/L, but, with larger media sizes, overall solids filtration is not as robust, and the effluent may be more comparable with high-quality secondary clarifier effluent.
5.4.5 Supplemental Carbon Requirements For tertiary denitrification systems, such as filters, BAFs, and MBBRs with a postanoxic fixed-film zone, the supplemental carbon source is vital to system operation. Methanol feed requirements can be estimated using eq 7.7, as follows (McCarty et al., 1969):
369
370
Biofilm Reactors
SM = 2.47 (NO3-N removed) + 1.53 (NO2-N removed) + 0.87 (dissolved oxygen removed)
(7.7)
Equation 7.7 applies specifically to methanol. The COD requirement for d enitrification varies, depending on the carbon substrate used. The amount of substrate COD stabilized by 1 mg of oxygen is as follows (Copp and Dold, 1998; Melcer et al., 2003): COD/NO3-N = 1/(1 − Y)
(7.8)
where Y = heterotrophic yield coefficient (mg biomass COD formed/unit substrate COD used). A value of 0.66 typically is used for aerobic respiration. For denitrification, a 2.86 conversion factor is incorporated to the equation, to account for the amount of nitrate required to accept the same number of electrons. This yields the following expression for COD requirements for anoxic growth: COD/NO3-N = 2.86/(1 − Y)
(7.9)
An estimate of supplemental carbon substrate requirements can be made by factoring the appropriate anoxic yield coefficient into the calculation. Determination of anoxic yield coefficients for different substrates has been a topic of recent research (Cherchi et al., 2008; Mokhayeri et al., 2006; Nichols et al. 2007). For example, incorporating a methanol anoxic yield coefficient of 0.38 to eq 7.9 results in a requirement of 4.6 mg COD/mg NO3-N denitrified. Based on a COD-to-methanol ratio of 1.5, the methanol requirement is 3.07 mg methanol/mg NO3-N denitrified.
5.4.6 Tertiary Denitrification Typical Operations Issues and Corrective Actions 5.4.6.1 Excess Backwashing Excess filter backwashing is an issue, with respect to both operations costs (power usage) and filter performance. In denitrification filters that use intermittent backwashes (e.g., downflow with sunken media and upflow with sunken or floating media), there is some reduction in nitrate removal performance in a filter cell that has just been backwashed because of the reduction in biomass in that filter. If the filter is backwashing too frequently, it may be difficult to maintain enough biomass for denitrification performance. Backwashing frequency can be reduced by increasing the time between backwashes. In the case of moving-bed filters, if denitrifying biomass is not maintained in the filters, it may be necessary to reduce the bed turnover rate.
Biological Filters
5.4.6.2 Gas (Nitrogen) Accumulation In downflow denitrification sand filters, nitrogen gas accumulation within the media increases the headloss through the filters. This is addressed through periodic nitrogen-release cycles. The nitrogen-release cycle typically is initiated based on the water level in the filters. In moving-bed filters, gas accumulation typically does not occur, because the sand (and nitrogen gas bubbles) is drawn continuously through the airlift. Similarly, in upflow filters that use intermittent backwashing regimes, the co-current direction of the wastewater flow and nitrogen gas bubbles eliminate the need for nitrogen-release cycles.
5.4.6.3 Solids Breakthrough Solids breakthrough to the filter effluent may be an indication of overloading of the filters or problems with the underdrain. For static-bed filters, this may be an indication that the backwash frequency should be increased. For moving-bed filters, it may be necessary to increase the bed turnover rate.
5.4.6.4 Nitrate/Nitrite Breakthrough Excess nitrate and nitrite in the filter effluent generally is the result of one of the following three things:
(1) The filter is in a startup mode, where biomass is still being developed; (2) There is inadequate carbon being added to support full denitrification; or (3) There may be insufficient phosphorus in the filter influent wastewater to support denitrification.
If the filter is in a startup mode or a period of transition (e.g., if influent nitrogen loads to the plant increase, and the resulting nitrate loading to the filters increases), it is likely just a matter of time until the biomass develops and stable operation is achieved. During startup and transition periods, the backwashing frequency may be minimized, and carbon dosing should be checked frequently. If the filter is not receiving adequate carbon (methanol) dosing for complete denitrification, elevated nitrate and nitrite levels in the effluent will result. It may be necessary to increase the chemical dosing slightly to meet actual process requirements. Elevated dissolved oxygen levels as high as 5 or 6 mg/L are not unusual in a high-quality secondary effluent entering a denitrification filter, and this increases the chemical requirements. In addition, denitrification filters should not be operated “half-way.” If denitrification filters are operated for partial denitrification only, some nitrate is converted to nitrite, but is not denitrified fully to nitrogen gas. This results
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in an inability to meet nitrogen limits and increased chlorine demand in downstream chlorine disinfection facilities. If it is not necessary to remove all of the nitrate in the filter, it is more reliable, from an operating standpoint, to fully denitrify only a portion of the flow and blend the two effluent streams than to attempt to only partially denitrify all of the wastewater flow. If the filter is not undergoing a startup or transition period and is receiving sufficient chemical for denitrification, but still has significant nitrate or nitrite breakthrough, it is possible that the process may be phosphorus-limited. If the available phosphorus is found to be insufficient for denitrification, phosphoric acid can be dosed to the filter influent wastewater.
5.4.6.5 Carbon Breakthrough If elevated BOD levels are observed in the effluent, it may be the result of excess methanol dosing. The effluent soluble BOD (or COD) should be compared with secondary effluent values, to confirm that there is not an upset in the upstream activated sludge process. If the soluble BOD is increasing across the filters, the methanol dosage can be decreased.
5.4.6.6 Phosphorus Management Plants that must meet very stringent effluent total phosphorus limits, in addition to operating a tertiary denitrification process, should closely monitor phosphorus, nitrate, and nitrite concentrations in and out of the filter, to ensure that sufficient phosphorus is available for denitrification. It may be worthwhile to provide the capability to dose phosphoric acid during periods of upset.
5.4.6.7 Operation During Peak Flow Events Operation during peak flow events can be challenging because of the increased hydraulic throughput and the resulting increases in headloss through the filters. Depending on the operation of the secondary clarifiers upstream, the plant may experience high solids loadings to the filters during this time also, which increases the necessary backwash frequency. While it is possible to continue denitrifying during peak wet-weather flow events, it can become difficult to manage denitrification needs with the additional loading considerations. If the effluent nitrogen limit is enforced on a quarterly or annual average basis, it generally is not necessary to fully denitrify during peak wet-weather events. In this situation, peak flows above a certain level can be bypassed around the filter complex to reduce the load to the filter. Alternatively, methanol dosing can be discontinued during peak wet-weather events.
Biological Filters
5.5 Phosphorus Removal Considerations for Biologically Active Filter Processes Similar to those discussed for activated sludge systems, methods available for removal of phosphorus are as follows: • Pre-precipitation using metal salts (typically iron or aluminum) in primary settling, • Precipitation using metal salts at the biofilter stage, and • Biological removal. Phosphorus precipitation at the primary treatment stage is used widely, because high-rate settling processes (with coagulant addition) often are combined with BAFs. Multipoint chemical dosing can be used to reach very low effluent concentrations, but phosphorus limitation also has to be prevented to allow efficient biological reactions (Odegaard, 2005). Precipitation by adding iron salts to the biofilter also is possible, but the increase in the mass of solids entering the filter results in a higher backwash frequency. When ferric chloride for phosphorus removal was used in a two-stage plant, and the precipitant was added to the nitrifying stage, filter run times were reduced, and a smaller media was required to retain the floc (Sagberg et al., 1992). Biological phosphorus removal has been developed for suspended-growth systems, in which the alternation of anaerobic and aerobic zones by recycling of the biomass within the process encourages phosphorus uptake (Barnard, 1974). To achieve biological phosphorus removal with fixed biomass, the alternation can only take place in time, not in space. Therefore, a two-reactor alternating system was tested using two pilot biofilters in series—anaerobic and aerobic—and a 6-hour cycle (Goncalves and Rogalla, 1992). The system successfully achieved biological phosphorus removal and was later adapted for nitrogen removal, with one cell out of five switched into anaerobic mode (Goncalves, Le Grand, and Rogalla, 1994; Goncalves, Nogueira, Le Grand, and Rogalla, 1994). However, the additional expense in valving arrangements has prevented its application on full-scale. At some facilities, the need to meet stringent effluent total phosphorus limits while operating a tertiary nitrification or denitrification process is difficult. Adequate phosphorus is needed for microbial growth, and insufficient phosphorus will limit the ability of the tertiary BAF to achieve treatment goals. An evaluation of pilot- and full-scale postdenitrification performance data suggests that nitrogen removal goals can be achieved in practice at a orthophosphorus-to-NOx-N ratio of 0.02 (deBarbadillo et al, 2006).
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6.0 Design Considerations There are a number of issues that must be considered in the design of the BAF reactors, supporting facilities, and upstream and downstream processes.
6.1 Preliminary and Primary Treatment Fine screening should be implemented, if possible, at multiple locations in the plant, depending on the type of BAF used. Though the BAF influent already may have been screened several times, it is imperative to include a screen immediately upstream of the BAF inlet for upflow BAFs with nozzled bottom floors. A simple bag screen or manual flat screen with less than 2.5-mm openings will protect the BAF if an automatic fine screen is provided upfront. However, if influent wastewater screening is poor (or if some inlet streams are not screened (e.g., septage and imported sludges) or the plant is surrounded by trees and/or channels are not covered, dedicated automatic screens are needed. To achieve an overall small footprint for the treatment plant, high-rate primary treatment often is used. This may be achieved by chemically enhanced primary treatment (CEPT), high-rate lamella settlers, or by a ballasted (sand or dense sludge) flocculation and settling process.
6.2 Backwash Handling Facilities The BAF backwash facilities and equipment typically include effluent clearwell; backwash water pumps (depending on the type of BAF); air-scour blowers; backwash waste equalization tank and return pumps; and all automatic valves, instruments, and controls required for automatic initiation and sequencing of the backwash procedure. Equipment and facilities must be sized adequately to handle the air and water rates and volumes necessary for effective backwashing. During backwashing, the total effluent flow from the BAF may decrease or stop, and this flow variation must be accounted for in the design and operation of any downstream treatment process, such as UV disinfection. In multiple-stage BAF systems, where there could be several different BAF cell sizes and different backwash water- and air-scour requirements for each cell size, facilities and equipment sizes typically are based on requirements for the largest cell, to avoid separate sets of backwash equipment for each BAF stage. Final effluent taken directly from the effluent channel of the last stage of a multistage BAF system or via a final clearwell typically is used for backwash water. The effect of backwash return
Biological Filters
streams must be accounted for in the BAF design. If an interstage clearwell is used, the ability to maintain minimum flow to the downstream BAF cells must be considered, to avoid flow interruptions and effects on downstream processes, such as disinfection. Provisions for mixing should be considered for large backwash waste tanks to prevent solids settlement. Backwash waste also may contain some media, and, although media loss may be small, it will accumulate over time. The backwash waste-return pumping system should be designed to avoid drawing the media into the pumps and rising mains. Some BAF installations use media-recovery systems, including transfer pumps, settlement zones, and/or baskets. Backwash waste typically is returned to the head of the treatment plant, and the solids are removed in the primary settling tanks. This has been shown to improve the performance of the primary settling tanks, as the BAF biosolids adsorb some BOD and can simplify the pumping and handling of primary solids by improving solids rheology (Michelet et al., 2005). Alternatively, the backwash waste stream can be treated using a dedicated solids-separation system. This can be of particular benefit at larger facilities (larger than 100 000 m3/d [26 mgd]), if the existing primary settling tanks have solids-handling limitations or if there are multiple BAF stages. There are several technologies that can be used, such as a ballasted flocculation and settling system, a solids contact/sludge recirculation system, or a dissolved air flotation thickener (DAF). The DAFs have been used in this application at a number of European installations.
6.3 Biologically Active Filter Process Aeration In biofilm systems, diffusion of oxygen and substrate (BOD or nitrogen) into the biofilm is the limiting factor for the kinetic performance of the reactor. Blowers or compressors supply process air, which is distributed either by a grid of pipework or by diffusers located at or near the bottom of the reactor. While the air flows up the reactor through the media, oxygen is dissolved in the water and diffuses into the biofilm. The passage of the air bubbles also helps to maintain clear flow channels through the medium (Rundle, 2009).
6.3.1 Oxygen-Transfer Efficiency All aeration studies on BAF plants have observed higher oxygen uptake rates than generally found in the activated sludge process, which is consistent with the BAF’s reduced volume and hydraulic retention time. Stensel et al. (1984) measured oxygen uptake rates from 121 to 250 mg/L·h in a 1.7-m-tall reactor, which was 3.0 to 3.2 times
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greater than the observed rates in clean water tests in the same equipment. In addition to the classical transfer of dissolved oxygen in water to the biomass, this was attributed to a large portion of the sparged air bubble area being in direct contact with the biofilm, allowing a second mechanism of direct transfer of oxygen from air bubbles to the biofilm. Canziani (1988) and Lee and Stensel (1986) had similar findings. In their pilot columns (1.2 m high), Fujie et al. (1992) noted higher transfer rates with media (before colonization) than without. On the other hand, they also noted a decline in oxygen-transfer rate with increasing biofilm concentrations and attributed the reduction to bubble coalescence caused by biofilms. Work in the United States with upflow laboratory-scale BAFs (Lee and Stensel, 1986; Reiber and Stensel, 1985) showed how direct transfer related to biofilm thickness. In larger field demonstration pilots, oxygen-transfer tests of BAFs at a full depth of 3.6 m using the offgas method showed process water oxygen-transfer efficiencies of 1.6 to 5.8%/m (0.5 to 1.8 %/ft) for floating media (average 4-mm diameter) and 3.9 to 7.9%/m (1.2 to 2.4%/ft) for mineral media (3 to 5 mm) at their nominal design conditions (Redmon et al., 1983). The efficiency of oxygen transfer (in terms of kg O2 transferred/kWh of power) increased for both types, with increased depth of diffuser submergence. While higher airflow rates increased the oxygen-transfer rate, they reduced the oxygen-transfer efficiency. Additional studies yielded the following results: • Rogalla and Sibony (1992) measured oxygen-transfer rates of 7 to 15%. • Pearce (1996) measured clean water oxygen-transfer efficiencies of 10 to 17% in a downflow BAF pilot with 2-m depth and 3.3-mm angular media. • Shepherd et al. (1997) measured oxygen-transfer efficiencies of 7.9 to 10.3% in upflow BAFs with a 4-m depth of 2- to 3-mm silica sand. • Laurence et al. (2003) reported oxygen-transfer efficiencies of approximately 20% from offgas testing of a 3-m depth upflow BAF with floating media and a 4-m depth upflow BAF with sunken media, during side-by-side pilot testing in New York City. • Leung et al. (2006) measured liquid side oxygen-transfer coefficients (kLa) under abiotic conditions and found them to increase as both gas and liquid superficial velocity increases. • Stenstrom et al. (2008) measured oxygen-transfer efficiencies of 5.8 to 21.1% for 3.6 m of 3- to 5-mm rock media and 13.1 to 29% for 3.6 m of 4-mm Styrofoam spheres in pilot reactors.
Biological Filters
6.3.2 Process Air Distribution Systems Process air distribution systems in BAFs include the following (Rundle, 2009):
(1) Simple pipes with sparge holes drilled at intervals positioned in media or near the floor of the filter. Coarse-bubble aeration through sparging pipes is used widely. (2) Diffusers placed on a pipe grid at the floor of the reactor to obtain even air distribution at low airflow rates, rather than to produce smaller bubbles for improved oxygen-transfer efficiency. Although diffusers are more efficient for oxygen transfer than coarse-bubble sparging in open aeration basins, a comparison between coarse- and fine-bubble aeration did not reveal any difference (Harris et al., 1996). In a comparison of coarse- and fine-bubble aeration in reactors with and without plastic random media, fine-bubble diffusers were found to be more efficient without packing (Hodkinson et al., 1998). The media-sheared coarse bubbles favored dispersion into smaller bubbles with a larger surface area and improved oxygen transfer. However, the fine bubbles coalesced into larger bubbles and reduced oxygen transfer. (3) Injection of air under the plenum, frequently used to scour filters during backwash, also can be used during filtration. In this design, an air blanket is formed under the false floor in the plenum chamber; air enters the cell via holes in the specially designed combined nozzle. At low airflows, only the upper holes are used, but, as airflow and pressure increase, blanket depth increases, and more holes are used. This system provides efficient aeration, but requires periodic chemical cleaning to prevent biological growth from blocking the air holes, causing poor air distribution and increasing energy costs (Holmes and Dutt, 1999; Springer and Green, 2005).
6.3.3 Process Air Control Process air control in BAF reactors is complicated by the following factors (Rundle, 2009):
(1) BAF plants operate primarily as plug-flow systems, so that the dissolved oxygen at the top of the reactor does not represent the dissolved oxygen concentration within the media; (2) Oxygen transfer not only takes place from dissolved oxygen in water, but also occurs by direct interfacial transfer from gas to biofilm, which cannot be accounted for with a dissolved oxygen probe; and
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(3) For BAF systems aerated by a coarse-bubble air grid, the minimum flow to provide effective distribution of air can exceed process requirements.
Blower selection is important for efficient plant operation. As solids accumulate in the media, the filter headloss increases, which can affect the airflow. When several BAF cells receive air from a common air main, the most recently backwashed cell or cells will have the lowest headloss and will take more airflow than the cells that have not been backwashed. This balancing issue can be mitigated by providing individual blowers for each BAF cell. For larger plants with a centralized blower station and a common air main, the air pipes feeding each cell are fitted with a mass flow meter (measuring velocity, pressure, and temperature). The meter is used to control a modulating valve, which balances airflow to the cells.
6.4 Supplemental Carbon Feed Requirements In tertiary denitrification systems and in some pre-denitrification, an external carbon substrate (electron donor) must be dosed to the BAF. Methanol typically has been used for this purpose. Increasingly, alternative carbon sources are being considered, including ethanol, acetic acid, and sugar solutions. The chemical properties for the selected carbon source must be evaluated and accounted for in design. Carbon dosage control is important for tertiary denitrification systems. Overfeeding wastes chemical, and could increase the BOD of the effluent, which could be an issue for plants with BOD limits of approximately 5 mg/L or lower. Underfeeding the carbon source reduces the amount of nitrate removed, and the plant may not achieve the desired effluent nitrate or total nitrogen concentration. Several alternatives exist for controling carbon dosing to BNR processes, as follows: • Manual control—for manual control of chemical dosing, all pumping rate adjustments and sampling are performed manually. • Flow-paced control—based on influent nitrate concentration and the required level of nitrate removal, the average carbon dose requirement is determined. The control system then is set to modulate pumping rate with fluctuations in wastewater flow. Typically, flow-pacing applies only to dry-weather operation. • Feed-forward control—a feed-forward control scheme, in which denitrification influent nitrate concentrations are measured and used in combination with flow to vary the carbon feed rate, offers the next level of automatic control.
Biological Filters
Because the carbon dose is based on both wastewater flow and concentration, it is feasible to operate in this mode during wet and dry weather. • Feed-forward and feedback with effluent concentration control—this represents the most complex level of chemical feed control. Systems with this capability currently are offered as proprietary packages by several denitrification filter system suppliers. Some are based on flow and nitrate only, while others incorporate nitrite and dissolved oxygen readings.
7.0 Biologically active filter Case Studies 7.1 Chemically Enhanced Primary Treatment Followed by TwoStage Biologically Active Filter for Total Nitrogen Removal: VEAS Wastewater Treatment Plant, Oslo, Norway The VEAS WWTP in Oslo, Norway, has a maximum peak design flow of 492 500 m3/d (130 mgd). As illustrated in Figure 7.17 the plant facilities consist of the following:
(1) Fine (3-mm) screens; (2) Aerated grit chambers with coagulant dosing; (3) Deep (11-m [36-ft]) primary clarifiers with 75-cm-long lamellae; and (4) Two stages of upflow BAF (Biofor), protected by a 1.5-mm fine screen • The first stage for carbon removal/nitrification with 24 Biofor cells, each with 87-m2 (936-sq ft) filtration area and 3.8-m (12.5-ft) media depth; and • The second stage for post-denitrification with methanol addition and 24 Biofor cells, each with 65-m2 (722-sq ft) filtration area and 3.0-m (9.8-ft) media.
Backwash wastewater from the BAFs is recycled back to the head of the plant, and the biological solids are co-settled in the primary clarifiers with chemical and primary solids. Solids-processing facilities include drum thickening, two-stage acid-gas digestion, and combined filter press dewatering/thermal vacuum drying. Digester biogas is burned in gas engines for combined heat and power production. Filtrate from the filter press is treated through ammonia stripping, and the stripped filtrate water is returned to the outlet of the aerated grit chambers. The VEAS plant is built entirely within a hill located near the town of Asker at the west side of the Oslo fjord, approximately 20 km (12 miles) south of Oslo. The
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Wash water return Flocculant
Polymer Sedimentation tank
Screens Inlet Aerated grit chamber
Internal return (I.R.)
Nitrification filter
Denitrification filter
Screen
Outlet Wash water tank
I.R.
Rag strainer Air Polymer
Electric power production
Fluegas
Methanol Drum thickeners
LT
HT
Gasengine Biogas
Decant (I.R) Sludge equalisation tank
Bio hydrolysis
Nitric acid
Ammonia stripping Air/ ammonia Stripped filtrate water (I.R.)
Air
Ammonia nitrate Slaked Buffer lime Anaetank Rag robic strainer digestion
HT Polymer
LT
HT
Heating
Filtrate water Combined filterpress and thermal vacuum dryer
Veas biosolids
Figure 7.17 VEAS WWTP schematic.
BAFs were installed between 1994 and 1996 as part of a major upgrade to meet the new nitrogen limits and more stringent phosphorus limits required by the North Sea Agreement. The VEAS is required to meet 90% phosphorus removal and 70% total nitrogen (TN) removal (<7 mg/L TN) as an annual average. Limited space availability was a major factor in the decision to implement BAF technology for nitrogen removal, as the biological treatment was inserted to existing tunnels first carved for primary settling. The retention time from plant influent to effluent is less than 3 hours at annual average flow. The influent wastewater is dilute and contains low levels of readily biodegradable organic carbon, as a result of the long detention time (5 to 7 hours during dryweather periods) in the aerated influent tunnel. The organic carbon is reduced further in the chemically enhanced primary clarifiers, which typically achieve 60 to 65% total organic carbon (TOC) removal. The wastewater temperature typically ranges from 5 to 16°C and has been as low as 2°C.
Biological Filters
Methanol is dosed upstream of the feed pumps to the denitrifying Biofor units. The methanol dose is controlled automatically using a feed-forward/feedback control loop with an effluent nitrate trim target value of 0.9 to 1.3 mg/L. The specific methanol dose is approximately 2.3 g methanol/g NO3-N denitrified. The Biofors are backwashed every 12 to 20 hours on an annual average, resulting in a backwash return flow of approximately 12% of the influent flow. Annual media loss is reported to be less than 1% per year. Table 7.9 provides typical annual average performance values for the two-stage Biofor total nitrogen removal system at VEAS. The data presented in the table are from 1997 to 1998, when the actual annual average flow was 265 753 m 3/d (70 mgd).
7.2 Chemically Enhanced Primary Treatment Followed by ThreeStage Biologically Active Filter for Total Nitrogen Removal: Siene Centre Wastewater Treatment Plant, Colombes, France The Seine Centre WWTP, one of five wastewater treatment facilities owned and operated by the Syndicat Inter départemental pour l’Assainissement de l’Agglomération Parisienne (SIAAP) (Paris, France), is located on a 4-ha (9.9-ac) urban site in Colombes, a near-northwest suburb of Paris. The entire plant is covered and has a rated capacity of 240 000 m3/d (63.4 mgd) with a peak wet-weather flow of 734 400 m3/d (194 mgd). The plant facilities consist of the following:
(1) Pre-treatment with bar racks, grit removal, and 6-mm fine screens; (2) CEPT using high-rate, solids contact clarifiers (Degrémont DensaDeg); and
Table 7.9 Performance results for two-stage BAF in Oslo, Norway. First-stage influent concentration (mg/L)
Second-stage effluent concentration (mg/L)
Volumetric loading (on 22 of 24 cells per stage) (kg/m3·d)
Removal efficiency (%)
Hydraulic loading (m/h)
TOC
25
13
0.56
48
Total nitrogen
19
6.2
0.43
67
First stage = 5.7 Second stage = 7.4
Parameter
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(3) Three stages of BAF, as follows, protected by 1.5-mm fine screening: • The first stage for carbon removal (Biofor C/N), • The second stage for nitrification (Biostyr N), and • The third stage for post-denitrification (Degremont Biofor DN).
The flow scheme is shown in Figure 7.18. Backwash wastewater from the BAFs is sent to DAF thickeners. The DAF subnatant flow is recycled back to the head of the plant, with dewatering centrate and other recycle flows. The DAF-thickened biological solids are combined with chemical and primary sludge from the Densadeg clarifier/ thickener and dewatered through centrifugation. Cake from the centrifuge is incinerated in fluidized-bed incinerators. During dry-weather flows, the three BAFs operate in series to eliminate residual BOD, nitrify, and then denitrify with the addition of methanol. When moderate rainfall occurs, the first and third biofilter stages run in parallel to eliminate organics, while the second stage carries out nitrification in series with the first. In high-flow
1 Preliminary Treatment
2 Primary Clarification and Phosphorus Removal
Raw Wastewater
Sewer
Bar Racks Grit Removal
Fine Screen 6 mm
DensaDeg Clarifier
Fine Screen 1.5 mm
3 Biological Nitrogen Removal I BOD Removal II Nitrification
III Denitrification BIOFOR C/N BIOSTYR N BIOFOR DN
Figure 7.18 Seine Centre WWTP liquid treatment schematic.
Effluent La SEINE
Biological Filters
conditions caused by severe wet-weather events and a peak factor of up to 4.3, all three biofilter stages operate in parallel to reduce BOD simultaneously, as nitrogen is diluted severely in those events. The specific methanol dose to the third-stage BAF is approximately 3.0 g methanol/g NO3-N applied. Total backwash wastewater flow from all three BAF stages is approximately 18% of the dry-weather influent flow. Table 7.10 provides annual average performance values from the year 2000 for the three-stage BAF total nitrogen removal system at Seine Centre WWTP. Tables 7.11 through 7.13 provide performance values for each stage individually. In addition to the three-stage configuration described above, the Seine Centre plant is capable of operating the first-stage Biofor cells in the pre-denitrification mode (Biofor Pre-DN) during dry-weather flow conditions. In this configuration, effluent containing nitrates produced in the second-stage Biostyr is recycled to the inlet of the first stage for denitrification using readily available carbon in the primary effluent. The recycle ratio is limited by the maximum allowable hydraulic loads on the first- and second-stage BAFs. The maximum hydraulic loads are 10 m/h (4.1 gpm/ sq ft) and 8 m/h (3.3 gpm/sq ft) on the first-stage pre-denitrification and secondstage nitrification, respectively. Consequently, the maximum allowable recycle ratio based on BAF hydraulic load limitations is 150%. Cost analyses performed by SIAAP in 2005/2006 concluded that a recycle ratio of 80 to 100% is optimal, in terms of net cost savings in methanol usage and aeration energy consumption. At Seine Centre’s operating recycle ratio of 50%, the pre-denitrification stage achieves approximately Table 7.10 Year 2000 average performance for three-stage BAF system at Seine Center WWTP.
Parameter COD BOD5 TSS TKN Total nitrogen
Three-stage BAF influent concentration (mg/L)
Volumetric loading* (kg·m3·d)
Final effluent concentration (mg/L)
Removal efficiency (%)
118 56 26 26 27
1.7 0.78 0.36 0.36 0.38
34 10 6 7 13
72 86 73 42 52
* 60 of 64 cells in filtration.
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Table 7.11 Year 2000 average performance for first-stage BAF system at Seine Center WWTP.
Parameter COD BOD5 TSS TKN Total nitrogen
First-stage influent concentration (mg/L)
Volumetric loading * (kg·m3·d)
First-stage e ffluent concentration (mg/L)
Removal efficiency (%)
118 56 26 26 27
5.9 2.8 1.3 1.3 1.3
33 8 7 15 21
72 86 73 42 22
* Corresponds to hydraulic loading of 5.7 m/h (2.3 gpm/sq ft).
70% nitrate removal. This removal efficiency has been observed to be independent of the influent soluble COD-to-nitrate ratio, indicating that there is a surplus of soluble COD in the primary effluent.
7.3 Total Nitrogen Removal in a Single-Stage Biologically Active Filter: Frederikshavn Wastewater Treatment Plant, Denmark The Frederikshavn WWTP has a rated capacity of 16 500 m3/d (4.4 mgd), with a peak flow of 36 300 m3/d (9.6 mgd). The plant facilities consist of screens, inlet pumping, Table 7.12 Year 2000 average performance for second-stage BAF system at Seine enter WWTP. C
Parameter COD BOD5 TSS TKN Total nitrogen
Second-stage influent concentration (mg/L)
Volumetric loading * (kg·m3·d)
Second-stage effluent concentration (mg/L)
Removal efficiency (%)
33 8 7 15 21.1
1.5 0.4 0.3 0.7 0.98
23 4 5 1.6 20.8
30 50 29 89 1.4
* Corresponds to hydraulic loading of 4.6 m/h (1.9 gpm/sq ft).
Biological Filters
Table 7.13 Year 2000 average performance for third stage BAF system at Seine enter WWTP. C
Parameter COD COD (including methanol)b TSS TKN NO3-N Total nitrogen
Third-stage influent concentration (mg/L)
Volumetric loading, (kg·m3·d)
Third-stage effluent concentration (mg/L)
Removal efficiency (%)
23 109
2.6 12.6
34 34
– 69
5 1.6 19.1 20.8
0.6 0.61 2.2 2.40
6 7 6.1 13.4
– – 68 36
Corresponds to hydraulic loading of 12.3 m/h (5.0 gpm/sq ft). Methanol dosage = approximately 3.0 kg methanol/kg NO3-N applied.
a
b
combined grit and grease removal, chemically enhanced rectangular primary settlers, and the parallel operation of two biological reactors (Figure 7.19), as follows: • A Modified Ludzack–Ettinger (MLE) activated sludge with circular clarifiers, receiving approximately 40%, or 6600 m3/d (1.7 mgd) of the flow and • A Biostyr BAF system, loaded with approximately 60%, or 9900 m 3/d (2.6 mgd). There are six Biostyr cells, each with 63 m2 (678 sq ft) of filtration area and 3.0 m (9.8 ft) of media depth. There are two aeration grids in each Biostyr cell—one located at the filter bottom, and another located approximately 1.4 m (4.5 ft) up from the filter bottom, to create an anoxic zone in the lower part of the filter bed. While this case study covers typical dry-weather operation (pre-denitrification/ nitrification in five filters, with post-denitrification in one filter), the Biostyr system at Frederikshavn also can be operated in other ways to handle wet-weather peaks, as follows:
(1) Single-stage pre-denitrification/nitrification in all six cells; (2) Pre-denitrification/nitrification in five filter cells, with post-denitrification in one filter cell (typical dry-weather operation);
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Max. hydraulic load, dry weather: 1.530 m3/h Max. hydraulic load, rainfall: 4.330 m3/h Backwash water from BioStyr: 3.000 m3/d
Recirculation
DN-talk
Inlet well
Grit
N-talk
Primary settlers
Discharge tower
Screen
Rainwater bypass: max. 2.800 m3/h
Clarifiers
BioStyr
Figure 7.19 Frederikshavn Central WWTP liquid treatment schematic.
(3) Full nitrification in all filter cells (operation during maximum wet-weather load); and (4) Controlled oxygen set-point, with simultaneous nitrification and denitrification in all filter cells.
While the total media volume is 1134 m3 (40 030 cu ft), for the pre-denitrification mode, approximately 794 m3 (28 030 cu ft) of the media volume is aerated for nitrification, and 340 m3 (12 000 cu ft) is unaerated. For the single post-denitrification filter cell, all 189 m3 (6671 cu ft) of the media is unaerated. Methanol is dosed upstream of the post-denitrification Biostyr. The methanol dose is controlled automatically using a feed-forward control loop. The specific methanol dose is approximately 3 g methanol/g NO3-N applied. The Biostyr units are backwashed every 24 hours, on average, resulting in a backwash return flow of approximately 33% of the influent flow to the BAF and 18% of the total influent flow to the plant. Backwash wastewater from the BAFs is returned to the primary clarifiers to be co-settled with primary solids. Primary sludge, which includes biological solids from the BAF, and waste activated sludge (WAS) from the MLE are stabilized through anaerobic digestion before disposal.
Biological Filters
Approximately 50% of the wastewater comes from the fish-processing industry. The remaining 50% is a mixture of other industrial and domestic wastewater. Table 7.14 provides annual average performance values from the year 1998 for the twostage BAF total nitrogen removal system at Frederikshavn Central WWTP, while the processes were without optimum control regarding aeration, methanol addition, and nitrate recirculation.
7.4 Nitrification and Denitrification: West Warwick, Rhode Island The Advanced Wastewater Treatment Facility (ATF) located in West Warwick, Rhode Island, treats an average daily flow of 40 000 m3/d (10.5 mgd), with a peak capacity of approximately 96 000 m3/d (25.34 mgd). The plant included screening, introduction of aluminum sulfate for phosphorus precipitation at the plant headworks, primary clarification, and secondary treatment, consisting of six aeration basins and two secondary clarifiers. The facility was upgraded with secondary effluent pumping; BAF for tertiary nitrification and denitrification; post aeration; and UV disinfection. Hydrated lime is added for supplemental alkalinity, and methanol provides supplemental carbon for denitrification. The BAF technology was chosen because of a constrained site and the need for the facility to meet stricter regulatory requirements (Figure 7.20). The BAF units consist of the following: Table 7.14 Year 1998 average performance for the two-stage BAF at rederikshavn Central WWTP. F
Parameter
Biostyr influent concentration (mg/L)
Biostyr effluent concentration (mg/L)
Volumetric loading* (kg/m3·d)
Removal efficiency (%)
COD Filtered COD BOD5 TSS TKN NH3-N NO3-N Total nitrogen
250 130 200 92 39 22 0 39
– – 2.7 4.9 – 0.6 – 5.4
2.2 1.1 1.8 0.8 0.3 0.2 – 0.3
– – 99 95 – 97 – 86
* Five of six cells in filtration; corresponds to hydraulic loading of 12.3 m/h (5.0 gpm/sq ft).
387
NITRIFICATION AND DENITRIFICATION FILTERS Electrical Room
Backwash pumps and blowers
Effluent Pipe
Influent Pumps
3N 3DN
1N
1DN
Pipe Gallery
388 Chemical Building
4DN
2DN 4N
2N
Clearwell
Mudwell
Denitrification Filters Nitrification Filters
Influent Pipe Process Air Blowers
Figure 7.20 Nitrification and denitrification BAF reactors at West Warwick, Rhode Island.
BAF Building
Biological Filters
• Four nitrification filters cells, each with a surface area of 100 m 2 (1080 sq ft) and containing an angular clay media with an effective size of 2.7 mm and designed on hydraulic loadings at annual average flow of approximately 132 m3/m2·d (2.25 gpm/sq ft); • Four denitrification filters cells, each with a surface area of 41.6 m2 (448 sq ft) and containing an spherical clay media with an effective size of 3.5 mm and designed on hydraulic loadings at annual average flow of approximately 320 m3/m2·d (5.4 gpm/sq ft). The effluent is discharged to the Pawtuxet River, a Narragansett Bay tributary, and is subject to seasonal nutrient removal requirements. Construction of the BAF facility began in the fall of 2002, and the BAF was completed 2 years later. Startup and performance testing occurred during 2005.
7.5 Post-Denitrification Sand Filters: Havelock, North Carolina The Havelock WWTP in eastern North Carolina is rated for 7200 m3/d (1.9 mgd) of monthly average flow. The plant facilities include influent pumping; screening; grit removal; an activated sludge system, consisting of two sets of aeration basins operating in series; final clarifiers; deep-bed denitrification filters; UV disinfection; and mechanical post-aeration. Solids-handling facilities include aerobic digestion, gravity belt thickening, and a sludge storage basin. Polyaluminum chloride is added to the clarifier influent for phosphorus precipitation, and methanol is added to the filter influent to provide a carbon source for denitrification. The Havelock WWTP is required to meet an effluent total nitrogen limit of 9700 kg/a (21 400 lb/yr), or 3.7 mg/L total nitrogen at the design flow. The activated sludge system is operated with a mixed-liquor suspended solids setpoint of 5000 to 6000 mg/L (corresponding to a solids retention time of 25 to 30 days). The plant has operated well in this mode, with effluent ammonia concentrations consistently lower than 0.2 mg/L. The influent nitrate concentration to the filters averages approximately 12 mg/L. Since denitrification in the filters began in 1998, the Havelock WWTP has consistently met its total nitrogen limit. The most significant challenge, with respect to meeting the nitrogen limit, initially was related to methanol feed control. In 1998, the National Pollutant Discharge Elimination System (NPDES) permit had specified a carbonaceous BOD (CBOD) limit of 3 mg/L (monthly average). This CBOD limit was the most stringent in the state
389
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Biofilm Reactors
and would have been difficult for any plant to meet. The addition of methanol to the filters for denitrification made it even more challenging. At first, the methanol feed was controlled through flow-pacing. While the system performed well, it sometimes was difficult to balance the requirements for very low CBOD with denitrification. In particular, it was difficult to closely match the carbon requirement under varying nitrate concentrations. As a result, an online nitrate analyzer was installed. The control algorithm was first modified to include feedback based on the effluent nitrate concentration and subsequently was modified to incorporate feed-forward and feedback controls based on flow and influent and effluent nitrate (patented TetraPace system [Severn Trent Water Purification Inc., Ft. Washington, Pennsylvania]). This has enhanced the operation and reliability of the denitrification process without risk of methanol overdose by more closely matching the methanol feed to the demand. It also has reduced variations in effluent nitrate. The plant staff estimates that methanol consumption has been reduced by approximately 30%. In addition, the plant achieved a more consistent effluent quality, as illustrated by the annual averages and standard deviations presented in Table 7.15. The Havelock WWTP now has over 9 years of operation with the denitrification filters, with excellent results. The effluent total nitrogen averaged 3 mg/L in 2001 Table 7.15 Havelock WWTP comparison of effluent BOD and total nitrogen average and variation.
Parameter CBOD average (mg/L) CBOD standard deviation Total nitrogen average (mg/L) Total nitrogen standard deviation
1998 (Before installation of TetraPace system [Severn Trent Water Purification, Inc.])
1999
2001a
2002b
2.8c 4.07 4.28
2.62c 0.85 3.09
2.81d 1.03 2.72
2.9d 0.99 3.32
2.96
1.76
1.51
2.28
April 2001 through March 2002. April 2002 through March 2003. c CBOD5. d BOD5. a
b
Biological Filters
and 2002. Approximately 2 mg/L of the effluent total nitrogen is in the form NOx-N. Currently, the main operations challenge is related to high flows during storm events. During peak flow events, the filters typically are taken out of denitrification mode and are operated for filtration only.
8.0 REFERENCES Bailey, W.; Tesfaye, A.; Dakita, J.; Benjamin, A.; McGrath, M.; Sadick, T.; Daigger, G.; Tucker, M. (1998) Demonstration of Deep Bed Denitrification and the Blue Plains Wastewater Treatment Plant. Proceedings of the 71st Annual Water Environment Federation Technical Exposition and Conference, Orlando, Florida, Oct 3–7; Water Environment Federation: Alexandria, Virginia. Barnard, J. L. (1974) Cut P and N without Chemicals. Water Wastes Eng., 11, 41–44. Brewer, P.; Martin, J. C.; Bedard, P. (1997) Lamella Plate Separators and Biological Aerated Filters at Poole STW. Proceedings of International Conference on Advances in Wastewater Treatment Processes, Leeds, United Kingdom, Sept 8–11; Aqua Enviro Ltd.: Wakefield, United Kingdom. British Standards Institution (1983) 1971 Specification for Media for Biological Percolating Filters, BS1438. British Standards Institution: London, United Kingdom. Brown, S. (1992) Treatment of Effluent Using the Colox Process. Paper Technol., 33 (9), 36–38. Cantwell, A.; Mosey, F. (1999) Recent Applications and Developments of the Biobead System. Proceedings of the BAF3 Conference, Cranfield University: Bedfordshire, United Kingdom. Canziani, R. (1988) Submerged Aerated Filters IV—Aeration Characteristics. Ingegneria Ambientale, 17, 627–636. Chen, J. J. (1980) Plant-Scale Operation of a Biological Denitrification Filter System. Paper presented at the American Society of Civil Engineers Annual Conference, Oct 27–31, Hollywood, Florida; American Society of Civil Engineers: Reston, Virginia. Cherchi, C.; Onnis-Hayden, A.; Gu, A. Z. (2008) Investigation of MicroCTM as an Alternative Carbon Source for Denitrification. Proceedings of the 81st Annual
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Water Environment Federation Technical Exposition and Conference, Chicago, Illinois, Oct 18–22; Water Environment Federation: Alexandria, Virginia. Churchley, J. H.; Jarvis, M.; Pickett, H. (1990) Aerated Filters, Submerged or Biological. Paper presented to the East Midlands Branch of the Institute of Water and Environmental Management. Cooper-Smith, G.; Schofield, I. (2004) Submerged Aerated Filters, Coming of Age for AMP4. Proceedings of the 2nd National CIWEM Conference, Wakefield, United Kingdom, Sept 13–15; Aqua Enviro Ltd.: Wakefield, United Kingdom, 55. Copp, J. B.; Dold, P. L. (1998) Comparing Sludge Production Under Aerobic and Anoxic Conditions. Water Sci. Technol., 38, 285–294. Daude, D.; Stephenson, T. (2004) Cost-Effective Treatment Solutions for Rural Areas; Design of a New Package Treatment Plant for Single Households. Water Sci. Technol., 48, 107–114. deBarbadillo, C.; Rectanus, R.; Canham, R.; Schauer, P. (2006) Tertiary Denitrification and Very Low Phosphorus Limits: A Practical Look at Phosphorus Limitations on Denitrification Filters. Proceedings of the 79th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 21–25; Water Environment Federation: Alexandria, Virginia. deBarbadillo, C.; Shaw, A.; Wallis-Lage, C. (2005) Evaluation and Design of Deep-Bed Denitrification Filters: Empirical Design Parameters vs. Process Modeling. Proceedings of the 78th Annual Water Environment Federation Technical Exposition and Conference, Washington, D.C., Oct 29–Nov 4; Water Environment Federation: Alexandria, Virginia. Degrémont (2008) Correspondence from Troy Holst providing recommended design loading rates for Biofor BAF units. May. . Degrémont (2007) Water Treatment Handbook, 7th ed.; Lavoisier SAS: Paris, France. Downing, A. L.; Painter, H. A.; Knowles, G. (1964) Nitrification in the Activated Sludge Process. J. Proc. Inst. Sew. Purif., 2, 130. Downing, A. L.; Tomlinson, T. G.; Truesdale, G. A. (1964) Effect of Inhibitors on Nitrification in the Activated Sludge Process. J. Proc. Inst. Sew. Purif., 6, 531–554. European Union (1991) Council Directive Concerning Urban Wastewater Treatment, May. European Union: Brussels, Belgium.
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Filtralite, Oslo, Norway (2008) Email correspondence from Paul Sagberg transmitting information from FoU Report for VEAS, Oslo, Norway (Table 6.1). Fitzpatrick, C. S. B. (2001) Factors Affecting Efficient Filter Backwashing. Proceedings of the International Conference on Advances in Rapid Granular Filtration in Water Treatment, London, United Kingdom, April 4–6; Chartered Institution of Water and Environmental Management: London, United Kingdom. Frankl, S. (2004) Wessex Water’s Experience of SAF Package Plants. Presented at Symposium on Small Sewage Treatment Works in the 21st Century, March 3. Froud, D. P. (1994) Br. Patent GB2,270,909. Fujie, K.; Hu, H.; Ikeda, Y.; Urano, K. (1992) Gas-Liquid Oxygen Transfer Characteristics in an Aerated Submerged Biofilter for Wastewater Treatment. Chem. Eng. Sci., 47, 3745–3752. German Association for Water, Wastewater and Waste [ATV-DVWK] (1997). Biological and Other Wastewater Purification (Biologische und weitergehende Abwasserreinigung), 4th ed.; Ernst & Sohn: Berlin, Germany (in German). German Association for Water, Wastewater and Waste [ATV-DVWK] (2000) Standard ATV-DVWK-A 131 E, Dimensioning of Single-Stage Activated Sludge Plants. German ATV-DVWK Rules and Standards; Ernst & Sohn: Berlin, Germany. Goncalves, R.; Rogalla, F. (1992) Continuous Biological Phosphorus Removal in a Biofilm Reactor. Water Sci. Technol., 26, 2027–2030. Goncalves, R. F.; Le Grand, L.; Rogalla, F. (1994) Biological Phosphorus Uptake in Submerged Biofilters with Nitrogen Removal. Water Sci. Technol., 29, 135–143. Goncalves, R. F.; Nogueira, F. N.; Le Grand, L.; Rogalla, F. (1994) Nitrogen and Biological Phosphorus Removal in Submerged Biofilters. Water Sci. Technol., 30, 1–12. Harremoes, P. (1976) The Significance of Pore Diffusion to Filter Denitrification. J. Water Pollut. Control Fed., 48, 377–388. Harris, S. L.; Stephenson, T.; Pearce, P. (1996) Aeration Investigation of Biological Aerated Filters Using Off-Gas Analysis. Water Sci. Technol., 34, 307–314. Hodkinson, B. J.; Williams, J. B.; Ha, T. N. (1998) Effects of Plastic Support Media on the Diffusion of Air Into a Submerged Aerated Filter. J. CIWEM, 12, 188–190.
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Holmes, J.; Dutt, S. (1999) Coln Bridge (Huddersfield) WWTW Biopur Plant Process Design and Performance. Proceedings of the BAF3 Conference, Cranfield University: Bedfordshire, United Kingdom. Hultman, B.; Jonsson, K.; Plaza, E. (1994) Combined Nitrogen and Phosphorus Removal in a Full-Scale Continuous Upflow Sand Filter. Water Sci. Technol., 29, 127–134. Husovitz, K. J.; Gilmore, K. R.; Delahaye, A.; Love, N. G.; Little, J. C. (1999) The Influence of Upflow Liquid Velocity on Nitrification in a Biological Aerated Filter. Proceedings of the 72nd Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Oct 9–13; Water Environment Federation: Alexandria, Virginia. International Water Association (2006) Recent Progress in Slow Sand and Alternative Biofiltration Processes, Gimbel, R., Graham, N. J. D., Collins, M. R. (Eds.); IWA Publishing: London, United Kingdom. Janning, K. F.; Harremoes, P.; Nielsen, M. (1995) Evaluating and Modelling of the Kinetics in a Full-Scale Submerged Denitrification Filter. Water Sci. Technol., 32, 115–123. Jolly, M. (2004) Aberdeen (Nigg) Wastewater Treatment Works—1st Year of Operation. Proceedings of the 2nd National CIWEM Conference, Wakefield, United Kingdom, Sept 13–15; Aqua Enviro Ltd.: Wakefield, United Kingdom, 103. Karschunke, K.; Sieker, K. (1997) Limits of Denitrification in Biofilters at the Wastewater Treatment Plant at Nyborg as an Example (Grenzen der Denitrification in der Biofiltrationstechnik am Beispeil der Klaranlage Nybor). Wasser Abwasser, 7, 337–343 (in German). Kent, T. D.; Williams, S. C.; Fitzpatrick, C. S. B. (2000) Ammonia Nitrogen Removal in Biological Aerated Filters, the Effect of Media Size. J. CIWEM, 14, 409. Kruger (2008) E-mail correspondence with Michele Kline of Kruger/Veolia regarding BAF design recommendations. Laurence, A.; Spangel, A.; Kurtz, W.; Pennington, R.; Koch, C.; Husband, J. (2003) Full-Scale Biofilter Demonstration Testing in New York City. Proceedings of the 76th Annual Water Environment Federation Technical Exposition and Conference, Los Angeles, California, Oct 11–15; Water Environment Federation: Alexandria, Virginia.
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Le Tallec, X.; Zeghal, S.; Vidal, A.; Lesouef, A. (1997) Effect of Influent Quality Variability on Biofilter Operation. Water Sci. Technol., 36, 111–117. Lee, K. M.; Stensel, H. D. (1986) Aeration and Substrate Utilization in a Sparged Packed-Bed Biofilm Reactor. J. Water Pollut. Control Fed., 58, 1066–1072. Leung, S. M.; Little, J. C.; Holst, T.; Love, N. G. (2006) Air/Water Oxygen Transfer in a Biological Aerated Filter. J. Environ. Eng., 132, 181–189. McCarty, D., Severn Trent, Coventry, United Kingdom (2008) Personal communication. McCarty, P. L.; Beck, L.; Amant, P. S. (1969) Biological Denitrification of Wastewaters by Addition of Organic Materials. Proceedings of the 24th Purdue Industrial Waste Conference, West Lafayette, Indiana, May 6–8; Purdue University: West Lafayette, Indiana, 1271–1285. Meaney, B. (2007) Operation of Submerged Filters by Anglian Water Services Ltd. Water Environ. J., 8, 327–334. Melcer, H.; Dold, P. L.; Jones, R. M.; Bye, C. M.; Takacs, I.; Stensel, H. D.; Wilson, A. W.; Sun, P.; Bury, S. (2003) Methods for Wastewater Characterization in Activated Sludge Modelling; Water Environment Research Foundation: Alexandria, Virginia. Metcalf and Eddy, Inc. (2003) Wastewater Engineering: Treatment and Reuse, Tchobanoglous, G., Burton, F. L., Stensel, H. D. (Eds.); McGraw-Hill: New York. Michelet, F.; Jolly, M.; Chan, T.; Rogalla, F. (2005) Troubleshooting SAF and BAF Biofilm Reactors on Full Scale. Proceedings of the 78th Annual Water Environment Federation Technical Exposition and Conference, Washington, D.C., Oct 29–Nov 2; Water Environment Federation: Alexandria, Virginia. Mokhayeri, Y.; Nichols, A.; Murthy, S.; Riffat, R.; Dold, P.; Takacs, I. (2006) Examining the Influence of Substrates and Temperature on Maximum Specific Growth Rate of Denitrifiers. Water Sci. Technol., 54, 155–162. Moore, R.; Quarmby, J.; Stephenson, T. (1999) Development of a Novel Lightweight Media for Biological Aerated Filters (BAFs). Proceedings of the BAF3 Conference, Cranfield University: Bedfordshire, United Kingdom. Newman, J.; Occiano, V.; Appleton, R.; Melcer, H.; Sen, S.; Parker, D.; Langworthy, A.; Wong, P. (2005) Confirming BAF Performance for Treatment of CEPT
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Effluent on a Space Constrained Site. Proceedings of the 78th Annual Water Environment Federation Technical Exposition and Conference, Washington, D.C., Oct 29–Nov 2; Water Environment Federation: Alexandria, Virginia. Nichols, A.; Hinojosa, J.; Riffat, R.; Dold, P.; Takacs, I.; Bott, C.; Bailey, W.; Murthy, S. (2007) Maximum Methanol-Utilizer Growth Rate: Impact of Temperature on Denitrification. Proceedings of the 80th Annual Water Environment Federation Technical Exposition and Conference, San Diego, California, Oct 13–17; Water Environment Federation: Alexandria, Virginia. Nicolavcic, B. (2002) Nitrogen Removal in Biofilms (Stickstoffentfernung in Biofiltern). Ph.D. thesis, Vienna University of Technology, Wiener Mitteilungen, Wasser-Abwasser-Gewässer, No. 172 (in German). Ninassi, M. V.; Peladan, G.; Pujol, R. (1998) Pre-Denitrification of Municipal Wastewater: The Interest of Up-Flow Biofiltration. Proceedings of the 71st Annual Water Environment Federation Technical Exposition and Conference, Orlando, Florida, Oct 3–7; Water Environment Federation: Alexandria, Virginia. Odegaard, H. (2005) Combining CEPT and Biofilm Systems. Proceedings of the IWA Specialized Conference on Nutrient Management in Wastewater Treatment Processes and Recycle Streams, Krakow, Poland, Sept 19–21; International Water Association: London, United Kingdom. Parkson (2004) E-mail communication with Miguel Gutierrez regarding loading rates for DynaSand filters operating for post-denitrification, May. Payraudeau, M.; Le Tallec, X. (2003) Method for Treating an Effluent Using Simultaneous Nitrification/Denitrification in a Biological Filter. U.S. Patent 6,632,365, Oct 14. Pearce, P. A. (1996) Optimisation of Biological Aerated Filters. Proceedings of the BAF2 Conference, Cranfield University: Bedfordshire, United Kingdom. Peladan, J.; Lemmel, G.; Tarallo, H.; Tattersall, S.; Pujol, R. (1997) A New Generation of Upflow Biofilters with High Water Velocities. Proceedings of the International Conference on Advanced Wastewater Treatment Processes, Leeds, United Kingdom, Sept 8–11; Aqua Enviro Ltd.: Wakefield, United Kingdom. Peladan, J-G.; Lemmel, H.; Pujol, R. (1996) High Nitrification Rate with Upflow Biofiltration. Water Sci. Technol., 34, 347–353.
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Phipps, S. D.; Love, N. G. (2001) Quantifying Particle Hydrolysis and Observed Heterotrophic Yield for a Full-Scale Biological Aerated Filter. Proceedings of the 74th Annual Water Environment Federation Technical Exposition and Conference, Atlanta, Georgia, Oct 13–17; Water Environment Federation: Alexandria, Virginia. Pickard, D. W.; Bizzarri, R. E.; Wilson, T. E. (1985) Six Years of Successful Nitrogen Removal at Tampa, Florida. Proceedings of the 58th Annual Water Environment Federation Technical Exposition and Conference, Kansas City, Missouri, Oct 6–10; Water Environment Federation: Alexandria, Virginia. Pujol, R.; Hamon, M.; Kandel, X.; Lemmel, H. (1994) Biofilter: Flexible, Reliable Biological Filters. Water Sci. Technol., 29, 33–38. Pujol, R.; Tarallo, S. (2000) Total Nitrogen Removal in Two-Step Biofiltration. Water Sci. Technol., 41, 65–68. Redmon, D. T.; Boyle, W. C.; Ewing, L. (1983) Oxygen Transfer Efficiency Measurements in Mixed Liquor Using Off-Gas Techniques. J. Water Pollut. Control Fed., 55, 1338–1347. Reiber, S. H.; Stensel, H. D. (1985) Oxygen Transfer in Fixed Film Systems. J. Water Pollut. Control Fed., 57, 135–140. Roennefahrt, K. W. (1986) Nitrate Elimination with Heterotrophic Aquatic Microorganisms in Fixed-Bed Systems with Buoyant Carriers. Aqua, 5, 283–285. Rogalla, F. (2004) 21 Years of Full Scale BAF: Grown-Up Technology or Adolescent Adventure? Water Sci. Technol., 49, 29–36. Rogalla, F.; Bourbigot, M-M. (1990) New Developments in Complete Nitrogen Removal with Innovative Biological Reactors. Water Sci. Technol., 22, 273–280. Rogalla, F.; Chan, T. F.; Michelet, F. (2005) BAF, SAF, and DBF: Challenges and Experiences. Proceedings of the Conference on Design and Operation of Activated Sludge and Biofilm Systems, Horan, N., Ed.; Aqua Enviro Ltd.: Wakefield, United Kingdom. Rogalla, F.; Ravarini, P.; DeLarminat, G.; Courtelle, J. (1990) Large Scale Biological Nitrate and Ammonia Removal. J. Inst. Water Environ. Manage., 4, 319–329. Rogalla, F.; Sibony, J. (1992) Biocarbone Aerated Filters—Ten Years After: Past, Present and Plenty of Potential. Water Sci. Technol., 26, 2043–2048.
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Rother, E. (2005) Optimising the Design and Operation of BAF Processes for Municipal Wastewater Treatment. Ph.D. Dissertation, Darmstadt University of Technology, Schriftenreihe WAR, Band 163, Darmstadt, Germany. Rundle, H. (2009) Good Practice in Water and Environmental Management: Biological and Submerged Aerated Filters. Chartered Institution of Water and Environmental Management (CIWEM), Aqua Enviro Technology Transfer: Wakefield, United Kingdom. Ryhiner, G.; Birou, B.; Gros, H. (1992) The Use of Submerged Structured Packings in Biofilm Reactors for Wastewater Treatment. Water Sci. Technol., 26, 723–731. Sagberg, P.; Dauthille, P.; Hamon, M. (1992) Biofilm Reactors: A Compact Solution for Upgrading of Waste Water Treatment Plants. Water Sci. Technol., 26, 733–742. Savage, E. S. (1983) Biological Denitrification Deep Bed Filters. Presented at the Filtration Society Filtech Conference; Filtration Society: London, United Kingdom. Schauer, P.; Rectanus, R.; deBarbadillo, C.; Barton, D.; Gebbia, R.; Boyd, B.; McGehee, M. (2006) Pilot Testing of Upflow Continuous Backwash Filters for Tertiary Denitrification and Phosphorus Removal. Proceedings of the 79th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 21–25; Water Environment Federation: Alexandria, Virginia. Schlegel, S.; Teichgraber, B. (2000) Operational Results and Experience with Submerged Fixed-Film Reactors in the Pre-Treatment of Industrial Effluents. Water Sci. Technol., 41, 453–459. Shepherd, D.; Young, P.; Hobson, J. (1997) Biological Aerated Filters and Lamella Separators; Evaluation of Current Status. Water Research Center Report No PT2061; WRc: Swindon, United Kingdom. Sibony, J. (1982) Development of Aerated Biological Filters for the Treatment of Waste and Potable Water. Presented at the Integrated Waste Services Association Conference, Zurich, Switzerland. Slack, D., Severn Trent, Coventry, United Kingdom (2004) Personal communication.
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Smith, A. J.; Edwards, W.; Hardy, P.; Kent, T. (1999) BAF’s Get Media Attention. Proceedings of the BAF3 Conference, Cranfield University: Bedfordshire, United Kingdom. Smith, A. J.; Hardy, P. J. (1992) High-Rate Sewage Treatment Using Biological Aerated Filters. J. Inst. Water Environ. Manage., 6, 179–193. Springer, A.; Green, S. (2005) Colne Bridge BAFF Process Improvements. Proceedings of the Conference on the Design and Operation of Activated Sludge and Biofilm Systems; Horan, N., Ed.; Aqua Enviro Ltd.: Wakefield, United Kingdom. Stensel, H. D.; Brenner, R. C.; Lubin, G. (1984) Aeration Energy Requirements in Sparged Fixed Film Systems. Proceedings of the International Biological Fixed Film Conference, Washington, D.C., July 10–12; U.S. Environmental Protection Agency: Washington, D.C. Stenstrom, M. K.; Rosso, D.; Melcer, H.; Appleton, R.; Occiano, V.; Langworthy, A.; Wong, P. (2008) Oxygen Transfer in a Full-Depth Biological Aerated Filter. Water Environ. Res., 80, 663–671. Stephenson, T. (1996) Development of a Recirculating Plastic Media Biological Aerated Filter (REBAF). Proceedings of the BAF2 Conference; Cranfield University: Bedfordshire, United Kingdom. Toettrup, H.; Rogalla, F.; Vidal, A.; Harremoes, P. (1994) The Treatment Trilogy of Floating Filters: From Pilot to Prototype to Plant. Water Sci. Technol., 29, 23–32. Tolley Process Engineering Ltd. (1981) Apparatus for Agitating and/or Aerating Liquids. Br. Patent GB2, 069, 353. Tschui, M.; Boller, M.; Gujer, W.; Eugster, J.; Mäder, C. (1993) Tertiary Nitrification in Aerated Biofilm Reactors. Proceedings of the European Water Filtration Congress, Ostend, Belgium, March 15–16. Tschui, M.; Boller, M.; Gujer, W.; Eugster, C.; Mäder, C.; Stengel, C. (1994) Tertiary Nitrification in Aerated Biofilters. Water Sci. Technol., 29 (10-11), 53–60. U.S. Environmental Protection Agency (1993) Nitrogen Control Manual, EPA625/R-93-010; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C.
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Wanner, O.; Gujer, W. (1985) Competition in Biofilms. Water Sci. Technol., 17 (2-3), 27–44. Water Environment Federation (1998) Biological and Chemical Systems for Nutrient Removal, Special Publication. Water Environment Federation: Alexandria, Virginia. Whitaker, J.; Evans, I. D.; Cantwell, A. D. C. (1993) Filtration Apparatus and Method. Br. Patent GB2, 260, 275.
Chapter 8
New and Emerging Fixed-Film Technologies 1.0 INTRODUCTION
2.4 Internal Circulation Reactor
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2.0 BIOFILM REACTORS WITH SUSPENDED CARRIERS OR GRANULES 402 2.1 Biofilm Airlift Suspension Reactor
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2.2 Upflow Anaerobic Sludge Blanket
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2.3 Expanded Granular Sludge Blanket
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3.0 ANAMMOX BIOFILM REACTORS
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4.0 MEMBRANE BIOFILM REACTORS
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5.0 REFERENCES
408
1.0 INTRODUCTION This chapter highlights fixed-film processes that are new, emerging, or not wellknown or widely used in the United States. A significant, relatively new development within fixed-film technologies is the use of granular sludge reactors, in which naturally formed sludge granules behave similarly to suspended-media biofilms. An interesting new fixed-film application is the anaerobic ammonium oxidation, or anammox, process, whereby nitrogen is removed from high-strength ammonium wastewaters anaerobically and without the need for a carbon source. Another new fixed-film technology is the membrane biofilm reactor, where membranes are used to deliver gaseous substrates, such as oxygen or hydrogen, to a biofilm through the attachment surface. Each of these technologies is described in detail in the following sections. 401
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Biofilm Reactors
2.0 BIOFILM REACTORS WITH SUSPENDED CARRIERS OR GRANULES Biofilm processes are ideal for obtaining high biomass concentrations without a downstream separation and recycle process (Nicolella et al. 2000a). Biofilm processes may be based on films attached to an immobile surface (static biofilms) (e.g., a trickling filter [Chapter 3]) or attached to carriers (e.g., moving-bed biofilm reactors [Chapter 5] and biological aerated filters [Chapter 7]). Whereas static biofilms have relatively low specific surface areas, biofilm carriers dramatically increase the reactor specific surface area, decreasing mass-transfer limitations and allowing high-rate applications with compact reactors. A special case of mobile carriers is the sludge granule (Lettinga et al., 1980). Granules are very large flocs, in diameters ranging from 1 to 3 mm, and have excellent settling properties, similar to carriers. Therefore, granular sludge systems can be considered a type of suspended carrier (Nicolella et al. 2000b). Several types of suspended carrier biofilm processes, shown schematically in Figure 8.1 are described below.
2.1 Biofilm Airlift Suspension Reactor Biofilm airlift suspension (BAS) reactors were developed in the Netherlands in the late 1980s for aerobic wastewater treatment, including the oxidation of biochemical oxygen demand (BOD), sulfide, and ammonia (Heijnen et al., 1993). Anaerobic versions of BAS reactors, called gas-lift reactors, use gases such as methane or nitrogen gas instead of air to provide the circulation. These gases typically are degradation byproducts formed in the reactor. The BAS reactors typically are in a tower configuration, which is divided vertically into riser and downcomer sections. Air is introduced at the bottom of the riser section, traverses the length of the reactor, and exits at the top. The upward bubble movement provides mixing, and sludge granules form, in response to the high upflow velocities, which wash out smaller particles. Commercial versions of the BAS process include CIRCOX, marketed in Europe by Paques B.V. (Balk, Netherlands) and marketed in the United States by Siemens Water Technologies (Warrendale, Pennsylvania); Turbo-Flo, developed in France by Lyonnaise des Eaux, currently part of Suez (Paris, France) (Lazarova and Manem, 1996; Mousseau et al., 1998); and BIOLIFT (Badot et al., 1994), developed in France by OTV, S.A, currently part of Veolia Water Systems (Saint-Maurice, France). As an
New and Emerging Fixed-Film Technologies
(a)
Effluent Gas collection dome Sludge blanket
Influent
Gas
(b)
Gas
Effluent
Gas-liquid-solid separation Expanded granular sludge blanket Influent
(d)
Gas
Gas-liquid separator
(c)
Annular space
Settling space
Gas Effluent Second separator
Effluent
First separator Expanded bed Influent
Air
Influent distribution system
Figure 8.1 Reactor configurations: (a) UASB, (b) EGSB, (c) BAS, and (d) internal circulation (modified from Nicolella et al., 2000b).
example, the CIRCOX system has a high loading capacity (4 to 10 kg COD/m3·d), short hydraulic retention times (0.5 to 4 hours), high biomass settling velocities (50 m/h), and high biomass concentrations (15 to 30 g/L) (Frijters et al., 2000; Nicolella et al., 2000b). Given the long solids retention time, nitrification is achieved easily with this process. A modified CIRCOX was developed to include an anoxic compartment for denitrification and was tested at pilot- and full-scale (Frijters, et al., 2000). The volumetric loading for nitrogen was 1 to 2 kg N/m3·d, and the denitrification efficiency was over 90%.
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2.2 Upflow Anaerobic Sludge Blanket The orginal upflow anaerobic sludge blanket was developed in the late 1970s in the Netherlands (Lettinga et al., 1980). Conceptually, the upflow anaerobic sludge blanket (UASB) is similar to a fluidized-bed reactor, but, instead of suspending an inert medium with the upward flow, large sludge granules form and constitute a suspended blanket. Typically, the reactor is in the form of a vertical cylinder. The influent enters the bottom of the reactor and passes through the bed, where soluble chemical oxygen demand (COD) is transformed into biogas—mainly methane. Gas bubbles adsorb to the granules, causing them to rise to the gas separator or settler, where the gas is dislodged and the granules return to the blanket. The upward and downward granule movement facilitates mixing. A key feature of UASBs is the formation of dense granular sludge. This provides very high solids concentrations and allows higher volumetric COD loadings than other anaerobic processes. Examples of UASB systems are BIOPAQ, marketed by Paques B.V, and the Biothane UASB, marketed by Biothane (Camden, New Jersey). Loading rates are in the range 10 to 15 kg COD/m3·d.
2.3 Expanded Granular Sludge Blanket The expanded granular sludge bed (EGSB) is a more recent modification of a UASB system, with a more elongated reactor and higher upflow velocities for liquid (10 m/h) and gas (7 m/h) (Seghezzo et al., 1998). The higher upflow velocity “expands” the granular sludge bed, improves internal mixing, and better utilizes the entire reactor volume. These reactors can be operated at much higher loading rates, from 15 to 30 kg COD/m3·d, making them suitable to treat high-strength industrial wastes. A disadvantage is that high upflow velocities preclude the removal of suspended solids. Numerous EGSB plants have been built by Biothane, which markets EGSB reactors as the Biobed process.
2.4 Internal Circulation Reactor The internal circulation reactor consists of two sequential UASB processes—one high rate and the second low rate (Pereboom et al., 1994). The reactor is in a tower configuration, where the lower part contains the high-rate reactor, and the upper part contains the low-rate reactor. The low-rate reactor polished the effluent from the high-rate reactor. In the lower tower, an expanded bed of granular sludge converts COD to biogas. The gas is collected in a separator and lifts water and sludge
New and Emerging Fixed-Film Technologies
to the upper compartment, where the gas is separated and the sludge is returned via a downer pipe. Paques B.V. has several internal circulation plants in operation (Nicolella et al., 2000b).
3.0 ANAMMOX BIOFILM REACTORS The anammox process is a novel technology that removes nitrogen from wastewaters, based on the unique metabolism of anammox bacteria (Strous et al., 1999). The process was developed by Delft University of Technology (Delft, Netherlands) and Paques B.V. anammox bacteria, which are chemolithoautotrophs and members of the Planctomycetales order, use ammonium as an electron donor and nitrite as an acceptor, producing dinitrogen gas without the need for a carbon source or electron donor (Strous et al., 1999). Nitrate is produced as a byproduct, at approximately 12% of the influent nitrogen concentration. The process is ideal for high-strength ammonium wastes (>0.2 g N/L) that are low in organic carbon (C:N ratio lower than 0.15), such as digester supernatant. This process is typically run in tandem with nitrite-producing processes, such as completely autotrophic nitrogen removal over nitrite (CANON) (Sliekers et al., 2003) and stable high ammonia removal over nitrite (SHARON) (van Kempen et al., 2001). Anammox bacteria are very slow growing, with a doubling time of approximately 11 days (Strous et al., 1998), but high volumetric loadings can be obtained using fixed-film anammox processes (Hippen et al., 2001). The anammox process has been studied with moving-bed biofilm reactors, rotating biological contactors, anaerobic biological filters, and granular sludge bioreactors (Abma et al., 2007). The anammox process is being marketed by Paques B.V., and several fullscale plants have been built and tested in Europe. The first was installed in 2002 at the wastewater treatment plant (WWTP) of the Waterboard Hollandse Deltain (Rotterdam, Netherlands), with a capacity of 500 kg N/d. Other plants are at foodprocessing, tanning, and semiconductor industries. At the Netherlands WWTP, the effluent from the sludge digester is dewatered, routed through an existing SHARON reactor, settled, and routed through the anammox reactor. The digester effluent contains 1000 to 1500 mg/L NH4+-N, and the effluent of the SHARON reactor contains equal amounts of ammonium (NH4+) and nitrite (NO2–). The configuration is similar to an internal circulation reactor (see above). It is a tower, where the influent is introduced at the bottom and the effluent leaves at the top. The influent is mixed at the bottom and then passed through a granular sludge bed, where most of the anammox
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activity takes place. Internal circulation is created by the produced nitrogen gas bubbles, which act as a gas lift. Nitrogen is removed at the top of the tower. A second compartment further polishes the effluent from the lower compartment, by removing the remaining NH4+ and NO2–. Given the slow growth rate of anammox bacteria and the lack of seed sludge, the expected startup time was 2 years. Because of operational difficulties, including problems with nitrite toxicity and sulfide inhibition, the actual startup time was 3.5 years. Startup of future plants should be faster, as existing plants will provide inocula. A key aspect of the anammox process is the formation of sludge granules, which greatly increase the biomass concentration. In full-scale tests, loading rates of up to 10 kg N/m2·d were achieved. The effluent NH4+ concentration was 60 to 130 mg N/L, while NO2– was 5 to 10 mg N/L, and NO3– was approximately 130 mg N/L (Abma et al., 2007). A second type of anammox process, the de-ammoniafication (DEMON) process, was developed to carry out partial nitrification and anammox in a single reactor (Sliekers et al., 2003). This process was tested at full-scale in Austria using an sequencing batch reactor process (Wett, 2006). In this system, special care is needed to prevent excessive dissolved oxygen concentrations, which increase nitrification rates and lead to NO2– toxicity.
4.0 MEMBRANE BIOFILM REACTORS Membranes have long been used for gas separation and gas transfer to liquids. In the late 1980s, researchers found that these same membranes could deliver a gaseous substrate, such as oxygen or hydrogen, to a biofilm naturally forming on the membrane’s outer surface ((Lee and Rittmann, 2000; Syron and Casey, 2008; Timberlake et al., 1988). When used to deliver oxygen, some researchers called them membrane-aerated bioreactors (MABRs) (Brindle and Stephenson, 1996a; Casey et al., 1999), but more generally can be called membrane biofilm reactors (MBfRs). For consistency, in this chapter, the membrane biofilm applications will be referred to as MBfRs. Hollow-fiber membranes commonly are used in MBfRs because, with outside diameters as small as 0.1 mm, they can provide high specific surface areas. Membrane sheets also are used. Microporous membranes typically provide much higher gas transfer rates than dense materials as gas molecules diffuse much more quickly through dry pores than through liquid-filled pores (Yang and Cussler, 1986). Hydrophobic materials are used to prevent the pores from wetting. Smaller pore sizes allow higher transmembrane pressures without gas bubbling, and a membrane
New and Emerging Fixed-Film Technologies
with a thin, dense layer also can discourage bubbling. Unlike membrane bioreactors (MBRs), where the membranes act as filters, the MBfR pores simply convey gas and therefore do not become fouled with solids or bacteria. When hollow fiber membranes are used, they often are collected into a gassupplying manifold at one end, while the opposite end may be open or sealed. Sealed ends are typically used when supplying gases are toxic, flammable, or expensive. Sealed end processes are highly efficient, as 100% of the gas supplied to the MBfR passes into the biofilm. The gas flux to the biofilm can be modulated by controlling the gas-supply pressure. Oxygen-based MBfRs have been studied since the late 1980s for wastewater applications, where the fibers served as a “passive” aeration device and concurrently could achieve BOD removal, nitrification, and denitrification (Brindle and Stephenson, 1996b; Brindle et al., 1998; Semmens et al., 2003; Suzuki et al., 1993; Timberlake et al., 1988). Nitrification typically occurs in the inner portions of the biofilm, close to the airor oxygen-filled membrane, and denitrification and BOD removal occur in the outer portions, when the bulk-liquid dissolved oxygen concentrations are low (Downing and Nerenberg, 2008b; Downing et al., 2010; Schramm et al., 2000; Semmens et al., 2003). Hydrogen-based MBfRs have been studied for the reduction of nitrate (Ergas and Reuss, 2001; Lee and Rittmann, 2002) and other oxidized contaminants in drinking water (Nerenberg and Rittmann, 2004). Pilot-scale tests have been conducted with MBfRs for nitrate and perchlorate from groundwater (Adham et al., 2004), showing high removals. A number of researchers have investigated air- or oxygen-based MBfRs for nitrogen removal from wastewater, mostly at the bench-scale (Hibiya et al., 2003; Pankhania et al., 1999; Semmens et al., 2003). Pilot-scale studies also have been conducted on air- and pure-oxygen-based MBfRs for wastewater treatment. For example, a pilot-scale MBfR was found to be effective in removing BOD in high-strength brewery wastewater, with organic removal rates of 27 kg COD/m3·d (Brindle et al., 1999). A pilot-scale MBfR was studied for concurrent removal of COD and total nitrogen using hollow-fiber and sheet membranes, and high removal efficiencies were found, but the performance decreased with time, as a result of excessive biofilm accumulation (Semmens, 2005). Bench- and pilot-scale tests also have been carried out on a hybrid (suspended- and attached-growth) MBfR for removal of BOD and nitrogen from wastewater (Downing and Nerenberg, 2007b, 2008a). This process is similar to a cord-type integrated fixed-film activated sludge system, where hollow-fiber membranes instead of cords are retrofitted into an activated sludge tank.
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Several researchers have used methane-based MBfRs to co-metabolically degrade trichloroethylene (Grimberg et al., 2000). Others have used MBfRs to treat gaseous contaminants that are passed through the lumen. For example, Min et al. (2002) used a nitrifying MBfR to oxidize nitric oxide (NO) from combustion gases to NO3–. Hydrogen (H2)-based MBfRs have been used to reduce a wide number of oxidized contaminants, including arsenate (Chung, Li, and Rittmann, 2006), bromate (Downing and Nerenberg, 2007a), chromate (Chung, Nerenberg, and Rittmann, 2006b), selenate (Chung, Nerenberg, and Rittmann, 2006a), and trichloroethane (Chung and Rittmann, 2007). Many other applications are possible and are likely to be developed as this technology becomes more well-known. The H2-based MBfRs are being developed commercially in the United States by Applied Process Technology, Inc. (Cincinnati, Ohio), although no full-scale applications have been built to date. More research is needed to determine the optimal membrane materials, diameter, packing density, and bulk liquid mixing strategy. A key need is for effective management of biofilm growth, as excessive growth reduces reactor efficiency.
5.0 REFERENCES Abma, W.; Schultz, C. E.; Mulder, J. W.; ven der Star, W. R. L.; Strous, M.; Tokutomi, T.; van Loosdrecht, M. C. M. (2007) Full-Scale Granular Anammox Process. Water Sci. Technol., 55 (8–9), 27–33. Adham, S.; Gillogly, T.; Nerenberg, R.; Lehman, G.; Rittmann, B. E. (2004) Membrane Biofilm Reactor Process for Nitrate and Perchlorate Removal #2804; Water Research Foundation: Denver, Colorado. Badot, R.; Coulom, T.; Delongeaux, N.; Badard, M.; Sibony, J. (1994) A FluidizedBed Reactor—The Biolift Process. Water Sci. Technol., 29 (10–11), 329–338. Brindle, K.; Stephenson, T. (1996a) The Application of Membrane Biological Reactors for the Treatment of Wastewaters. Biotechnol. Bioeng., 49 (6), 601–610. Brindle, K.; Stephenson, T. (1996b) Nitrification in a Bubbleless Oxygen Mass Transfer Membrane Bioreactor. Water Sci. Technol., 34 (9), 261–267. Brindle, K.; Stephenson, T.; Semmens, M. J. (1998) Nitrification and Oxygen Utilisation in a Membrane Aeration Bioreactor. J. Membr. Sci., 144 (1–2), 197–209.
New and Emerging Fixed-Film Technologies
Brindle, K.; Stephenson, T.; Semmens, M. J. (1999) Pilot-Plant Treatment of a High-Strength Brewery Wastewater Using a Membrane-Aeration Bioreactor. Water Environ. Res., 71, 1197–1204. Casey, E.; Glennon, B.; Hamer, G. (1999) Review of Membrane Aerated Biofilm Reactors. Res. Conserv. Recycl., 27 (1–2), 203–215. Chung, J.; Li, X. H.; Rittmann, B. E. (2006) Bio-Reduction of Arsenate Using a Hydrogen-Based Membrane Biofilm Reactor. Chemosphere, 65 (1), 24–34. Chung, J.; Nerenberg, R.; Rittmann, B. E. (2006a) Bioreduction of Selenate Using a Hydrogen-Based Membrane Biofilm Reactor. Environ. Sci. Technol., 40, 1664–1671. Chung, J.; Nerenberg, R.; Rittmann, B. E. (2006b) Bio-Reduction of Soluble Chromate Using a Hydrogen-Based Membrane Biofilm Reactor. Water Res., 40, 1634–1642. Chung, J.; Rittmann, B. E. (2007) Bio-Reductive Dechlorination of 1,1,1-Trichloroethane and Chloroform Using a Hydrogen-Based Membrane Biofilm Reactor. Biotechnol. Bioeng., 97 (1), 52–60. Downing, L.; Downing, L. S.; Bibby, K. J.; Esposito, K.; Fascianella, T.; Tsuchihashi, R.; Nerenberg, R. (2010) Nitrogen Removal from Wastewater Using a Hybrid Membrane-Biofilm Process: Pilot-Scale Studies. Water Environ. Res., 82, 195–201. Downing, L.; Nerenberg, R. (2007a) Kinetics of Microbial Bromate Reduction in a Hydrogen-Oxidizing, Denitrifying Biofilm Reactor. Biotechnol. Bioeng., 98 (3), 543–550. Downing, L.; Nerenberg, R. (2007b) Performance and Microbial Ecology of the Hybrid Membrane Biofilm Process (HMBP) for Concurrent Nitrification and Denitrification of Wastewater. Water Sci. Technol., 55 (8–9), 355–362. Downing, L.; Nerenberg, R. (2008a) Effect of Bulk Liquid BOD Concentration on Activity and Microbial Community Structure of a Nitrifying, MembraneAerated Biofilm. Appl. Microbiol. Biotechnol., 81, 153–162. Downing, L.; Nerenberg, R. (2008b) Effect of Oxygen Gradients on the Activity and Microbial Community Structure of a Nitrifying, Membrane-Aerated Biofilm. Biotechnol. Bioeng., 101, 1193–1204.
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Ergas, S. J.; Reuss, A. F. (2001) Hydrogenotrophic Denitrification of Drinking Water Using a Hollow Fibre Membrane Bioreactor. J. Water Supply Res. Technol. Aqua, 50 (3), 161–171. Frijters, C.; Vellinga, S.; Jorna, T.; Mulder, R. (2000) Extensive Nitrogen Removal in a New Type of Airlift Reactor. Water Sci. Technol., 41 (4–5), 469–476. Grimberg, S. J.; Rury, M. J.; Jimenez, K. M.; Zander, A. K. (2000) Trinitrophenol Treatment in a Hollow Fiber Membrane Biofilm Reactor. Water Sci. Technol., 41 (4–5), 235–238. Heijnen, J. J.; Vanloosdrecht, M. C. M.; Mulder, R.; Weltevrede, R.; Mulder, A. (1993) Development and Scale-Up of an Aerobic Biofilm Airlift Suspension Reactor. Water Sci. Technol., 27 (5–6), 253–261. Hibiya, K.; Terada, A.; Tsuneda, S.; Hirata, A. (2003) Simultaneous Nitrification and Denitrification by Controlling Vertical and Horizontal Microenvironment in a Membrane-Aerated Biofilm Reactor. J. Biotechnol., 100 (1), 23–32. Hippen, A.; Helmer, C.; Kunst, S.; Rosenwinkel, K. H.; Seyfried, C. F. (2001) Six Years’ Practical Experience with Aerobic/Anoxic Deammonification in Biofilm Systems. Water Sci. Technol., 44 (2–3), 39–48. Lazarova, V.; Manem, J. (1996) An Innovative Process for Waste Water Treatment: The Circulating Floating Bed Reactor. Water Sci. Technol., 34 (9), 89–99. Lee, K.-C.; Rittmann, B. E. (2000) A Novel Hollow-Fiber Membrane Biofilm Reactor for Autohydrogenotrophic Denitrification of Drinking Water. Water Sci. Technol., 41 (4–5), 219–226. Lee, K.-C.; Rittmann, B. E. (2002) Applying a Novel Autohydrogenotrophic Hollow-Fiber Membrane Biofilm Reactor for Denitrification of Drinking Water. Water Res., 36 (8), 2040–2052. Lettinga, G.; Vanvelsen, A. F. M.; Hobma, S. W.; Dezeeuw, W.; Klapwijk, A. (1980) Use of the Upflow Sludge Blanket (USB) Reactor Concept for Biological Wastewater-Treatment, Especially for Anaerobic Treatment. Biotechnol. Bioeng., 22(4), 699–734. Min, K. N.; Ergas, S. J.; Harrison, J. M. (2002) Hollow-Fiber Membrane Bioreactor for Nitric Oxide Removal. Environ. Eng. Sci., 19, 575–583.
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Mousseau, F.; Liu, S. X.; Hermanowicz, S. W.; Lazarova, V.; Manem, J. (1998) Modeling of Turboflo—A Novel Biofilm Reactor for Wastewater Treatment. Water Sci. Technol., 37 (4–5), 177–181. Nerenberg, R.; Rittmann, B. E. (2004) Reduction of Oxidized Water Contaminants with a Hydrogen-Based, Hollow-Fiber Membrane Biofilm Reactor. Water Sci. Technol., 49 (11–12), 223–230. Nicolella, C.; van Loosdrecht, M. C. M.; Heijnen, S. J. (2000a) Particle-Based Biofilm Reactor Technology. Trends Biotechnol., 18 (7), 312–320. Nicolella, C.; van Loosdrecht, M. C. M.; Heijnen, J. J. (2000b) Wastewater Treatment with Particulate Biofilm Reactors. J. Biotechnol., 80 (1), 1–33. Pankhania, M.; Brindle, K.; Stephenson, T. (1999) Membrane Aeration Bioreactors for Wastewater Treatment: Completely Mixed and Plug-Flow Operation. Chem. Eng. J., 73 (2), 131–136. Pereboom, J. H. F.; Vereijken, T. (1994) Methanogenic Granule Development in Full-Scale Internal Circulation Reactors. Water Sci. Technol., 30 (8), 9–21. Schramm, A.; De Beer, D.; Gieseke, A.; Amann, R. (2000) Microenvironments and Distribution of Nitrifying Bacteria in a Membrane-Bound Biofilm. Environ. Microbiol., 2 (6), 680–686. Seghezzo, L.; Zeeman, G.; van Lier, J. B.; Hamelers, H. V. M.; Lettinga, G. (1998) A Review: The Anaerobic Treatment of Sewage in UASB and EGSB Reactors. Bioresour. Technol., 65 (3), 175–190. Semmens, M. J. (2005) Membrane Technology: Pilot Studies of Membrane-Aerated Bioreactors; Water Environment Research Foundation: Alexandria, Virginia. Semmens, M. J.; Dahm, K.; Shanahan, J.; Christianson, A. (2003) COD and Nitrogen Removal by Biofilms Growing on Gas Permeable Membranes. Water Res., 37, 4343–4350. Sliekers, A. O.; Third, K. A.; Abma, W.; Kuenen, J. G.; Jetten, M. S. M. (2003) Canon and Anammox in a Gas-Lift Reactor. FEMS Microbiol. Lett., 218 (2), 339–344. Strous, M.; Fuerst, J. A.; Kramer, E. H. M.; Logemann, S.; Muyzer, G.; van de Pas-Schoonen, K. T.; Webb, R.; Kuenen, J. G.; Jetten, M. S. M. (1999) Missing Lithotroph Identified as New Planctomycete. Nature, 400, 446–449.
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Strous, M.; Heijnen, J. J.; Kuenen, J. G.; Jetten, M. S. M. (1998) The Sequencing Batch Reactor as a Powerful Tool for the Study of Slowly Growing Anaerobic Ammonium-Oxidizing Microorganisms. Appl. Microbiol. Biotechnol., 50, 589–596. Strous, M.; Kuenen, J. G.; Jetten, M. S. M. (1999) Key Physiology of Anaerobic Ammonium Oxidation. Appl. Environ. Microbiol., 65, 3248–3250. Suzuki, Y.; Miyahara, S.; Takeishi, K. (1993) Oxygen-Supply Method Using Gas-Permeable Film for Waste-Water Treatment. Water Sci. Technol., 28 (7), 243–250. Syron, E.; Casey, E. (2008) Membrane-Aerated Biofilms for High Rate Biotreatment: Performance Appraisal, Engineering Principles, and Development Requirements. Environ. Sci. Technol., 42, 1833–1844. Timberlake, D.; Strand, S.; Williamson, K. (1988) Combined Aerobic Heterotrophic Oxidation, Nitrification and Denitrification in a Permeable-Support Biofilm. Water Res., 22, 1513–1517. van Kempen, R.; Mulder, J. W.; Uijterlinde, C. A.; Loosdrecht, M. C. M. (2001) Overview: Full Scale Experience of the SHARON Process for Treatment of Rejection Water of Digested Sludge Dewatering. Water Sci. Technol., 44 (1), 145–152. Wett, B. (2006) Solved Upscaling Problems for Implementing Deammonification of Rejection Water. Water Sci. Technol., 53 (12), 121–128. Yang, M.-C.; Cussler, E. L. (1986) Designing Hollow-Fiber Contactors. Am. Inst. Chem. Eng. J., 32, 1910–1916.
Chapter 9
Clarification 414
3.6.5 Floor Slope
2.0 Solids-Separation Choices
416
3.6.6 Effluent Weir and Launder 440
3.0 Design Approach
417
3.6.7 Sludge Collectors 443
1.0 Introduction
3.6.8 Sludge Hopper
438
445
3.1 Types of Settling Regimes
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3.1.1 Type I
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3.7 Rectangular versus Circular Clarifiers
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3.1.2 Type II
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3.8 Design Example
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3.1.3 Type III
419
3.1.4 Type IV
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3.9 Clarifier Following MovingBed Biofilm Reactor, Trickling Filter, Rotating Biological Contactor, and Biotower 448
3.2 Special Considerations for Nutrient Removal Sludges 419 3.4 Wastewater Flocculation 422
3.9.1 Secondary (Integrated Fixed-Film Activated Sludge) Clarifiers 451
3.5 Flocculation Criteria
3.9.2 Sludge Hopper
454
3.9.3 Process Performance
455
3.3 Clarifier Enhancements 420
424
3.6 Clarifier Design Details 429 3.6.1 Influent Column 431 3.6.2 Energy-Dissipating Inlet 431 3.6.3 Feed Well (Flocculating Type) 433 3.6.4 Side Water Depth, Clear Water Zone, and Overflow Rate 435
3.10 Other Considerations
458
3.10.1 Modeling
458
3.10.2 Interaction with Other Facilities 458 3.10.3 International Practices 4.0 REFERENCES
413
458 459
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1.0 Introduction Fixed-film reactors convert soluble biochemical oxygen demand (BOD) (for carbonaceous, aerobic systems) and grow biological solids for nitrifying and denitrifying systems. These solids contain BOD, phosphorus, and nitrogen, which must be removed to meet treatment objectives. Various solids-separation devices are now available. In the past, and to a limited extent today, the performance of fixed-film reactors (excluding activated biofilters [BAFs]) reflects the designs of clarifiers recommended for trickling filters (TFs). The Recommended Standards for Sewage Works (Great Lakes–Upper Mississippi River Board of State Sanitary Engineers, 1971) suggested that a 2.1-m (7.0-ft) side water depth (SWD) and an average overflow rate (OFR) of 1.7 m/h (1000 gpd/sq ft) represented an adequate design for trickling filter clarifiers. These clarifier recommendations can and did result in many trickling filters being unable to meet secondary standards of 30 mg/L 5-day BOD (BOD5) and total suspended solids (TSS). Similar design criteria also would have resulted in failure for many activated sludge (AS) plants. The updated Recommended Standards for Sewage Works (Great Lakes–Upper Mississippi River Board of State Sanitary Engineers, 1990) has increased the side water depth to a minimum of 3.0 m (10 ft), which may not be adequate for higher or cost-effective OFRs. Many of the previous clarifier design limitations are recognized today. Experienced design engineers are providing up-to-date clarifier designs that feature a deep SWD; tangential inlets to large hydraulically flocculated feed wells (Parker, 1983; Walker Equipment Company, 1953); and proper effluent launder configuration, such as cantilevered inboard (Anderson, 1945; Parker, 1983) or wall-mounted launders with Lincoln baffles (Stukenberg et al., 1983) or Stamford baffles (Water Environment Federation, 1992) located below the effluent weirs or as part of the effluent launder. Combined with a follow-on short-term aeration or solids-contact basin, trickling filter effluent qualities of ≤10 mg/L carbonaceous BOD5 and TSS may be routinely produced when these concepts are incorporated to the clarifier design. For example, Parker (1983, 1991) and Parker and Stenquist (1986) reported on the performance of deep (4.9 to 6.1 m) flat-floor secondary suction clarifiers with large feed wells and internal mechanical flocculators. These units provided treatment capacity up to and exceeding an OFR of 2.0 m/h (1180 gpd/sq ft) without a significant loss in efficiency; that is, effluent TSS was maintained at approximately 10 mg/L. The authors of these papers use cantilevered inboard launders as part of their design philosophy to control the density currents and launder approach velocity.
Clarification
Other engineers advocate the use of the Lincoln and Stamford baffle arrangements (IWA, 1992; Stukenberg et al., 1983; Water Environment Federation, 2005; WEF et al., 2009). The maximum hydraulic capacity of the deep, relatively flat-floor, rapid sludge removal flocculating clarifiers is consistent with the depth versus limiting OFR formulations (IWA, 1992) presented later for primary, intermediate, and secondary clarifiers using scrapers and central draw-off. Any review of historical trickling filter data should consider the performance bias inherent in a poorly designed clarifier; for example, the effluent TSS will be higher, and the associated BOD5 will increase the effluent BOD5 proportionally, to BOD5 = SBOD5 + fb (TSS)
(9.1)
Where
fb = BOD5 content of effluent TSS (mg BOD5/mg TSS) and SBOD5 = soluble BOD5 (mg/L) measured on the filtrate from the TSS test.
BOD5 in TSS, mg BOD5/mg TSS
As with the activated sludge process, the effect of effluent TSS on effluent quality is a function of the BOD5 loading (kg/m3·d) and the temperature. Reduced organic loadings and higher operating temperatures will reduce the value of fb. The relationship between fb and the loading on the trickling filter was reported in a study by Brown and Caldwell (1978), and other additional data are shown in Figure 9.1.
0.8
0.6 Seattle Metro Malmö, Sweden Garland, Tex. S. Davis Co., Utah Stockton, Calif.
0.4
0.2
0
50
150 100 BOD5 Loading, Ib/1000 cu ft/day
200
Figure 9.1 BOD5 content of effluent versus TSS and trickling filter loading (lb/d/1000 cu ft × 0.016 02 = kg/m3·d).
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In the operating plant, the reduction of the oxygen demand of the TSS is influenced by wastewater temperature; cold wastewater will reduce the rate of endogenous decay. However, incubation at 20°C in the BOD5 test will accelerate the rate of organic matter oxidation, and, consequently, the reported BOD5 of the TSS would increase in the winter. That is,
Fb = fb20 (Θt-20)
(9.2)
Where Θ = 1.020 to 1.025. Colder water, which reduces the rate of soluble BOD5 removal and the settling velocity of particles, contributes to higher effluent BOD5 and TSS typically occurring in the colder periods of the year. Reduced bioactivity in the fixed bed also may result in less bioflocculation. The benefits of improving effluent TSS and BOD5 quality by follow-on contact or short-term aeration were observed during all seasons, especially in cold weather. The latter of these seasons supported the bioflocculation premise.
2.0 Solids-Separation Choices There are five general choices of solids separation equipment typically used in wastewater treatment.
(1) Conventional clarifiers, (2) High-rate clarifiers, (3) Dissolved air flotation thickeners (DAFs), (4) Filters, and (5) Membranes.
Conventional clarifiers are the types most commonly used with fixed-film reactors and the only type discussed in detail in this chapter. The DAFs are commonly used in Sweden following moving-bed biofilm reactors (MBBRs). In some tertiary applications, where concentrations of solids leaving the fixed reactor are low enough, only filters may be needed. The DC Water and Sewer Authority (Washington, D.C.) is considering the latter for tertiary MBBR application at its Blue Plains wastewater treatment plant (WWTP), for example. For information on high-rate clarifiers, including augmentation with lamella separators and chemicals, refer to Manual of Practice (MOP) FD-8, Clarifier Design (Water Environment Federation, 2005). For DAFs, filters,
Clarification
and membranes, see MOP 8, Design of Municipal Wastewater Treatment Plants (WEF et al., 2009). Filtration also is discussed in Chapter 10 of this manual. The BAFs are a unique form of fixed-film reactor, which combines the biological reactor with the solids-separation media, and generally are not followed by any solids separation. When they are, it is typically a downflow filter.
3.0 Design Approach The approach taken in this manual is the same as in MOP FD-8 (Water Environment Federation, 2005), though more details can be found in MOP FD-8. In that manual, a single, general approach for sizing all types of clarifiers is suggested, as described below.
(1) Characterize the settling velocity or settling velocity distribution of wastewater (discussed below). (2) Select design settling velocity, VD, in m/h (typically >1 m/h; see below for how to choose). (3) Calculate the ideal clarifier area, Aideal (in m2), as follows: Aideal = QR/(VD *24)
(9.3)
where QR = maximum wastewater flow to clarifier (m3/d). QR is typically the maximum daily flow for fixed-film systems.
(1) Determine the degree of non-ideality expected and express it as a design efficiency (DE). Design efficiency is the ratio of the clarifier area required by an ideal clarifier to that of a particular design. An ideal clarifier would have a design efficiency of 1.0, and, for example, the literature (Ozinzky et al., 1994; Watts et al., 1996) suggests that typical shallow circular clarifiers have a DE of approximately 0.7 to 0.8; Ekama et al. (1997) suggests that certain rectangular clarifier designs may have a design efficiency of 0.8 to over 1.0. In these cases, the “ideal clarifier” was one that performed according to one-dimensional flux theory. It is expected that, in the near future, vendors will include this ratio in their designs and documentation to support their claims (2) Determine design surface area, Ad : Ad = Aideal /DE
(9.4)
417
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Biofilm Reactors
where DE = design efficiency and is a characteristic of the particular clarifier design details. (3) Select depth and design details (inlet and outlet designs, baffling, collectors, etc.) to achieve most cost-effective design. Clarifiers treating effluents from most fixed-film reactors, including trickling filters, biotowers, rotating biological contactors (RBCs), and MBBRs, historically are categorized as type I or type II settlers. Details of this type classification may be found below. For these, VD is chosen by developing a distribution of settling velocities of the particles in the wastewater. This can be a cumulative frequency distribution (fraction of suspended solids settling faster than stated value). The VD is chosen to correspond to the percent removal desired (i.e., fraction of solids settling faster than VD). With these types of clarifiers, it is possible to increase the VD (i.e., reduce design clarifier area Ad), by providing flocculation and/or by adding chemicals and/or ballasting agents. Another option for these type clarifiers is to include tubes or plates to increase the settling area available for a given footprint. Details of how to do this are discussed in each appropriate chapter and in most detail in Chapter 3 of MOP FD-8 (Water Environment Federation, 2005). Clarifiers following integrated fixed-film activated sludge (IFAS) or TF/solids contact (SC) systems treat higher concentrations of suspended solids, which settle as a uniform mass at a uniform initial settling velocity (ISV). This traditionally is referred to as type III or zone settling (see below for more details of this type designation). These solids are called mixed-liquor suspended solids (MLSS). Here, VD = ISV and is a function of the biology (primarily how many and which type of filamentous organisms are in the MLSS) and the concentration of the MLSS . Typically, the mixed-liquor quality is represented by tests like sludge volume index (SVI) and/or settling constants, such as VO and k. Chapter 4 of Clarifier Design, MOP FD-8 (Water Environment Federation, 2005), describes methods to measure and estimate the ISV.
3.1 Types of Settling Regimes The following four types, or classifications, of settling are recognized:
(1) Type I, free or discrete particle settling; (2) Type II, flocculent settling; (3) Type III, hindered or zone settling; and (4) Type IV, compaction.
Clarification
These four regimes occur in most wastewater clarification processes to a greater or lesser degree, depending on the design criteria used and the operation of the units.
3.1.1 Type I Free or discrete particle settling typically occurs in grit chambers. In this type of settling, particles are assumed to settle independently and at a constant velocity as a function of Newton’s and Stake’s laws. The design of the type I basin should minimize turbulence in the clarification zone.
3.1.2 Type II Flocculent settling is defined as a condition where particles initially settle independently but agglomerate with time within the clarification unit. As a result, the settling velocity of the particle aggregates increases. No mathematical models are available to predict the settling velocity or changes in velocity with depth, and laboratory long (1.2 to 2.0 m)-column settling tests should be conducted to define the effects. Encouraging flocculation will improve the settling rate and reduce the overflow TSS.
3.1.3 Type III Hindered or zone settling is typical of the settling characteristics in the lower zones of an activated sludge clarifier or a gravity thickener, wherein the initial suspended solids concentrations are higher than in flocculent settling. As a result, the mass of solids form and settle as a zone or blanket, and the particles remain in the same position relative to each other. As a result of zone settling, a relatively clear water zone (CWZ), or free-settling zone (type II), is formed above the sludge blanket.
3.1.4 Type IV Compression settling occurs when there is a structure formed between the particles and the bottom of the tank. Further consolidation occurs by compression as a function of time. A transition zone also occurs between type III and type IV settling. This type of settling occurs in the bottom area of primary clarifiers and thickeners. It may occur in secondary clarifiers, to a limited degree, depending on the process design and operation.
3.2 Special Considerations for Nutrient Removal Sludges Nutrient removal can refer to nitrogen and/or phosphorus removal. Sludges from nitrifying systems can contain significant denitrifying microorganisms and can float in clarifiers if nitrate comes into contact with them. Typically, this is minimized by
419
420
Biofilm Reactors
removing enough nitrate in the reactors, maintaining high dissolved oxygen going into reactors, and/or removing sludges quickly. For systems incorporating biological phosphorus removal, phosphorus can be released if sludges are allowed to go anaerobic. This typically is avoided by maintaining high dissolved oxygen concentrations going into reactors and/or removing sludges quickly or by using DAFs.
3.3 Clarifier Enhancements The overall performance of any biological process (e.g., trickling filters, biotowers, RBCs, and activated sludge) is inherently tied to the clarity of the effluent from the solids–liquid separation step. Trickling filter, biotower, and RBC effluent TSS concentrations generally are rather low, at a range of 30 to 80 mg/L for domestic wastewater. Thus, the settling regime is type II with effluent flocculation. This occurs because effluent solids settle as discrete particles or separate particle aggregates because the concentration of TSS in the effluent is too low to produce a well-flocculating mass (which would improve clarification). Understanding this situation, investigators (Norris et al., 1982) developed the follow-on solids contact process, which is a short-term aerated flocculation step following trickling filters. It is equally adaptable to RBCs, activated biofilters (BAFs), and other fixed-growth reactor (FGR) effluents. The activated sludge portion of the dual biological processes ranges from shortterm aeration (TF/SC) designs, where soluble substrate conversion essentially is completed in the trickling filter, to dual systems, in which the TF/AS systems share the soluble substrate conversion. For these systems, the clarification and return sludge requirements are similar to those of a single-stage activated sludge system. The clarification environments of common trickling filter systems may be categorized as shown in Table 9.1. The most difficult clarification modes are roughing trickling filters (highly loaded), where the effluent TSS represent mostly un-metabolized primary effluent solids and young biological cells. The clarified effluent typically will be turbid because of the reduced level of bioflocculation. The common practice of using comparatively high OFRs for intermediate clarifiers following heavily loaded FGRs partially defeats the goal of removing the slow-to-separate trickling filter effluent TSS. The efficiency of the intermediate clarifier has a direct influence on the sizing requirements and performance of the subsequent trickling filter or activated sludge stages. Good clarification of roughing filter effluent requires the same design considerations that are necessary for high removals of TSS by primary clarifiers.
Clarification
Table 9.1 Clarification characteristics of trickling filter systems in municipal WWTPs.
Mode
Treatment
BOD5 TSS/ load MLSSa 3 (kg/m ·d) (mg/L)
Trickling filter
Roughing
>1.0
40–120
Trickling filter
bod5r
0.2–1.0
40–120
–
Type ii
Trickling filter
BOD5R/NODR
0.1–0.3
40–120
–
Type ii
Trickling filter
Tertiary NODR
0.1–0.2
10–30
–
Types I–II
ABF
bod5r
0.2–0.8
500–1000
5–10
Types II–III
ABF/AS
bod5r
0.8–0.3
1000–3500
40–70
Types II–III
Trickling filter/SC
Flocculation
0.2–1.0
500–1500
10–30
Types II–III
Trickling filter/AS
bod5r
1.0–3.0
1000–3500
40–100
Types II–III
* *
RAS (%Q)
Settling regime Types I–II
MLSS = mixed liquor suspended solids. NOD = nitrogenous oxygen demand.
a *
These same considerations extend to trickling filters designed for complete treatment, BOD5R, and combined BOD5R and nitrogenous oxygen demand, where the TSS concentrations are low; while the solids may be flocculent, the low TSS levels do not produce the enhanced settling associated with type III settling. Therefore, these clarifier designs must provide for good entrance conditions with low turbulence, adequate retention time, low OFRs, and launder-weir arrangements, to control velocity currents. The Activated Bio Filter (ABF; a proprietary system that combines fixed-film and suspended-growth systems) mode of operation at BOD5 loadings of 0.2 to 0.8 kg/m3·d is designed primarily to enhance settling, by increasing the concentration of flocculent solids in suspension—that is, to change the settling regime from type II to type III. There is some additional oxidation that occurs in the short residence time of the recycled TSS in the filter, which may improve the flocculent nature of the solids. At higher organic loads in the ABF/AS system, a significant portion of the oxidation is in the suspended-growth reactor (SGR) portion of the dual system. The TF/SC process is designed to enhance the separation characteristics of the trickling filter biomass by promotion of the already flocculent nature of the solids through an increase in TSS concentration and reactivation by aeration. The aeration
421
422
Biofilm Reactors
regenerates the flocculent nature, which was reduced in the oxygen-limited clarifier sludge blanket. The aeration solids retention time (SRT) of these designs is short, between 0.2 to 1.0 days, because bioconversion (removal of soluble BOD5) essentially is completed in the trickling filter. The settling characteristics of the solids in the clarifier influent are type II and type III, and excellent solids removal can be expected. Settling rates are high because of the denser biomass produced by the trickling filter. The TF/AS systems produce MLSS settling characteristics similar to an activated sludge facility of the same MLSS and diluted SVI (DSVI). If the trickling filter provides ≥70% soluble BOD5R, the resulting solids typically possess a low DSVI. However, the activated sludge system still should be a plug-flow design with initial selector zone(s). This will prevent filamentous growth during overloading of the filters and soluble BOD5 breakthrough. The TF/AS settling regime is also type II and type III, and advanced secondary treatment (<10/20 mg/L carbonaceous BOD5/TSS) can be achieved readily without effluent filtration.
3.4 Wastewater Flocculation The mechanical flocculation of raw wastewater to improve the removal of TSS and BOD5 has been practiced for a long time. Darby (1939) and Fischer and Hillman (1940) discussed the performance of centrally driven mechanical flocculators in circular clarifiers. It was found that these units, without chemical addition, performed better than conventional clarifiers. The enhanced performance is recognized today to be the result of optimizing the flocculation potential of the incoming particulates. The designer should note that many of these devices also provide greater control of the distribution of flow to a clarifier, an energy-reduction step that is known to improve clarifier hydraulics. As a consequence of both of these effects, good clarifier design will incorporate this important feature. Flocculation of biological effluents, trickling filters, and activated sludge began in the 1950s, if not earlier. The Walker Process Equipment Company (Aurora, Illinois) patented the Carillon inlet with a tangential discharge into the feed well in 1953, which was installed in several activated sludge facilities by 1970. This feed-well design used hydraulic energy in a tangential flow mode to circulate the feed-well contents, encourage solids contact, and promote floc growth. In the late 1960s, the use of multiple mechanical mixers in clarifier feed wells also evolved, first for water treatment using alum and then for primary and secondary wastewater streams. During the period 1965 to 1980, there were many mechanical and hydraulic flocculating feed
Clarification
wells installed in activated sludge clarifiers. However, the use of these devices did not constitute common or standard practice. Parker et al. (1970, 1971, 1972) studied the aggregation and breakup of activated sludge floc in continuous-flow, bench-scale flocculators. In studying activated sludge with three different SRTs, they found that the optimum energy gradient (G, second−1) values for mechanical flocculation were from 20 to 70 second−1, and optimum retention times were 20 to 60 minutes. This is opposed to the G values found in 14 operating plants, which ranged from 88 to 220 second−1. In pilot and field studies at six locations, they determined that floc breakup exceeded aggregation in aeration basins and/or aerated distribution channels and that flocculation would improve the effluent qualities of the plants. This work led to renewed interest in flocculator–clarifier designs and the enlargement of feed-well diameters with mechanical mixing to accomplish flocculation. While Argaman and Kaufman (1970) and Parker et al. (1970, 1971, 1972) laid the foundation for understanding the mode and role of flocculation of activated sludge particles from bench-scale studies, there was little supporting data collected from many full-scale facilities installed in the late 1960s and 1970s. While there were perhaps more than 50 wastewater plants with flocculating clarifiers constructed during that period, flocculating units were not considered standard practice until the late 1980s. Norris et al. (1982), Parker (1983), and Stukenberg et al. (1983) recommended the use of flocculating feed wells to encourage floc growth to offset breakup in the aeration basin. Parker (1983) and Stukenberg et al. (1983) reported that multiple mixers in the feed wells were not effective in producing conditions optimal to floc aggregation in full-scale operations. The mechanical mixers have high tip speeds (0.6 to 1.8 m/s) and increased floc-shearing effects compared with centrally driven flocculators at 0.15 to 0.6 m/s. At comparable tip speeds, smaller turbines are unable to impart sufficient mixing energy, and effluent TSS concentrations were equal or lower with the mixers off in the larger feed wells. During the late 1980s and 1990s, the use of influent hydraulic energy to induce flocculation of wastewater solids in the feed-well zone became widely accepted and was used for all types of clarification applications. In activated sludge, TF/AS, and TF/SC processes, these floc-promoting final clarifiers can produce effluents of 4 to 10 mg/L TSS. It also is logical to assume that these same devices in primary, trickling filter, and RBC clarifiers will enhance performance significantly.
423
424
Biofilm Reactors
3.5 Flocculation Criteria The performance of the clarification device is enhanced by reducing, by flocculation, the number of smaller particles. The result is a higher unit capacity and improved overflow quality because of the formation of faster settling, larger agglomerates from the finer particles. Floc shear and breakup in the aeration system is not necessarily detrimental, because this increases the surface area and decreases diffusion resistance (LaMotta, 1976). Because the aeration system will cause floc breakup, conditions that favor deflocculation/reflocculation need to be defined. Fortunately, laboratory and recent full-scale studies provide the information required. There are a number of aspects to be considered in the flocculation of FGR and SGR effluent before the secondary clarifier. While there may be others, four of the primary factors are the following:
(1) Effects of prior treatment on the biological floc, (2) Level of energy, (3) Method of flocculation, and (4) Time required to achieve flocculation.
To define the design requirements for good flocculation, first it is necessary to establish whether prior treatment of the floc will be a factor in the design requirements for flocculation. Prior treatment is defined by the type of aeration equipment and the intensity of aeration (e.g., kW/m3 [hp/1000 cu ft] or L/min air/m3 [scfm/1000 cu ft]) and the mode of mixed-liquor transport from the aerator to the clarifier. Parker (1983) and Parker et al. (1971) reported that the energy intensity and subsequent floc-shearing level can affect conventional clarification of MLSS. Specifically, aeration systems with a high-power input/unit volume (kW/m 3) and aerators with high shear (high tip speed, motor-speed surface aerators, etc.) may adversely affect clarification efficiency. Excessive shearing of floc agglomerates will result in higher effluent TSS concentrations, even at low OFRs, unless there is provision for reflocculating. The question of whether the activated sludge will reflocculate regardless of prior treatment was researched by Wahlberg et al. (1992). A field study of 21 WWTPs evaluating 30 different activated sludge systems revealed that flocculating and subsequent settling produced comparable effluent TSS concentrations, regardless of the initial aggregative state of the sludge or the aeration device used. Of these plants, nine used mechanical aerators, seven used coarse-bubble facilities, and five used fine-bubble aeration facilities.
Clarification
100
12
80
9
Turbidity (NTU)
Suspended Sollds Conc. (mg/L)
While Wahlberg et al. (1992) found that higher levels of turbulence could extend the flocculating time, this condition was not considered significant. The flocculation was 99% complete within 10 minutes; typical flocculation results from these studies are shown in Figure 9.2. The designer should note that these are batch tests, which would simulate a perfect plug-flow flocculator. A significantly longer (2 to 3 times greater) flocculation time would be required in a completely mixed flocculator. Nevertheless, the results demonstrate that reflocculation can be accomplished relatively quickly, regardless of the prior turbulence level.
60 40
6 3
20 0
0 0
10
20
30 40 50 60 Turbidity (NTU)
70
0
80
10
16
60
12 8 4 0
10
20
30
30
40
50
60
70
60
70
(b) 75 Turbidity (NTU)
Turbidity (NTU)
(a) 20
0
20
Flocculation Time (min)
40
50
Flocculation Time (min) (c)
60
70
45 30 15 0 0
10
20 30 40 50 Flocculation Time (min) (d)
Figure 9.2 (a) Suspended solids concentration and turbidity data collected throughout the study with a linear relationship function and 95% suspended solids concentration prediction interval estimate; (b) turbidity/time data collected from the Camp Creek plant on August 13, 1990, with batch flocculation curve fit; (c) turbidity/time data collected at the Utoy Creek plant on July 19, 1990, with batch flocculation curve fit; and (d) turbidity/time data collected at the Coneross plant on May 30, 1991, with batch flocculation curve fit.
425
426
Biofilm Reactors
Parker (1983) presented data on the effects of the aeration mode, intensity of the floc size, and effect on effluent suspended solids. That study illustrates the effect of higher aeration intensity on final effluent quality. These data were generated without the benefit of a flocculation stage before settling. In the same paper, the aeration G factor was presented as a function of the chemical oxygen demand removal rate (kg/mgd), and, as expected, the high-rate aeration systems produced a higher G value, as shown in Table 9.2. Flocculation G factors are effective in the range 15 to 25/s, although breakup may not occur up to 50 to 70/s (Figure 9.3). The effect of the method of agitation on flocculation–clarification was studied by an equipment company in the mid-1970s. Eimco PMD (1974) evaluated mixing, aeration, and tangential stirring of the mixed liquor before settling. Aeration was found to be the poorest means of agitation, while stirring with the low shear of tangential mixing consistently proved to be the best method of producing a high-clarity effluent. Typical results of the several tests (Eimco PMD, 1974) conducted are shown in Figures 9.4 and 9.5. The initial observations by Eimco (Salt Lake City , Utah) personnel that led to these flocculation studies occurred at a Dutch Carrousel oxidation ditch at Wolfgang WWTP in 1974. Settling vessels set up at intervals around the 20- to 25-minute circuit
40
Study I
Study II
50
100
30
20
10 0
0
0.1
0.25
150 G, sec –1 0.5
200 1.0
250 1.5 2.0
300 2.5
Air Rate, cfh/gal
Figure 9.3 Effect of coarse-bubble aeration rate on effluent solids (cfh/gal × 0.007 482 = n3/L·h).
Table 9.2 Energy gradient (G) values as a function of aeration type. Oxygen transfer SOTR at standard (kg O2/MJ) conditions (%) [lb/hp-hr]*
Aerator type
G (second–1) COD removal rate 0.34 d–1
0.86 d–1
1.75 d–1
Coarse bubble (low efficiency)
8
0.38 (2.26)
115
183
261
11
0.46 (2.70)
96
153
218
–
0.34 (2.0)
122
194
277
–
0.59 (3.5)
92
147
209
20
0.95 (5.6)
73
116
166
30
1.33 (7.9)
62
98
140
Coarse bubble (high efficiency) Mechanical aerator (low efficiency) Mechanical aerator (high efficiency) Fine bubble (low efficiency) Fine bubble (high efficiency) * Wire-to-water basis. Turbidity - JTU 35
Tangential Stirring Aeration - Fine Bubble Mechanical Mixing None
30 25 20 15 10
TSS = 2.0 Turbidity - 14.0 (0.25 mg/L TSS)
5 0 0
4
8
12
16
20
Settling Time - Minutes
Figure 9.4 Flocculation procedures, test number 1. 427
24
28
32
428
Biofilm Reactors
Turbidity - JTU 60
TSS = 3.93 Turbidity - 33 (0-25 mg/L TSS)
50
Tangential Stirring Aeration - Fine Bubble Mechanical Mixing None
40 30 20 10 0
0
4
8
12
16
20
24
28
32
Settling Time - Minutes
Figure 9.5 Flocculation procedures, test number 2a. showed a significant enhancement in clarity after 5 to 15 minutes of travel in the looped channels flowing at 0.27 to 0.36 m/s. The surface aerator impeller tip speed was approximately 4.3 m/s. The required reflocculating time after the higher shear aeration zone was in the range 3 to 10 minutes, which was consistent with the results shown in Figures 9.4 and 9.5. Das et al. (1993) evaluated and quantified aeration effects on floc particles in 24 WWTPs. This study confirmed earlier results, which showed that energy levels in excess of G values of 70 to 80/s would cause floc breakup. It also was shown that the floc would re-form rapidly, if the energy level was reduced before clarification. Das et al. (1993) quantified reflocculating and improved clarification as a function of distance from the aerator, similar to the effect previously noted by the Carrousel oxidation ditch supplier (Eimco, Salt Lake City, Utah). Tapered aeration, common to the biological nutrient removal designs, was beneficial to clarification because of lower G values (less shear) in the effluent end. The mode of MLSS transport from aerator to clarifier was found to be a factor in clarifier performance (Das et al., 1993). In some cases, there was higher percentage of dispersed TSS in the outlet, which was attributed to the flow-control structure. The transfer piping velocity and direction changes of the flow typically did not significantly influence the level of dispersed TSS in the ranges of velocity common to biological treatment.
Clarification
Heavy aeration of the transfer channels to the clarifiers will break up floc, which may not agglomerate in standard or conventional feed-well designs. While flocculating feed wells may repair the disrupted floc, aeration in transfer channels should be turned down to the minimal rate necessary to prevent sedimentation. Fine-bubble diffusers cause less shear than coarse-bubble diffusers. Also, there should be minimal free-fall zones between the aerators and clarifiers. Das et al. (1993) and Wahlberg et al. (1992) summarized questions raised by earlier researchers about the use of the G value to define flocculation energy. While admitting to its shortcomings, they concluded that it is a relatively simple calculation that does have some significance. This information and results of field experiences may provide guidance for the designer in establishing effective flocculating systems. While the aeration basin may encourage floc breakup, the treatment of the MLSS as it leaves the aeration basin and is transported to the clarifier is effective for reflocculating. Measures that should be taken include the following: • Avoid waterfalls from the aeration basins of more than 200 mm, • Eliminate control structures and piping with high (>0.6 m/s) velocities, • Eliminate or minimize channel aeration using fine-bubble devices, and • Provide 5 minutes or more hydraulic flocculation before the clarification zone. This flocculation requirement involves a separate compartment for a rectangular clarifier with a low headloss diffusion inlet in the clarifier. The flocculation zone in circular clarifiers is the vertical projection of the flocculating feed well to the dense sludge blanket to extend the flowing zone (density current) emanating from the feed well.
3.6 Clarifier Design Details This discussion in this section will be oriented primarily to circular clarifiers, because they are the dominant form used for trickling filter flow sheets. Recommendations for the design of rectangular clarifiers and additional secondary clarifier design discussions are provided elsewhere (IWA, 1992; Water Environment Federation, 2005; WEF et al., 2009). The design of a clarifier should be approached with a methodology that accounts for the necessary compromises because of the interactions of the components of the clarifier. For example, it is an obvious advantage to introduce the feed from the feed well to the clarification zone with the least amount of turbulence; thus, the larger feed
429
430
Biofilm Reactors
well was developed. However, it is also obvious that, at some point, increasing the feed-well size will intrude on the clarification volume and reduce its ability to equally distribute the feed, because the contained energy level will be too low. The interrelation of those variables, shown in Table 9.3, is not well-understood, and many conclusions have been drawn based on site-specific conditions and the inherent limitations of existing equipment and design. For example, a clarifier with a poor feed-well design will produce better results if the clarifier is larger or deeper than required. This may compensate for the inadequacies of the feed well. The effects of the several clarification factors in Table 9.3 show that many compromises must be made. However, these compromises should evolve from some common understanding and methodology to minimize the adverse effects of any compromise. Further, a common methodology will be a basis for developing a better understanding of clarification and result in further improvements. If the last 5 to 30 mg/L TSS in secondary treatment behave as type II particles, then the clarifier design that produces the best results with raw wastewater suspended solids also will produce the best secondary effluent. When there is an increase in influent solids because of solids recycle, the appropriate volume for sludge inventory should be added to the clarifier depth, and adequate sludge transport capacity should be provided to ensure timely removal of the settled solids. In designing clarifiers, design engineers also should consider the volume of the recycle flow, its energy input, and its disposition (i.e., overflow or underflow). Underflow recycle flows can generate their own directed energy currents as velocities increase at the point of removal. These currents also can be deleterious to clarification; that is, high recycle rates to the sludge removal mechanism can induce localized velocity currents along the mechanism. Generation of localized currents will destabilize the clarifier, regardless of their location, inlet, effluent, or sludge removal.
Table 9.3 Clarifier design features/interactions. Surface area
Side water depth
Slope of floor
Center column and inlet configuration
Weir configuration
Diameter and depth of the feed well
Scraper velocity
Stamford baffle/cantilevered launder
Sludge hopper placement
Sludge collector configuration
Clarification
The previous conclusion was that the same clarification principles can be applied in the design of all clarifiers, whether type I, type II, or type III conditions exist. Thus, the following design discussion is suitable for the secondary clarification considerations set forth in Table 9.1 and for primary and intermediate clarifiers.
3.6.1 Influent Column The sizing of the influent column for a secondary clarifier is based on the total flow to the clarifier. The maximum velocity in the influent pipe typically is 0.6 to 0.75 m/s (2 to 2.5 ft/sec). The outlet ports of the clarifier should be located as high as possible in the column to contain turbulence in the upper portions of the feed well. The outlets should not extend more than 0.75 m below the liquid level or closer than within 0.9 m of the bottom of the feed well when the flow is discharged directly to the feed well. For best effluent quality, an energy-dissipation device should be used to control and direct the energy in the outlet flow from the center column. The maximum velocity, and hence port area, is dependent on the inlet-feed-well design. Where the influent is discharged directly to the feed well, the maximum velocity should not exceed 0.60 m/s (2.0 ft/sec). The outlet port area is developed from this velocity limitation. Typically, the center column outlet area should be 125 to 150% of the center column area. Moreover, it typically is desirable for the influent to be discharged to an energy dispersion inlet, which reorients the flow to a horizontal tangential discharge to the flocculating feed well. In that case, the inlet velocity can be increased to 0.75 m/s (2.5 ft/sec), and the bottom of the center column outlet ports typically would be at ≤1.0 m of the surface.
3.6.2 Energy-Dissipating Inlet The energy-dissipating inlet (EDI), which was first patented as the Carillon Inlet (Walker Equipment Company, 1953), provides uniform and lower energy input in the influent flow to the feed well (see Figure 9.6). It also provides for the introduction of the influent to the upper 0.4 to 0.7 m of the feed well, thus reducing the required depth of the feed well. Where organ pipe rapid sludge removal units are used, the outlet ports are lower, because the sludge collection box and influent flow must be collected by the EDI under the sludge box and then discharged nearer to the water surface. The port area of the improved EDI unit is based on a maximum velocity of 0.61 m/s (2.0 ft/sec) and an average velocity of 0.30 to 0.36 m/s (1.0 to 1.2 ft/sec), whichever produces the maximum port area. Some designs have used moveable gates to
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Biofilm Reactors
Distribution trough Influent well ClariFlow port
Figure 9.6 EDI and hydraulic flocculating feed well.
modify the velocity. The long-term operability and necessity of this refinement should be considered carefully, because experiences of proper use are minimal. The depth of the outlet ports, as shown in Figure 9.7, should be limited to ≤0.8 m, to minimize the depth of the feed well, as discussed in the Feed Well (Flocculating Type) section. The volume of the EDI inlet structure is small, and the device typically will be 8 to 10% of the basin’s diameter and deep enough to collect the flow from the center
Clarification
Figure 9.7 General arrangement of the improved clarifier inlet structure.
column outlets. The minimum retention time will be approximately 8 to 10 seconds, based on the maximum total influent flow, but not significant if higher. The number of outlet ports may vary from 4 to 12 in the largest clarifiers. Installing an outlet port at every 1.2 to 1.5 m of the perimeter of the EDI should be considered by the designer. It also is important to provide a fully closed bottom (including port openings) and to extend the flow guidance vane beyond the opening (approximately 0.2 to 0.3 m), as shown in Figure 9.7, to fully direct the incoming flow in a horizontal and tangential pattern.
3.6.3 Feed Well (Flocculating Type) External flocculation has been used for rectangular basins, while circular basins typically use an enlarged feed well. Mechanical flocculator designs that have proven efficient have used a low tip speed (0.15 to 0.6 m/s) with centrally driven paddles. Flocculators designed with two to four individual flocculators in the feed well operate with tip speeds of 0.6 to 1.8 m/s and will break up floc particles; consequently,
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this design should not be used. At a lower tip speed of 0.2 to 0.6 m/s, these multiple units cannot provide sufficient mixing for flocculation. Current practice favors the use of tangential-directed, energy-dissipating inlets, where hydraulic energy is used to promote floc growth in the feed-well zone. The low shear developed by the tangential flow has been found to promote floc growth, improve settling, and result in lower effluent TSS in a cost-effective design. The feed-well design should provide for a large and uniform reduction of the inlet velocity and discharge of the influent to the clarifier at the lowest possible energy that still maintains good flow distribution. The objective is best attained with the inclusion of the aforementioned EDI device. It also is important that the feed well further reduce the average velocity as the flow passes down and out of the feed-well cylinder. Simultaneously, the inlet and feed well must produce a gentle hydraulic motion, which encourages flocculation of biological solids. The tangential inlets provide a low-shear, low-energy circulating motion, which results in improved clarification. Together, the tangential discharging EDI and feed well are considered to be a hydraulic flocculating feed well. Currently, there is no consensus of opinion as to the optimum maximum downward velocity in the feed well, although larger-diameter feed wells typically reduce the average velocity to less than 1.0 m/s. It is suggested that the downward velocity be limited to ≤0.75 m/min (2.5 ft/min) at the maximum total influent flow to the feed well and that the average velocity be ≤0.5 m/s (1.6 ft/min). Some references (Water Environment Federation, 2005; WEF et al., 2009) may acknowledge the use of higher feed-well velocities, but these rates are not suggested when the best clarification results are to be achieved. The depth of the feed well is a function of the configuration and operating conditions of the clarifier. The design should be based on a zone for flow passage between the dense-phase sludge layer and the bottom of the feed well, which does not result in an increase in the average velocity of the outward flow. That is, the vertically projected cylindrical area under the feed well and above the dense-phase blanket should equal or exceed the plan area of the feed well. As a result, the average outward flow velocity will be ≤0.50 m/s. Figure 9.8 illustrates the approach used to develop the feed-well depth. The minimum depth of the feed well is ≥0.9 m (3.0 ft) deeper than the outlets of the tangential outlets. The feed-well depth should be ≤1.22 m (4.0 ft) for smaller (<15 m) clarifiers and ≤1.8 m (6.0 ft) for units up to 60 m in diameter. A decision on the depth
Clarification
Feedwell f EDI f
Feedwell depth Side water depth
DH
Vv VH
Clear water & free settling zone (type I - II)
Effluent
dl v
Peak flow SBD
ds
Average SBD
Return sludge Influent
Figure 9.8 Dense sludge blanket profile: average and peak flows.
should be based on SWD, CWZ, slope of the floor, limiting scour of sludge, and type of process feed rate to the clarifier (e.g., influent wastewater, trickling filter effluent, and FGR/SC or FGR/AS effluent). Deep feed wells should be avoided because of the high potential for excessive sludge blanket scour, especially in shallow and large clarifiers. For TF/SC and TF/AS clarifiers, it is necessary to define the CWZ that occurs at peak flow conditions considering the transfer of a portion of the MLSS inventory from the aerator to the clarifier. This increased depth of sludge is denoted as div in Figure 9.8 and should not exceed 0.6 to 0.9 m. For primary, intermediate, and final FGR clarifiers, div. is essentially zero, as the sludge blanket is below the SWD. The TF/SC and TF/AS clarifiers may have reduced capacity because of the reduction of clarification volume during peak flows. Thus, general practice is to provide a deeper or larger clarifier to allow for the inventory shift in FGR/SGR processes.
3.6.4 Side Water Depth, Clear Water Zone, and Overflow Rate There is general consensus within the profession that SWD, CWZ, and OFR significantly affect clarifier performance. There also is a general consensus that these factors are interdependent. For example, greater CWZ depth allows higher OFRs, and vice versa. Unfortunately, little consensus exists in terms of quantitative relationships and tradeoffs. This section presents relevant available information. However, the reader
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must recognize that it is preliminary in nature, and greater experimental verification is required. It is hoped that the publication of this information will spur greater research and testing, so that an improved database will be available in the future. Currently, it is recognized that clarifiers should be constructed deeper when the diameters are larger, because hydraulic retention efficiency decreases as a function of diameter/SWD (IWA, 1992). Data collected on full-scale pulp and paper clarifiers by Erdman (circa 1958) showed a relationship between the clarifier center water depth (CWD)—hence SWD—and the hydraulic retention efficiency; the results are presented in Figure 9.9. During the 1960s, the recommendation to use deeper clarifiers similar to industrial practices of the 1950s (Erdman, circa 1958) was initiated and supported by later publications (Boyle, 1975; WEF et al., 2009). The need to increase SWD as a function of the diameter to maintain the retention efficiency also was presented (WEF et al., 2009). The SWD for circular clarifiers, as a function of the basin diameter, may be limited by site conditions. While it typically is most economical to use deeper, higher loaded clarifiers, this is not always possible, and the OFR must be modified as a function of the SWD (CWZ). Moreover, as experience has shown, good performance can be achieved with shallow basins, if the proper hydraulic rate is applied.
30
100 ft f
60 ft f
25 20 CWD (ft)
40 ft f 15 20 ft f
10 5 SWD = CED -
f 24
0 35
40
45
50
55
60
65
70
75
Detention – %
Figure 9.9 Clarifier detention: efficiency versus CWD (ft × 0.348 = m).
Clarification
For example, the following relationship between OFR, SWD, and CWZ have been reported (IWA, 1992): For primary clarifiers, primary clarifiers–thickeners (PC-Ts), intermediate clarifiers, and trickling filter clarifiers with floor slope ≥1:12,
Maximum OFR (m/h) ≤ 0.182 SWD2 (SWD = 1.83 to 3.05 m)
(9.5)
Average OFR (m/h) ≤ 0.092 SWD2 (SWD = 1.83 to 3.05 m)
(9.6)
Maximum OFR (m/h) ≤ 0.556 SWD (SWD = 3.05 to 4.57 m)
(9.7)
Average OFR (m/h) ≤ 0.278 SWD (SWD = 3.05 to 4.57 m)
(9.8)
For TF/SC and TF/AS with floor slope ≤1:12,
Maximum OFR (m/h) ≤ 0.182 CWZ2 (CWZ = 1.82 to 3.05 m)
(9.9)
Average OFR (m/h) ≤ 0.091 CWZ2 (CWZ = 1.82 to 3.05 m)
(9.10)
Maximum OFR (m/h) ≤ 0.556 CWZ (CWZ = 3.05 to 4.57 m)
(9.11)
Average OFR (m/h) ≤ 0.278 CWZ (CWZ = 3.05 to 4.57 m)
(9.12)
The CWZ is differentiated from the SWD, as shown in Figure 9.8. The value of div will increase at higher MLSS, sludge loading rates, OFRs, and peak-to-average flows because of the increased MLSS inventory shift. Average and maximum OFRs will generate two minimum clarifier areas. A designer should consider the practice of using the larger of the two areas generated from the above formulations based on the average of the maximum day/month and maximum month/year flows. While further testing of relationships such as these is needed, they illustrate the nature of the tradeoffs between these variables. While not mathematically presented, WEF and ASCE (2009) indicates a similar relationship of OFRs to SWD. When flat floors are used (WEF et al., 2009), the allowable OFR should be reduced by 0.34 m/h (200 gpd/sq ft). This compensates for the reduced clearance (DH) between the feed well and the sludge blanket. However, an alternate approach is to increase the SWD by 1.0 to 1.5 m (Parker, 1983, 1991; Parker and Stenquist, 1986) to maintain higher OFRs. The clearance, DH, under the feed well then will be similar for both sloped and flat-floor designs. The OFRs, as a function of SWD and CWZ, suggested above are based on the inlet, feed well, floor slopes, allowance for peak-load inventory shift, and weir
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arrangements set forth in this section. Inadequately designed configurations may cause short-circuiting and reduced performance and capacity. Further discussion of the recommended OFRs as a function of SWD and other design features for primary and secondary clarifiers can be found in earlier cited reports on deep clarifier performance and published design manuals (IWA, 1992; U.S. EPA, 1987; Water Environment Federation, 2005; WEF et al., 2009). While general practice is to design deeper clarifiers to minimize the surface areas for clarification, shallow units can provide equivalent performance when rated properly. Existing units often require low-cost upgrading of inlets, feed wells, effluent structures, and sludge-scraper mechanisms to provide their optimum performance and maximum capacity (Albertson and Alfonso, 1995; Albertson and Hendricks, 1992; Water Environment Federation, 1992).
3.6.5 Floor Slope The floor slope and resulting CWD must be consistent with the operating plan for the clarifier. Existing units may not be suitable for the most practical mode of plant operation, and this should be recognized before expanding the facility. For example, shallow or deep primary clarifiers with 1:12 ratio floor slopes were considered not suitable for co-thickening of primary and secondary sludges because of excessive sludge retention required to develop an adequate depth over the hopper. The floor slope will be a significant factor in the ability to provide good clarification and the ability to process the settled solids in an orderly manner. The design manual, Dewatering Municipal Wastewater Sludges (U.S. EPA, 1987), provides recommendations for the floor slope of primary clarifiers, as set forth in Figure 9.10. Primary and secondary sludges can be co-thickened efficiently in primary clarifiers that are configured properly. This is an important consideration in smaller plants, where separate thickening is not a desirable feature because of capital and attendant operation and maintenance costs. However, ignoring floor slope, basin configuration, and sludge transport requirements for thickening may result in the failure of center draw-off units to be efficient in sludge thickening. The floor slope also should vary as a function of the design requirements and size of the clarifier. The conical slope should provide for a minimum depth over the withdrawal sump to prevent “rat-holing” or short-circuiting of influent to the sludge hopper. As an example, minimum slopes for clarifiers with scraper collection systems are set forth in Table 9.4.
d
Type A
Recommended operation
dcwz
dswd
Separate thickening Slope = 1:12
vs ds
Type B
Combined clarification and thickening dcwd
d ≤ 15 m Slope = 2.5:12
Combined clarification and thickening Type C d ≥ 15–40 m Slope = 1:12 and 2. dt = 0.5d
dt
Combined clarification and thickening Type D
d ≥ 15–40 m
dh
Slope = 1:12 and 2. dt = 0.4d dh = 0.6–1.2 m
dt Legend
d = Diameter dcwz = Depth of clear water zone ds = Depth of sludge dswd = Side water depth
dcwd = Center water depth dt = Diameter at thicknener zone dh = Height of vertical wall
Figure 9.10 Recommendations for primary clarifier floor configuration.
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Table 9.4 Example minimum floor slopes versus diameter for clarifiers with s crapers. Diameter (m)
PC, IC, TFC*
Floor slope, (m/m) PC–Tb
TF/SC, TF/AS
10
0.20:1.0
0.30:1.0
0.24:1.0
20
0.120:1.0
0.25:1.0
0.14:1.0
30
0.083:1.0
0.20:1.0
0.100:1.0
40
0.083:1.0
0.05/0.2:1.0
0.041/0.125:1.0
50
0.083:1.0
0.05/0.2:1.0
0.041/0.125:1.0
60
0.083:1.0
0.05/0.2:1.0
0.041/0.125:1.0
Thickening of sludges is not practiced; PC = primary clarifiers, IC = intermediate clarifiers, and TFC = trickling filter clarifiers. b Co-thickening of primary and secondary sludges. Thickening zone is 50% of diameter. a
Most recent experiences revealed that improving sludge collection/transport and removal practices can result in both efficient clarification and thickening in primary clarifiers with conventional floor slopes. Memphis, Tennessee, and Phoenix, Arizona, after modifications, are producing excellent results as PC-T units (Albertson and Walz, 1997).
3.6.6 Effluent Weir and Launder The understanding that wall-mounted effluent baffles (Figure 9.11) improve the capacity and performance of final clarifiers evolved in the early 1980s in two separate locations (with the Stamford baffle in Stamford, Connecticut, and the McKinney baffle in Lincoln, Nebraska) (Stukenberg et al., 1983; Water Environment Federation, 2005). The Lincoln design used an inward horizontal extension of the bottom of the wall-mounted launder, while the Stamford design used a sloping baffle mounted below the launders. The McKinney baffle installed at the Lincoln WWTP (Stukenberg et al., 1983) produced excellent results, with average OFRs up to 1.75 m/h (1 035 gpd/sq ft). The Stamford baffle test results (Semon, 1982) are shown in Figure 9.11 and demonstrate that the operating capacity was extended from 1.36 m/h (800 gpd/sq ft) to approximately 1.7 m/h (1000 gpd/sq ft), without significant loss of suspended solids. Using a similar design concept, there was no deterioration (TSS = 5 to 9 mg/L) in the effluent quality at Columbus Southerly WWTP (Ohio) (Albertson et al., 1992) at monthly average OFRs up to 1.55 m/h (915 gpd/sq ft), solids loadings of 5.7 kg/m2·h (28 lb/d/sq ft), and weir rates up to 562 m3/m·d (45 400 gpd/lin ft).
Clarification
60
TSS mg/L
50 40
Stamford Baffle 2 - 39.6 m f 3.96 m SWD Feedwell - 4.40 m f
99
84
With/Baffle W/O Baffle
30 20 10 SM – 90 TO 130 cc/gm MLSS – 2000 TO 3000 mg/L gal/sq ft.d 600 m/hr 1.02
O‘Flow rate 800 700 1.38 1.19
900 1.52
1000 1.59
Figure 9.11 Clarifier improvement with Stamford baffle. Industrial secondary clarifiers that are ≥90 m diameter (Ø) and typically deep use a single weir on a wall-mounted launder at high rates without problems of solids carry-over. The Santa Rosa, California, WWTP has four 35.1 m × 5.5 m (115 ft diam × 18 ft SWD) final clarifiers, which are equipped with draft tube rapid sludge removal (organ pipe) collectors. Three units were equipped with cantilevered inboard launders. The new unit is equipped with a wall-mounted launder and a McKinney effluent baffle. The bottom of the wall-mounted launder extends inward 1.47 cm to prevent upwardflowing wall currents. Comparative stress tests conducted at Santa Rosa (Buttz, 1992) demonstrated that the baffled inside-wall launder was equally effective as the cantilevered double weir-launder located at 26.8 m diam (88 ft) in the 35.1 m diam × 5.5 m (115 ft diam × 18 ft) SWD clarifiers. These studies were conducted to evaluate the relative efficiency of the two launder designs over a wide range of flow conditions, resulting in OFRs up to 3.0 m/h and solids loading rates (SLRs) up to 8.2 kg/m2·h (40 lb/d/sq ft). The results of the studies are summarized in Table 9.5. The extensive data were summarized incrementally using the ranges of the OFR to evaluate the effect of OFR and SLR on the effluent TSS for each configuration. A designer should note that there
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Table 9.5 Results of full-scale study of launder positioning. Wall mounted
Cantilevered inboard OFR (m/h)
No. of samples
TSS (ng/L)
SLR (kg/m2.h)
–
0.93
5
5
4.0
5
4.8
1.15
15
5
4.5
16
6
5.2
1.38
12
8
5.0
1.65
15
10
5.8
1.58
14
8
5.4
1.85
18
7
6.1
1.87
8
10
6.0
2.12
15
8
6.6
2.15
10
5
6.6
2.35
15
6
7.1
2.35
11
8
7.1
2.64
27
7
7.7
2.56
7
7
7.5
2.86
8
7
8.2
–
–
–
–
OFR (m/h)
No. of samples
TSS (ng/L)
–
–
–
1.13
9
1.37
SLR (kg/m2.h)
are only negligible differences in the results from the two effluent launder designs. The daily and average results did not show any measurable effect of OFR or SLR on effluent TSS. The 10-mg/L TSS reported for each of the launder types had 50th percentile values of 7 to 8 mg/L, suggesting that there was little interruption in performance over the flow and solids loading range tested. At low values of TSS, the distribution of data becomes one-sided, and averages are biased by single or infrequent events. In summary, it appears that the newer approaches, such as the cantilevered double launder design and new wall-mounted baffle designs, produce similar results. It also is clear that either of these designs also perform better than the conventional outside-the-wall launder configuration without baffles (Stamford or Lincoln). Figure 9.12 illustrates three designs that will affect the Lincoln/Stamford baffle. The recommended horizontal extension of the baffle is defined by the following:
Stamford baffle (cm) = 460 mm + 16.7 (φ – 9 m) Stamford baffle ≥ 460 ≤ 1200 mm Baffle = 18 in. + 0.2 in. /ft (φ – 30 ft) Baffle ≥ 18 ≤ 48 in.
(9.13)
Clarification
Scum Baffle
Launder cover
Launder cover
Scum Baffle
1 1 1
SB
SB Type B
Type A 1 Original McKinney (Lincoln) baffle extended horizontally. A slope of >45 degrees prevents sludge accumulation and the baffle may be constructed of concrete, steel, aluminum, or fiberglass.
Scum Baffle
SB Type C
Figure 9.12 McKinney (Lincoln)/Stamford baffle arrangements.
The extensions may be either a part of the effluent launder or a separate baffle mounted on the wall. The baffle should be placed at a sufficient distance below the water surface of secondary clarifiers to discourage the growth of algae on the baffle surface. It is recognized that existing state and regulatory standards may suggest maximum weir rates requiring double or even triple effluent weir arrangements. However, these requirements should be reevaluated in light of experiences with the McKinney/ Stamford baffle (Albertson et al., 1992; Buttz, 1992; Semon, 1982; Stukenberg et al., 1983) and use of deeper basins. The reader should note that there are more than 12 plants with Stamford baffles in Connecticut and that the state mandated these effluent baffles for all new treatment facilities (Water Environment Federation, 1992).
3.6.7 Sludge Collectors For primary, intermediate, and FGR clarifiers, the conventional collector in the United States has been the multiplied straight scraper design. For TF/SC and TF/AS,
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both scraper and suction-type collectors have been used. While the scraper straight-bladed and spiral mechanisms typically have been used overseas for clarifier applications, the suction (draft tube or header) unit has been the preferred mechanism for secondary clarifiers in the United States. Wahlberg et al. (1993) studied full-scale scraper and organ pipe rapid sludge removal performance and found similar effluent TSS concentrations and noted that the scraper mechanism was unaffected by changing settling and flocculation characteristics. More recently, there has been an increase in the number of scraper clarifiers used in the United States because of the desire to use higher MLSS for advanced WWTPs without significantly increasing the sludge return rate. The hypothesis that there are no process advantages to using the more costly and higher operations and maintenance suction or rapid sludge removal clarifier is being questioned. There are two types of rapid sludge removal clarifiers—the suction header (orifice tube) and the multiple (uptake) draft tube or (organ-pipe) design, which has individually adjustable sludge-withdrawal pipes. The header design requires a flat floor and a somewhat deeper (1.0 to 1.5 m) SWD. It has been suggested that the current organ-pipe design is improperly configured for best performance (Albertson, 1991). The organ-pipe design requires operator attention because of plugging of the withdrawal pipes. Both designs have underwater seals that require maintenance. The spiral scraper has advantages over the straight-blade collector (Albertson and Hendricks, 1992; Albertson and Okey, 1992; Billmeier, 1988; Günthert, 1984; Warden, 1981), in that the centerward transport rate of solids is much faster, because the solids remain in contact with the blade. The primary clarifier, intermediate clarifier, and TFC scraper designs are straightforward because of the relatively low solids loading. The PC-T, TF/SC, and TF/AS clarifiers must be evaluated on a sludge transport–sludge depth basis (Albertson and Hendricks, 1992; Albertson and Okey, 1992; Albertson et al., 1992; Billmeier, 1988; IWA, 1992). That is, it is necessary to calculate the sludge depth as a function of the operating conditions to configure the scraper blade for TF/SC and TF/AS applications. Albertson and Okey (1992), Billmeier (1988), International Association on Water Quality (1992), and Warden (1981) set forth procedures for determining the design of the scrapers for the TF/SC and TF/AS influent conditions using spiral scrapers in a central sludge draw-off design. The design procedure determines the quantity and depth of sludge to be transported to establish the scraper angle of attack, rake speed, and depth of blade. The results are tapered spiral scrapers of 200- to 300-mm depth at the wall increasing to 800 to 1000 mm in the center in 40- to 60-m diam clarifiers.
Clarification
In larger clarifiers (>30 to 35 m diam) receiving high MLSS, it often is necessary to provide four arms to assist sludge movement to the hopper. The two additional arms typically extend to 25 to 50% of the diameter. These secondary scrapers can significantly reduce the short-circuiting of influent to underflow of secondary clarifiers (and primary clarifiers) and often are required for primary clarifiers that also co-thicken sludges.
3.6.8 Sludge Hopper The size and location of the sludge hopper is important to the optimum performance (Albertson and Orris, 1994) of a circular secondary clarifier. The quantity and depth of sludge to be removed increases geometrically as the sludge approaches the hopper. If the hopper is located at 10% of the radius, then 99% of the sludge must be transported to 1% of the basin area. If the hopper is too small, then the high velocities will cause short-circuiting as the overlying MLSS breaks through the sludge blanket. A long, narrow sludge hopper design that extends from approximately 5 to 7% to 23 to 25% of the radius will provide the best results. The area of the hopper should limit the vertical velocity of the maximum return activated sludge (RAS) flow through the top plane of the hopper to 1.5 to 3.0 m/min. Smaller clarifiers with a lower overlying sludge blanket should be designed with a lower entrance velocity. The velocity should not exceed approximately 1.5 (CWS-SWD, m), to limit short-circuiting. The length-to-width ratio of the hopper would be 3 to 3.5:1. The hopper depth is less significant, and multiple outlets (two to four) to the RAS pipe will reduce the potential for vortexing and subsequent short-circuiting. A general arrangement of the sludge hopper is shown in Figure 9.13. The hopper would be centered at 14 to 17% of the basin radius.
3.7 Rectangular versus Circular Clarifiers Properly designed, both types of clarifiers will produce equivalent effluent TSS. However, there is a general preference for circular basins for secondary clarifiers and for smaller units. Space considerations still may dictate the use of rectangular units. There are references to length/width of the rectangular basin as being significant design criteria. In reality, length may be the most important factor to dissipate the influent energy. Typically, rectangular clarifiers should have a minimum length of 40 m (130 ft) and preferably 60 to 80 m (200 to 260 ft). Longer basins have consistently proven more efficient.
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22–25% R Center Column
6 – 8% R
RAS
Q + RAS Influent Pipe
Sludge Hopper Center Column Support
Figure 9.13 General arrangement of sludge hopper.
Like circular basins, rectangular basins will perform more efficiently when constructed with adequate depth. The increased depth reduces the forward-flow velocity; hence, lower turbulence occurs. Because longer basins handle more flow, depth also should increase as a function of length to reduce the forward-flow velocity. The effluent depth in larger basins (60 to 80 m) should be 4.0 to 4.8 m. The floor slope is a function of the application. The floor slope should provide at least 1.2 to 1.5 m of sludge depth at the sludge hopper without the sludge blanket extending to the effluent end floor–wall junction at average loading conditions. For peak flow conditions, inventory shift from the aeration basin during peak flows must be considered in dual biological designs. The sludge hopper may be located at the influent end, one-third or one-half length, or at the effluent end of the basin. For mixed liquors (like that found in IFAS plants), the half-length and effluent-end location is best. Concurrent flow and transported sludge units have demonstrated excellent performance (Wilson, 1991). Transport requirements (e.g., sludge depth and relative scraper capacity) will determine the best location of the sludge hopper(s).
3.8 Design Example An example of the integration of the various design factors is presented. Table 9.6 sets forth design requirements for this example, while Figure 9.14 illustrates the type of clarifier to be designed.
Table 9.6 Design basis for primary and secondary clarifiers. Primary Number of units
Secondary
3
3
Average influent flow (m /s)
1.2
1.2
Peak influent flow (m /s)
2.0
2.0
Maximum RAS flow (m /s)
–
0.720
MLSS (mg/L)
–
3500
SVI ≤ mL/g
–
120
SWD ≤ m*
4.0
4.6
3
3
a
3
RAS = return activated sluge. SVI = sludge volume index. * Limited by the site conditions for the clarifiers. Aeration volume is 35 000 m3. a
b
Sub Arm (>30m Diameter)
Direction of Rotation
Ports
One Blade (>30 m Diameter)
Two Blade (>30 m Diameter) Scum Skimmer Scum Baffle
EnergyDissipating Inlet Feed Well
Tangential Diffuser Ports
Effluent Launder and Baffle
Scum Box Stub Arm Rotated 90 deg
10 Influent Pipe
Return Sludge Pipe
Figure 9.14 General arrangement of improved wastewater clarifier. 447
448
Biofilm Reactors
3.9 Clarifier Following Moving-Bed Biofilm Reactor, Trickling Filter, Rotating Biological Contactor, and Biotower These are designed similarly to primary clarifiers and are designed to maximize TSS and associated BOD 5 removal and also can be used to thicken the waste sludge, although current practice typically is to provide separate thickeners (Water Environment Federation, 2005). The clarifier OFR based on a 4.0-m SWD is established as follows: oFr aVg = 0.278 (4.0 m SWd) = 1.11 m/h oFr MaX = 0.556 (4.0 m SWd) = 2.26 m/h The area required for the clarifier is as follows: A MaX =
(2.2 m 3 /s) (60 s/min) (60 min/h) = 3504 m 2 (2.26 m/h)
A aVg =
(1.2 m 3 /s) (3600 s/h) = 3892 m 2 (governs) (1.11 m/h)
The size of the clarifiers is as follows: CWd = SWd + 0.05 m/m(φ /4) + 0.02 m/m(φ /4) = 4.0 + 0.05(41 m/4) + 0.2(41 m/4) = 6.56 m Because only fixed-film reactor sludge is being collected, the design would be similar to the primary clarifier sizing, except that a simple sloped floor would be 1:12 m/m. The CWD of the clarifier would be as follows: 3892 m 2 φ= (3 units)(0.785)
0.5
= 40.7 m
∴ 3 to 41 m φ x 4.0 m SWd clarifiers The average and maximum flow/clarifiers = 0.40 (QAVG) and 0.773 m3/s (QMAX). Center column (CC): VMAX ≤ 0.75 m/s at (QMAX) = 0.733 m3/s Outlet ports (four) ≈ 125% CC area and ≤ 0.70 m deep (sets EDI minimum depth). Center column design: Diameter ≥ 1.12 m CC VMAX ≤ 0.75 m/s
Clarification
Ports = four 0.44 m width × 0.70 m water depth. Ports VMAX = 0.60 m/s EDI: Detention ≥ 8 to 10 seconds at QMAX Depth = CC outlet port + 0.1 m ≥ 0.8 m Vp ≤ 0.36 m/s average, ≤ 0.61 m/s maximum Number of ports = 8 Depth of ports = 0.60 m Width of opening =
(0.733 m 3 /s)(1000 mm/m) (8)(0.6 m)(0.61 m/s)
= 250 m EDI design: Size = 3.5 m diam × 0.8 m water depth Detention = 10.5 seconds at QMAX Ports = Eight at 250-mm width × 600-mm water depth A typical arrangement is illustrated in Figure 9.15. (0.733 m 3 /s) (60 s/min) Feed well ≥ (0.75 m/min) (0.785) ≥ 8.24 m (use 8.75 m)
0.5
The area of the flocculating feed well is based on a maximum downward velocity (VV) of 0.75 m/min. The diameter of the feed well is as follows: The minimum clearance (DH) between the bottom of the feed well and the maximum sludge blanket height is defined by the cylindrical area through which the flow discharges at a velocity ≤0.9 of the vertical velocity (VV) in the 8.75 m diam feed well. The outlet area is ≥111% of the feed-well area. (8.75 m ) 2 (0.785) (8.75 m) (3.14) (0.9) ≥ 2.43 m
DH ≥
The maximum DH depth available is based on the EDI port depth (0.7 m) and the clearance between the bottom of the EDI port and bottom of the feed well (minimum 0.6 m, target 0.9 m).
449
450
Biofilm Reactors
Fit to 30 mm of center column 3.5 m DIA EDI
Overlap
Outlet Port 250 mm
300 mm
300 mm
Typical port Total 8
150 mm 600 mm
800 mm
Clarifier: Feedwell: Avg/max infl. flow:
41 mØ x 4.0 m SWD 8.5 mØ x 1.5 m WD 0.40/0.733 m3/sec
Grit Drain 4 – 50 mm x 100 mm
Figure 9.15 EDI primary clarifier.
D H = SWd − feed well depth = 4.0 − 0.7 − 0.9 = 2.4 m (< 2.43 m d h) The feed-well depth of 0.7 + 0.8 m, or 1.5 m, is recommended to maximize the DH value to minimize scour. The effluent weir and launder will use a type C Lincoln/Stamford baffle design (Figure 9.12), the horizontal extension of which will be as follows: Stamford baffle = 460 mm + 16.7 mm/m (41 – 9, m) = 994 mm (use 1000 mm)
Clarification
3.9.1 Secondary (Integrated Fixed-Film Activated Sludge) Clarifiers The secondary clarifiers will be operated in a FGR/AS or IFAS mode, which will result in an activated sludge design procedure. If the clarifiers were operating on the fixed-growth effluent without any recycle, then the design would be similar to the primary clarifier sizing, except a simple sloped floor would be 1:12 m/m. Co-thickening of primary and fixed-growth sludge also would have been evaluated. The three secondary clarifiers will be designed on the basis of the CWZ above the maximum sludge depth. As shown in Table 9.6, the maximum available SWD is 4.6 m, and the dIV (Figure 9.8) allowance (preliminary) for peaking conditions is 0.7 m, leaving a CWZ of 4.6 minus 0.7, or 3.9 m. The design OFR will be determined, with the allowable SLR as a function of the SVI, RAS suspended solids (RSS), and MLSS concentration. OFRMAX = 0.556 (3.9 m) = 2.17 m/h OFRAVG = 0.278 (3.9 m) = 1.08 m/h Total AMAX =
(2.20 m 3 /s)(3600 s/h) (2.14 m/h)
= 3652 m2 Total AAVG =
(1.2 m 3 /s)(3600 s/h) (1.08 m/h)
= 4000 m2 (governs) The size of the clarifiers (preliminary) is as follows: 4000 m 2 φ= (3 units) (0.785) = 41.2 m
0.5
3 = 41.5 m diam × 4.6 m SWD. (2000 l/s) (3600 s/min) (1000 l/ m 3) (3) (42) 2 (0.785) = 1.73 m/h
oFr max =
451
452
Biofilm Reactors
With the preliminary sizing of the clarifiers completed, the initial assumption of 0.7 m depth (div) for the inventory shift at the maximum day flow can be evaluated. During average flow conditions, the sludge blanket would be at or below the SWD. Operating conditions for the clarifiers are developed in Table 9.7, and the inventory shift is determined with the 35 000-m3 aeration volume and three clarifiers online. The maximum day inventory shift would be 0.49 m or less than the initial 0.7-m assumption. With a CWZ of 4.1 m, the clarifier diameter would be reduced to 41.0 m (same as the primary clarifier), and the div value will be 0.50 m. The secondary clarifiers will be 3 – 41 mφ × 4.6 m SWD The check of the allowable OFRs with a CWZ of 4.6 to 0.5 m or 4.1 m is the aximum OFR—2.27 m/h (IWA, 1992)/2.00 m/h design and average OFR—1.28 m m/h (IWA, 1992)/1.09 design. The maximum day sludge loading rate is 7.67 kg/m2·d. Daigger (1995) projected that the underflow concentration (RSS) at this loading and a 120-mL/g SVI would be 13 000 mg/L. At the average loading of 5.27 kg/m2·d, the RSS would be 15 000 mg/L. A well-designed sludge collector will be capable of producing the 12 000-mg/L RSS used for design.
Table 9.7 Review of secondary clarifier operating conditions. Average day
Maximum day
Flow (m3/s)
0.400
0.733
RAS (m /s)
0.165
0.240
MLSS (mg/L)
3500
2960
RSS (mg/L)
12 000
12 000
SVI (mL/g)
≤120
<120
OFR (m/h)
1.07
1.96
5.27
7.67
0
6300
0
0.49
3
SLR (kg/m ·d) 2
Inventory shift (kg)
a
Inventory shift depth ( dI V ) , m
b
Inv. shift = (3.5–2.96 kg/m )(35 000 m )/3 = 6300 kg. Inv. shift depth = (6300 kg)/(12 kg/m3)(0.8)(42.5 m)2)(0.785) = 0.49 m.
a
b
3
3
Clarification
Center column (CC) and outlet ports: CC VMAX ≤ 0.6 m/s (Q+RASMAX) = 0.973 m3/s Ports = 125% of the area of the CC Area =
0.973 m 3 /s = 1.62 m2 0.60 m/s
CC Diameter = 1.44 m (use 1.45 mφ) Area of ports =
(1.25)(1.62 m 2) = 0.51 m2 each 4 ports
Center column design Diameter = 1450 mm Ports = 4 = 750 mm deep × 675 mm wide EDI: (Q + RAS)AVG/MAX = 0.60/0.973 m3/s Port VAVG/VMAX = 0.36/0.61 m/s (0.36 m/s governs) Detention ≥ 10 seconds Depth = 0.8 m Area of port =
0.60 m 3 /s = 0.208 m2 each (8)(0.36 m/s)
EDI design: Size = 40 m diam × 0.8 m water depth Ports = 8 = 700 mm deep × 300 mm wide Flocculating feed well (FFW): Average/maximum Q + RAS = 0.60/0.973 m3/s Average/maximum VV = 0.5/0.75 m/min (0.75 m/min governs) (0.973 m 3 /s)(60 s/min) Feed well φ = (0.75 m/min)(0.785)
0.5
= 9.6 m (use 10.0 m) The minimum FFW depth would be the EDI port depth of 0.70 m + 0.75 to 0.9 m clearance to the bottom of the feed well, or 1.45 m minimum, and the target depth is 1.6 m.
453
454
Biofilm Reactors
Minimum recommended clearance (DH) under the feed well is as follows: (10.0 m 2) (0.785) (0.9 V h / V V ) (10.0 m) (3.14) = 2.78 m
recommended D h =
Maximum DH available is as follows: Available DH = SWD – div – FFWD = 4.6 – 0.5 – 1.45 = 2.65 m Thus, the feed well will be 10.0 m diam × 1.45 m deep. The objective of having 0.9 VH/VV velocity at maximum dIV could not be achieved. Deepening the feed well will encourage peak-flow scouring of the sludge blanket because of higher VH values. The volume of the vertical projection from the 10.0-m-diam feed well to the 4.6-m SWD is 361 m3. This volume will provide 6.2 minutes of retention for MLSS flocculation at 0.973 m3/s and 15 minutes at a 0.6-m3/s average flowrate. The design will be a single weir mounted on the launder or the inside wall acting as a Lincoln/Stamford baffle. The horizontal projection (for the Stamford baffle) of the type C launder baffle (Figure 9.12) will be as follows: Stamford baffle = 460 mm + 1.67 mm/m(41 m φ - 9 m) = 994 mm (use 1000 cm) The clarifier floor would be transversed by two full- and two half-diameter spiral scrapers. The depth of the spiral scrapers would be tapered from 250 mm at the wall to 750 mm at the sludge hopper following the design procedure set forth by Albertson and Okey (1992). A variable-speed drive is suggested, with a tip speed range of 3 and 7 m/min. The scraper drive would operate at the higher speed when the blankets exceed the conical zone, which reduces the CWZ and clarification capacity. The transport requirements—and hence scraper speed—can change as a result of the SLRs and the SVI of the MLSS.
3.9.2 Sludge Hopper The sludge hopper in the 41-m-diam clarifier would process a maximum of 0.24 m3/s of return sludge. The clarifier is of intermediate size and would have a CWD-SWD
Clarification
of 1.7 m. The selected RAS velocity into the plane of the hopper (Figure 9.13) was 0.7 m/s (2.4 ft/sec), based on ≤1.5 (CWD-SWD). The hopper design is as follows: area = w=
(0.24 m 3 /s)(60 s/min) = 6 m2 2.4 m/min 6 m2 = 1.31 m 3.5 w
The hopper size is 1.3 m wide × 4.6 m long × 0.75 m deep. The 600-m-diam RAS pipe at a 0.85-m/s flowrate would have three uniformly spaced openings, each with an area of 94 200 mm2 (307 mm × 307 mm or equivalent area). Design procedures for rectangular and rapid sludge removal circular basins can be found in the referenced literature (for example, Water Environment Federation; 2005; Water Environment Federation et al., 2009). The omission of a procedure does not imply that these designs cannot provide equivalent effluent quality. The circular clarifiers are more common and scraper discharge to the central hopper will produce maximum underflow solids and minimize RAS flows at the equivalent effluent quality.
3.9.3 Process Performance The treatment capabilities of FGRs with improved clarifier designs represent a significant departure from historical data. This is particularly true for facilities that include a solids contact system after the FGR. A U.S. Environmental Protection Agency (1991) assessment of single-stage nitrification provides data on units typically constructed after 1970. Some of these reported results are summarized in Table 9.8. In general, the TF/SC and the trickling filter designs with good clarifier configurations produce an excellent effluent quality. The clarifier OFRs were not provided in the report. When trickling filters are in a nitrifying mode, clarification is good to excellent, but typically can be further improved with solids contact. When the trickling filter is in the carbonaceous BOD5 removal range and providing complete soluble BOD5 conversion, the TF/SC will likewise improve performance by enhanced suspended solids removal. When soluble BOD5 conversion is required in the dual biological process, then the TF/AS process will provide a stable system, with effluent quality of approximately 10 mg/L CBOD5 and 10 mg/L TSS.
455
456
Biofilm Reactors
Table 9.8 Effluent quality of single-stage trickling filter nitrification plants. Effluent quality (mg/L) Location
Process
Type media Depth (m) BOD5
TSS
NH4-N
Palm Springs, California
Trickling filter
Slag
2.9
7
9
0.5
Waconda, Illinois
Trickling filter/ SC
Plastic
8.5
<10
<5
0.1
Bremen, Ohio
Trickling filter
Plastic/rock
9.8/1.8
<10
<10
<2.7
Kenallville, Indiana Trickling filter
Rock/ plastic
1.7/2.0
10
5
<1.0
Rochester, Indiana
Trickling filter
Rock/ plastic
1.8/5.5
12
21
0.6
Amherst, Ohio
Trickling filter
Plastic
5.2
<10
<10
2.4
Youngstown, Ohio
Trickling filter/AS Plastic
4.8
5
10
0.3
Ashland, Ohio
Trickling filter/ SC
Plastic
9.1
6
7
1.5
Pickerington, Ohio
Trickling filter/ SC
Plastic
8.2
<2
<6
0.1
Buckeye Lake, Ohio Trickling filter/ SC
Plastic
12.8
2
5
0.3
Wauseon, Ohio
Trickling filter/ SC
Plastic
4.3
<10
<15
5.0
Allentown, Pennsylvania
Trickling filter
Plastic
9.8
12
11
5.3
Ozark, Alabama
Trickling filter
Plastic
6.1
<10
<10
<1.0
Boulder, Colorado
Trickling filter/ SC
Rock/ plastic
2.4/4.9
15
10
5.0
The Ohio Environmental Protection Agency (Columbus, Ohio) surveyed 15 newer trickling filter and TF/SC facilities and summarized 5 to 6 years of operating results. The purpose of the survey was to determine the nitrification capabilities of trickling filters operating in combined BOD5R and nitrification mode. The survey
Clarification
found that the 50th percentile values for effluent carbonaceous BOD 5, TSS, and NH4-N were less than what a mechanical treatment plant would be expected to achieve (best available developed control technology [BADCT]). Based on the survey results (Table 9.9) and BADCT criteria, it was recommended that 2 mg/L for summer NH4-N limits be used for a trickling-filter-only system—that is, without solids contact or sand filtration.
Table 9.9 Clarified/filtered effluent quality of Ohio nitrifying trickling filter plants. Final effluent (mg/L)
City/ county wastewater treatment organization
Period of record
Village of Shreve
1991–1996
8
9
0.51
0.48
Glendale
1991–1996
9
10
0.60
0.50
Leipsic
1991–1996
9
10
2.00
1.00
Zanesville
1991–1996
11.5
9
4.92
5.19
Average
–
9.4
9.5
2.00
1.80
Ashland
1991–1996
5
8
0.50
0.20
Licking County
1991–1996
4.2
7
0.08
0.09
Wauseon
1991–1996
2.5
6
0.31
0.38
Gallipolis
1991–1996
6.5
8.5
0.40
1.50
Springfield
1991–1995
3.5
9
0.06
0.43
Urbana
1991–1996
9.0
13
0.20
0.21
Fostoria
1991–1996
4.5
8.5
0.63
1.71
Blufflon
1991–1996
8.1
10.5
0.35
4.30
Average
–
5.4
8.8
0.32
1.10
Orville
1991–1995
1.3
1.5
0.08
0.09
Chardon
1991–1996
1.3
2
0.06
0.11
Norwalk
1991–1996
2.5
9
0.19
3.71
Average
–
1.7
4.2
0.11
1.30
NH4-H
Unit operationsa Solids contact
SF
CBODj
TSS
Summer Winter
457
458
Biofilm Reactors
The settled carbonaceous BOD5 and TSS results were 9.4 and 9.5 mg/L, respectively, for trickling-filter-only systems. With an add-on solids contact, the average effluent carbonaceous BOD5 was 5.4 mg/L, with a TSS of 5.4 mg/L. The use of a TF/SC process followed by sand filtration and an average effluent of 1.7 mg/L CBOD5 and 4.2 mg/L TSS was achieved. While these results are excellent, these systems typically do not have the latest improvements in secondary clarifier design. Additional information on TF/SC clarification performance is contained in papers by Parker (1983), Parker and Stenquist (1986), and Parker and Matasci (1989). These papers reflect improved secondary clarifier designs.
3.10 Other Considerations Modern clarifier design also is affected by the following: • Modeling, • Interaction with other facilities, and • International practices.
3.10.1 Modeling Chapter 6 of Clarifier Design, MOP FD-8 (Water Environment Federation, 2005), provides an overview of the state-of-the-art of clarifier modeling. A review of the applicable theory is given with discussions of the use of models in design, troubleshooting, and plant operations. All models are idealizations of reality and, as such, have built-in limitations. Thus, the results of all models need to be interpreted and treated with a certain amount of caution.
3.10.2 Interaction with Other Facilities Clarifiers are an inseparable and integral part of every conventional WWTP. Their performance efficiency is affected by the upstream wastewater collection and treatment facilities and has a significant effect on downstream biological treatment and solids-handling facilities. Chapter 12 of Clarifier Design, MOP FD-8 (Water Environment Federation, 2005), addresses the interaction of plant clarifiers and the other wastewater treatment and solids-handling processes.
3.10.3 International Practices Apart from site-specific innovations, design practice worldwide is similar, with differences being in emphasis rather than philosophy. Most of the regional differences
Clarification
between clarifier (or settling tanks, as they tend to be called) designs are either in the equipment fitted into the tanks, including sludge collection and inlet arrangements, or a preference for rectangular or circular tank designs. Multi-level clarifier floors also are more prevalent in parts of Europe, such as Germany. Chapter 11 of Clarifier Design, MOP FD-8 (Water Environment Federation, 2005), addresses many of these differences, with an emphasis on United Kingdom practices and preferences.
4.0 REFERENCES Albertson, O. E. (1991) Improving the Rapid Sludge Removal Collector. Proceedings of the 64th Annual Water Environment Federation Technical Exposition and Conference, Toronto, Ontario, Canada, Oct 7–10; Water Environment Federation: Alexandria, Virginia. Albertson, O. E.; Alfonso, P. (1995) Clarifier Performance Upgrade. Water Environ. Technol., 7 (3), 56–59. Albertson, O. E.; Hendricks, P. (1992) Bulking and Foaming Organism Control of Phoenix, AZ, WWTP. Water Sci. Technol., 26 (3), 461–472. Albertson, O. E.; Okey, R. W. (1992) Evaluating Scraper Designs. Water Environ. Technol., 4 (1), 62. Albertson, O. E.; Orris, E. (1994) Sludge Hopper Design for Activated Sludge Clarifiers. Proceedings of the National Conference on Environmental Engineering, Boulder, Colorado, July 11–13; American Society of Civil Engineers: Reston, Virginia. Albertson, O. E.; Scott, R. F.; Stensel, H. D.; Okey, R. W. (1992) Expansion and Upgrading of Columbus, Ohio WWTPs to Advanced Wastewater Treatment. Water Sci. Technol., 25 (4), 1. Albertson, O. E.; Walz, T. (1997) Optimizing Primary Clarification and Thickening. Water Environ. Technol., 9 (12), 41–45. Anderson, N. E. (1945) Design of Final Settling Tanks for Activated Sludge. Sew. Works J., 17, 50–65. Argaman, Y.; Kaufman, W. J. (1970) Turbulence and Flocculation. Am. Soc. Civ. Eng. J. Sanit. Eng. Div., 96, 223–241. Billmeier, E. (1988) The Influence of Blade Height on the Removal of Sludge from Activated Sludge Settling Tanks. Water Sci. Technol., 20 (4), 165–175.
459
460
Biofilm Reactors
Boyle, W. H. (1975) Don’t Forget Sidewater Depth. Water Waste Eng., 32. Brown and Caldwell (1978) Fixed Growth Reactors for the Municipality of Metropolitan Seattle. West Point Pilot Study, Vol. III; Brown and Caldwell: Walnut Creek, California. Buttz, J. (1992) Laguna WWTP—Secondary Clarifier Stress Test. Report to City of Santa Rosa, California; CH2M Hill: Oakland, California. Daigger, G. T. (1995) Development of Refined Clarifier Operating Diagrams Using an Updated Settling Characteristics Database. Water Environ. Res., 67, 95–100. Darby, W. A. (1939) Flocculation in Theory and Practice. Water Works Sew., June. Das, D.; Keinath, T. M.; Parker, D. S.; Wahlberg, E. J. (1993) Floc Breakup in Activated Sludge Plants. Water Environ. Res., 65, 138–145. Eimco Pmd. (1974) Technology and Development Report on Activated Sludge Flocculation (Internal Report). Eimco Pmd.: Salt Lake City, Utah. Ekama, G. A.; Barnard, J. L.; Gunthert, F. W.; Krebs, P.; McCorquodale, J. A.; Wahlberg, E. J. (1997) Secondary Settling Tanks: Theory, Modeling Design and Operation, IAWQ Scientific and Technical Report No. 6; International Association on Water Quality: London, United Kingdom. Erdman, A. (circa 1958) Internal Report on Effect of DEW/Dia Ratio on Clarification Efficiency. Dorr-Oliver Inc.: Stamford, Connecticut. Fischer, A. J.; Hillman, A. (1940) Improved Sewage Clarification by PreFlocculation without Chemicals. Sew. Works J., 12, 280–306. Great Lakes–Upper Mississippi River Board of State Sanitary Engineers (1971) Recommended Standards for Sewage Works. Health Education Services: Albany, New York. Great Lakes–Upper Mississippi River Board of State Sanitary Engineers (1990) Recommended Standards for Sewage Works. Health Education Services: Albany, New York. Günthert, F. W. (1984) Thickening Zone and Sludge Removal in Circular Final Settling Tanks. Water Sci. Technol., 16, 303–316. International Association on Water Quality (1992) Clarifier Design. In Process Design Manual for Biological Nutrient Removal; Technomic Publishing Company: Lancaster, Pennsylvania.
Clarification
LaMotta, E. J. (1976) Internal Diffusion and Reaction in Biological Flocs. Environ. Sci. Technol., 10 (8), 765–769. Norris, D. P.; Parker, D. S.; Daniels, M. L.; Owens, E. L. (1982) High Quality Trickling Filter Treatment without Tertiary Treatment. J. Water Pollut. Control Fed., 54, 1087–1098. Ozinzky, A. E.; Ekama, G. A.; Reddy, B. D. (1994) Mathematical Simulations of Dynamic Behaviour of Secondary Settling Tanks, Research Report W85. University of Cape Town, South Africa. Parker, D. S. (1983) Assessment of Secondary Clarification Design Concepts. J. Water Pollut. Control Fed., 55, 349–359. Parker, D. S. (1991) The Case for Circular Clarifiers. Water Eng. Manage., 22, April. Parker, D. S.; Kaufman, W. J.; Jenkins, D. (1970) Characteristics of Biological Flocs in Turbulent Regimes, Rep. No. 70-5. Sanitary Engineering Research Laboratory, University of California: Berkeley. Parker, D. S.; Kaufman, W. J.; Jenkins, D. (1972) Floc Breakup in Turbulent Flocculation Processes. ASCE J. Sanit. Eng. Div., 98, 79–99. Parker, D. S.; Kaufman, W. J.; Jenkins, D. (1971) Physical Conditioning of Activated Sludge Floc. J. Water Pollut. Control Fed., 43, 1817–1833. Parker, D. S.; Matasci, R. N. (1989) The TF/SC Process at Ten Years Old: Past, Present and Future. Proceedings of the 62nd Annual Water Environment Federation Technical Exposition and Conference, San Francisco, California, Oct 15–19; Water Environment Federation: Alexandria, Virginia. Parker, D. S.; Stenquist, R. J. (1986) Flocculator–Clarifier Performance. J. Water Pollut. Control Fed., 58, 214–219. Semon, J., City of Stamford, Connecticut (1982) Personal communication. Stukenberg, J. R.; Rodman, L. C.; Touslee, J. E. (1983) Activated Sludge Clarifier Design Improvements. J. Water Pollut. Control Fed., 55, 341–348. U.S. Environmental Protection Agency (1991) Assessment of Single-Stage Trickling Filter Nitrification, EPA-430/09-91-005; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C. U.S. Environmental Protection Agency (1987) Design Manual: Dewatering Municipal Wastewater Sludges, EPA-625/1-87-014; U.S. Environmental Protection Agency, Office of Research and Development: Cincinnati, Ohio.
461
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Biofilm Reactors
Wahlberg, E. J.; Keinath, T. M.; Parker, D. S. (1992) Influence of Activated Sludge Flocculation on Secondary Clarification. Proceedings of the 65th Annual Water Environment Federation Technical Exposition and Conference, New Orleans, Louisiana, Sept 20–24; Water Environment Federation: Alexandria, Virginia. Wahlberg, E. J.; Peterson, M. A.; Flancher, D. M. (1993) Field Application of the CRTC’s Protocol for Evaluating Secondary Clarifier Performance: A Comparison of Sludge Removal Mechanisms in Circular Clarifiers. Presented at the Rocky Mountain Water Pollution Control Association Annual Meeting, Albuquerque, New Mexico, Sept 20; Rocky Mountain Water Pollution Control Association: Albuquerque, New Mexico. Walker Equipment Company (1953) ClariFlow Inlet Device. U.S. Patent 2,635,757. Warden, J. H. (1981) The Design of Rakes for Continuous Thickeners. Filtr. Sep., 18 (2), 113–116. Water Environment Federation (2005) Clarifier Design, 2nd ed., Manual of Practice No. FD-8; McGraw-Hill: New York. Water Environment Federation (1992) Problem Solvers—Baffles Reduce Clarifier Turbulence. Water Environ. Technol., 4 (12), 64. Water Environment Federation; American Society of Civil Engineers; Environmental and Water Resources Institute (2009) Design of Municipal Wastewater Treatment Plants, 5th ed., WEF Manual of Practice No. 8, ASCE Manuals and Reports on Engineering Practice No. 76; McGraw-Hill: New York. Watts, R. W.; Svoronos, S. A.; Koopman, B. (1996) One Dimensional Modeling of Secondary Clarifiers Using Concentrations and Feed Velocity Dispersion Coefficient. Water Res., 30, 2112–2124. Wilson, T. E. (1991) Rectangular Clarifiers Should Be Considered. Water Eng. Manage., 138 (4), 20.
Chapter 10
Effluent Filtration 1.0 INTRODUCTION
463
3.0 References
469
2.0 Process performance 465
1.0 INTRODUCTION Effluent filtration is used to polish the clarified effluent from biological (e.g., s uspended-growth, fixed-growth, and dual biological systems) and chemical precipitation processes. Effluent filtration has been specified when an effluent total suspended solids (TSS) limit of <10 mg/L (monthly average) is required, when very low effluent nitrogen and/or phosphorus limits are specified, for effluent reuse applications (sometime mandated by state regulatory agencies), and to improve effluent turbidity for more effective disinfection. There are several types of filtration systems, such as shallow-bed (traveling-bridge), deep-bed, moving-bed, and disc-type filters, where the flow may be up, down, radial, continuous, or discontinuous. Membrane filters also are becoming more common and are used in situations for indirect and direct reuse. Complete descriptions of each technology are in the design manuals and reports from Metcalf and Eddy (2003), U.S. Environmental Protection Agency (1974, 1977), and Water Environment Federation et al. (2009). The purpose of this section is to provide general information on effluent filtration and issues inherent with fixed-film biological treatment systems and not to provide detail design criteria and operating parameters, which are available in these other resources. Early applications of granular-media filters for wastewater treatment essentially followed the design procedures developed for the treatment of potable water. 463
464
Biofilm Reactors
Because wastewater is significantly different in physical and chemical characteristics from most natural waters, wastewater filtration entails special design considerations. In general, wastewater filters receive larger, heavier, and more variable particle sizes and uneven solids loadings. Filtration mechanisms are complex and may consist of a combination of factors, including straining (mechanical and chance contact), interception within media, gravity settling, inertial impaction of the particles with adhesion to the filtering medium, and growth of biological solids within the filter bed, which further enhances solids removal. Because the performance of wastewater filters is affected by many factors, a designer should consider pilot studies in cases where strict effluent quality limits must be met or for applications not previously experienced. Currently, there are numerous filter technologies that are being brought to the market and finding favor with plant engineering and operations and maintenance staff. These technologies include the use of “cloth” media systems constructed in a variety of configurations. Filters often are referred to by the type of media, configuration of flow, and/or type of backwashing device. The solids-removal filtration alternative evaluation consists of a qualitative and quantitative analysis to determine the recommended technology for solids removal. The following solids-removal filtration technologies currently are being considered by engineers and plant operators: • Traveling bridge, • Continuous backwash, • Deep-bed downflow, • Cloth, • Compressible media, • Mono media, • Dual media, and • Membrane. As noted in this list, the description of the different technologies sometimes refers to the type of media used, mechanism for backwashing, or configuration of the filter. While effluent sand filtration technology is well-developed, there are so many types of filters, media, and operating modes that technology is too complex and diverse to be summarized in this manual. Some examples of different filters are shown in Figures 10.1 to 10.3.
Effluent Filtration
Disk
Drive Motor
Overflow Weir Effluent Weir
Overflow Valve Influent Weir
Backwash Assably
Backwash Waste Solids Valve
Backwash Solids Backwash/ Valve Collection Solids Pump Manifold
Figure 10.1 Cloth filters (courtesy of Aqua–Aerobics).
Discussion of effluent filters in this manual will be restricted to managing the solids from fixed-film processes.
2.0 Process performance Sand filtration can be expected to remove 50 to 80% of the influent TSS without chemical addition and 75 to 90% with chemical addition. Polymer alone or metal salts in combination with polymers are the typical chemicals used. The improved performance of clarifiers results in lower TSS feed concentrations, and the filtered effluent quality is enhanced. Washwater collection pumps Influent ports Influent chamber
Control panel Fiberclass lumber
Polyprophylene wear strip Backwash pump Backwash shoe Individual cell effluent/backwash port and module header Granular media-typically 11” sand single media dual media and activated carbon also aviailable Porous plate media support Top support retainer
Fiberglass cell divider sheet
Effluent chamber clearwell
Figure 10.2 Traveling-bridge filter (courtesy of Aqua–Aerobics).
465
466
Biofilm Reactors
Actuator for Movable Plate
Actuator for Movable Plate
Movable Plate
Movable Plate
Effluent Compressible Media
Fixed Plate Influent Filtration Cycle
Flush Water
Compressible Media
Fixed Plate Influent
Actuator for Movable Plate
Wash Water
Movable Plate
Flush Cycle
Compressible Media
Fixed Plate Influent (washing Water)
Washing Air Wash Cycle
Figure 10.3 Compressible media filters (courtesy of Schreiber).
The removal of TSS removes a portion of the carbonaceous 5-day biochemical oxygen demand (BOD5). The quantity of carbonaceous BOD5 removed is a function of the carbonaceous BOD5 associated with the effluent TSS that can be removed through the filter. Depending on the upstream treatment process loading and whether chemical coagulants are used, the associated carbonaceous BOD5-to-TSS ratio can range from 0.3 to 0.8 mg/mg. Effluent quality depends on influent quality, as illustrated by the results of six different types of pilot filters on activated sludge effluent. As shown in Figure 10.4 (Metcalf & Eddy, 1991), all filters produced similar results with a high-quality influent. However, with decreasing influent quality, the effluent quality deteriorates, both in terms of percent removal and effluent turbidity. As shown in Figure 10.4, some types of filters perform better on poorer-quality influent than others. Generally, the TSS of fixed-growth reactors have been more difficult to filter than the effluent TSS from activated sludge systems. Historically, the clarified trickling filter TSS have been higher than newer facilities with good clarifier configurations. Highly loaded oxygen-limited or partially fouled trickling filters produce particles
Effluent Filtration
12
Effluent turbidity reading, NTU
10 8 6 4
2 0 0
5
10 15 20 Influent turbidity reading. NTU
25
30
Figure 10.4 Performance data for six different types of granular-medium filters tested on the same activated sludge effluent at a filtration rate of 2.7 L/m2 s (4 gpm/sq ft) (Metcalf and Eddy, 1991).
that are more difficult to remove than those particles produced by fully aerobic, lower loaded fixed-growth reactors (FGRs). Anaerobiosis in the FGR will produce finer, difficult-to-capture particulates. Solids settleability issues with FGR units have been recognized for many years at some facilities that have witnessed poor effluent solids from the final clarifiers. These same solids are also troublesome for removal through sand/anthracite media filters. Filters downstream of fixed-film facilities also needed to contend with the breakthrough of some soluble carbonaceous BOD5. While many trickling filter and rotating biological contactor (RBC) facilities achieved 85% removal of carbonaceous BOD5, some soluble substrate would be discharged to the downstream effluent filters. With aerobic conditions existing in the filter bed, growth of heterotrophic organisms can occur, thus increasing the solids accumulation and potential for “mudballing” within an effluent filter. Mudballing refers to a condition within a granular media where the media clumps together and forms balls of media (from 50 to 200 mm). This decreases the filter performance and allows short-circuiting of the filter influent. While
467
468
Biofilm Reactors
mudballing can occur as a result of poor backwashing and the passage of scum, oil, and grease from the upstream biological treatment process, close examination of the effluent quality from the upstream fixed-film process unit is required to understand whether excess biological growth may occur in the downstream filter unit. Plant operators have reported that backwashing the filters once per day, whether the filters required it or not, was one successful means to avoid mudballing. Also, some facilities will add a small maintenance dose of chlorine (1 mg/L chlorine) to the filters to reduce the potential for biological growth and improve the downstream chlorination efficiency. This practice does raise the concern of increasing disinfection byproducts in the plant effluent. Improved performance of the filters can be realized using the same techniques to address clarifier performance via the use of a solids contact system following the FGR. Combined fixed-film and suspended-growth systems typically will have a better solids capture efficiency and, thus, improved effluent filterability. Improving the filterability of the influent solids through the use of inorganic, organic, or both types of coagulant aids upstream of the sedimintation tank and, if necessary, before the filter, can improve the effluent filterability. Critical to success are good mixing and flocculation zones before the sedimentation and filter units, to generate sufficient particle size of solids for removal. This also will decrease the chemical dosing requirements. Pierce (1978) reported on the plant performance of four types of sand filters used to upgrade trickling filter effluents. These results are presented in Table 10.1.
Table 10.1 Full-scale filtration of settled trickling filter effluent. Sand filter influent
Sand filter effluent BOD5 (mg/L)
Type
Dosing rate (m/h)
BOD5 (mg/L)
TSS (mg/L)
Pressure-mixed media
5.43
31
19
22*
7
Backwashed gravity-mixed media
6.88
16
15
5
2
Backwashed gravity-sand
4.09
24
17
4
5
Backwashed intermittent-sand
0.022
40
8
40
9
*High BOD5 may be because of nitrification in the BOD5 test.
TSS (mg/L)
Effluent Filtration
Weaver (1989) summarized the experiences of upgrading the Cibolo Creek wastewater treatment plant (Randolph, Texas) with tertiary sand filters. The effluent quality averaged 2.4 mg/L BOD5, 2.0 mg/L TSS, and 0.86 mg/L ammonium-nitrogen (NH4-N). The Wauconda, Illinois, facility (U.S. EPA, 1991) average effluent BOD5, TSS, and NH4-N values were <10, <5, and <0.1 mg/L, respectively, following sand filtration. Lin and Heck (1987), reporting on a trickling filter–solids contact process at Wauconda, Illinois, noted that nitrification was occurring in the effluent sand filters. Boller and Gujer (1986) also described additional nitrification occurring in tertiary sand filters following nitrifying trickling filters. They concluded that up to 1.7 mg/L NH4-N could be oxidized before the process became oxygen-limited in a 1400-mmdeep filter. The dissolved oxygen in the influent would need to be approximately 7 to 8 mg/L to achieve oxidation of 1.7 mg/L NH4-N, as shown in Figure 10.4. High nitrification rates (700 to 900 mg/m3·d) were not adversely affected by intense air or water backwashing. Sand filtration can readily upgrade good trickling filter and RBC effluent to less than 5 mg/L BOD5 and TSS when advanced wastewater treatment is required. Trickling filter–solids contact with effluent sand filtration long-term results from three plants (see Table 9.9) reported by the Ohio Environmental Protection Agency (1997) averaged 1.7 mg/L carbonaceous BOD5 and 4.2 mg/L TSS. For non-nitrifying systems, an effluent of 10 to 15 mg/L BOD5 and 5 to 10 mg/L TSS is achievable with effluent sand filtration.
3.0 References Boller, M.; Gujer, W. (1986) Nitrification in Tertiary Trickling Filters Followed by Deep Bed Filters. Water Res., 20, 1363–1373. Lin, C. S.; Heck, G. (1987) Design and Performance of the Trickling Filter/Solids Contact Process for Nitrification in a Cold Climate. Proceedings of the 60th Annual Water Environment Federation Technical Exposition and Conference, Philadelphia, Pennsylvania, Oct 4–8; Water Environment Federation: Alexandria, Virginia. Metcalf and Eddy, Inc. (2003) Wastewater Engineering: Treatment and Reuse, Tchobanoglous, G., Burton, F. L., Stensel, H. D. (Eds.); McGraw-Hill: New York. Metcalf and Eddy, Inc. (1991) Wastewater Engineering: Treatment, Disposal and Reuse, 3rd ed.; McGraw-Hill: New York.
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Ohio Environmental Protection Agency (1997) Trickling Filter Alternative BADCT. Internal communication from D. S. Rector to P. Novak, April 7; Ohio Environmental Protection Agency: Columbus, Ohio. Pierce, D. (1978) Upgrading Trickling Filters, Tech. Rep. MCD-42, EPA-430/9-78004; U.S. Environmental Protection Agency: Washington, D.C. U.S. Environmental Protection Agency (1991) Assessment of Single-Stage Trickling Filter Nitrification, EPA-430/09-91-005; U.S. Environmental Protection Agency, Office of Wastewater Management: Washington, D.C. U.S. Environmental Protection Agency (1974) Wastewater Filtration Design Considerations. Technology Transfer Seminar Publication; U.S. Environmental Protection Agency: Washington, D.C. U.S. Environmental Protection Agency (1977) Wastewater Filtration Design Considerations. Technology Transfer Seminar Publication; U.S. Environmental Protection Agency: Washington, D.C. Water Environment Federation; American Society of Civil Engineers; Environmental and Water Resources Institute (2009) Design of Municipal Wastewater Treatment Plants, 5th ed., WEF Manual of Practice No. 8; ASCE Manuals and Reports on Engineering Practice No. 76; McGraw-Hill: New York. Weaver, T. G. (1989) Sand/Anthracite Filtration Complements Trickling-Filter Systems. Water Eng. Manage. Sept, 47.
Chapter 11
Development and Application of Models for Integrated Fixed-Film Activated Sludge, MovingBed Biofilm Reactors, Biological Aerated Filters, and Trickling Filters 1.0 INTRODUCTION
473
2.0 MODELING
478
2.1 Numerical Approach Using Semi-Empirical Equations for Biofilm (Steady-State and Dynamic Simulation) 2.1.1 AmmoniumNitrogen Uptake Rate
478 480
Nitrification Rates from Pilot Studies 482 2.1.1.3 AmmoniumNitrogen Uptake Rate by Nitrifiers in Mixed-Liquor Volatile Suspended Solids 485 2.1.1.4 Mass Balance for Ammonium-Nitrogen in Each Reactor 490
2.1.1.1 AmmoniumNitrogen Uptake Rate by Nitrifiers in Biofilm 480
2.1.2 Chemical Oxygen Demand Removal
2.1.1.2 Biofilm
2.1.3 Biomass Production 499 471
494
(continued)
472
Biofilm Reactors
2.1.3.1 Mixed-Liquor Volatile Suspended Solids 500
3.1.2 Integrated Fixed-Film Activated Sludge Plant Operation 515
2.1.3.2 Biofilm
3.1.2.1 Data from December 2006
501
2.1.4 Fraction of Nitrifiers 502 2.1.5 Denitrification
503
2.1.6 Oxygen
504
2.2 Numerical Approach to Solve One- and TwoDimensional BiofilmDiffusion Models (Steady-State and Dynamic Simulation) 504 2.2.1 AmmoniumNitrogen
506
2.2.2 Linkage to Equations 11.1 to 11.42 Presented Earlier 508 2.2.3 Chemical Oxygen Demand, Biomass (Volatile and Total Suspended Solids), Dissolved Oxygen, and NOx-N 510 2.2.4 Biofilm Thickness, Growth, and Fraction Nitrifiers 510 3.0 MODEL APPLICATIONS TO FULL-SCALE FACILITIES 512 3.1 Integrated Fixed-Film Activated Sludge Plant Description and Modeling
515
3.1.2.2 Flow and Recycle 515 3.1.2.2.1 Primary Effluent
515
3.1.2.2.2 Aerobic Cells
518
3.1.2.2.3 Secondary/ Plant Effluent 518 3.1.2.2.4 Discussion of the Data 518 3.1.3 Modeling Integrated Fixed-Film Activated Sludge in Aquifas 518 3.1.3.1 Results from Aquifas 519 3.1.3.2 Key Inputs to Aquifas Biofilm OneDimensional Model 526 3.1.3.3 Discussion of Aquifas Model and Accuracy of Results 526 3.1.4 Modeling in BioWin 527 3.1.4.1 Framework 527
513
3.1.4.2 Results from BioWin 531
3.1.1 Integrated Fixed-Film Activated Sludge Plant Description 514
3.1.4.3 Discussion of Results from BioWin 531 (continued)
Development and Application of Models
3.2 Moving-Bed Biofilm Reactor Plant Description and Modeling 534 3.2.1 Moving-Bed Biofilm Reactor Modeling with GPS-X 537 3.2.1.1 Introduction 537 3.2.1.2 Example 542 3.2.2 Moving-Bed Biofilm Reactor Modeling with Aquifas 542
3.2.3 Moving-Bed Biofilm Reactor Modeling— General Comments 543 3.2.4 Integrated Fixed-Film Activated Sludge and Moving-Bed Biofilm Reactor Modeling—General Observations 552 4.0 REFERENCES
553
1.0 INTRODUCTION The first portion of this chapter discusses how activated sludge and biofilm models can be integrated to create models for integrated fixed-film activated sludge (IFAS) and moving-bed biofilm reactor (MBBR) systems. The parameters used are based on the International Water Association (London, United Kingdom) (IWA) activated sludge model and the IWA Task Group on Biofilm Modeling (Wanner et al., 2006). The second portion of this chapter shows how models can be applied to simulate IFAS and MBBRs. The modeling of trickling filters is found in Chapter 3. The modeling of biofilm reactors is more complex compared with activated sludge models. This is principally because of the introduction of diffusion. Further, the modeling of IFAS is more complex than biofilm “pure” MBBRs, where it can be assumed that, for low-strength wastewaters, the mixed-liquor volatile suspended solids (MLVSS) concentration is so low that there is very little removal of chemical oxygen demand (COD), nitrification, and denitrification by the suspended solids (Figure 11.1). However, this assumption is not valid for MBBRs operating with moderate- to high-strength wastewaters observed in some of the developing economies around the world. In these MBBRs, where MLVSS levels can be between 200 and 500 mg/L, there is a need to integrate removals in the suspended solids with the removals in the biofilms. The next level of complexity is the introduction of multiple cells in series in a reactor in which different cells can be operated under different conditions (e.g., anaerobic, anoxic, aerobic, and post-anoxic with second substrate). A typical process configuration is shown in Figure 11.2. The addition of recycle streams, such as nitrate
473
474
Biofilm Reactors
Primary Effluent
Secondary Effluent 1A. Activated Sludge
1A RAS
IB. IFAS – fixed bed Eg: Ringlace, Bioweb(cord)
1B
Screen 1C
Screen 1D
IC. IFAS – Moving Bed Eg: Captor, Linpor(sponge) Kaldnes, Hydroxyl, Entex (plastic)
ID. MBBR Kaldnes, Entex, other (plastic)
Clarifier size requirements decrease from 1A to ID Operating MLSS decrease from 1A to ID HRT requirements decrease from 1A to ID Biofilm media surface area requirements increase from 1A to ID
Figure 11.1 Plant configurations (Sen and Randall, 2008a) (manufacturers not mentioned in text are as follows: Ringlace [Ringlace Products Inc., Troutdale, Oregon], Bioweb [Entex Technologies, Raleigh, North Carolina], Hydroxyl [HeadworkUSA, Houston, Texas]).
recycle and return activated sludge (RAS), under multicell simulations, requires the use of numerical methods and modeling techniques that solve for the effluent iteratively instead of analytical techniques that can be applied to simpler biofilm reactors. Within the realm of biofilm modeling, there are different levels of complexity. The simplest models use a set of equations that can be solved analytically. The analytical simplifies the set of differential equations, but require more assumptions, such as knowledge of the rate-limiting substrate in each cell within a reactor. Also, it assumes knowledge of the limiting substrate in the layers within the biofilm. The numerical onedimensional model relaxes these assumptions (Reichart, 1998; Sen and Randall, 2008b). The one-dimensional model can be extended to multiple dimensions. Some researchers
Development and Application of Models
IFAS Tank and Media Configuration 10 reactors (cells) in Series Anaerobic
Anoxic
Aerobic
Post Anoxic
Influent
Reactor #
Reair
Supplemental Carbon
1
2
3
4
5
6
7
8 9 10
RAS Except in MBBR Mode
Figure 11.2 Schematic of IFAS and MBBR system with 10 cells (reactors) in series operating in enhanced nutrient removal configuration (adapted from Sen and Randall, 2008c). treat the length of a reactor or a multicell reactor as the second dimension (also referred to as a pseudo-2D). Others treat the spatial diversity within the biofilm as a second dimension, wherein the biofilm is allowed to grow and cover more of the media surface. There is recognition that incorporating the concepts of two-dimensional modeling is important in IFAS and MBBR media systems, where the thickness of the biofilm can lead to loss in specific surface area on the inside surface of plastic cylinders, which may be partially compensated for by the increase in the biofilm that develops on the outer surface. This chapter discusses the one- and two-dimensional modeling of the biofilm. One of the difficulties in applying the one- and two-dimensional modeling of biofilms is the need to apply a finite difference of finite element method. This slows down the computation. One strategy that can be applied to overcome this is to develop semi-empirical equations of the kinetics of COD removal, nitrification, denitrification, and biomass production by the biofilm operating within a multicell system, which apply terms to simulate the effect of boundary layer (stagnant liquid layer) thickness, liquid and biofilm-diffusion coefficients, biofilm thickness, and substrate utilization rates. The experimental studies are complex because the data required to develop the equations require operating the biofilm under different COD, ammonium-nitrogen (NH4-N), and COD/NH4-N conditions. This can be done by operating a multicell reactor (recommend three cells in series in the aerobic zone) under different mixed-liquor
475
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Biofilm Reactors
suspended solids (MLSS) mean cell residence times (MCRTs) ranging from MLSS MCRT = hydraulic retention time (HRT) for MBBRs to a series of higher MLSS MCRTs for IFAS. The substrate uptake rates must be measured for the biofilm that develops, and a set of equations can be developed for the biofilm (Sen, 1995; Sen and Randall, 2008a, 2008c; Sen et al., 2000; Sriwiriyarat et al., 2005). This approach then is used in models for fluxes of substrates and biomass in the biofilm in IFAS and MBBRs that can be solved much faster. The effect of variation in the thickness of the stagnant liquid layer on the flux rates can be incorporated by using multiple Monod terms to simulate the relationships between the maximum substrate uptake rates in the semi-empirical equations with the turbulence and media design, both of which can affect the thickness of the boundary layer and the biofilm. The effect of substrate concentrations in layers of the biofilm, fraction nitrifiers, and biofilm thickness, as influenced by external electron acceptor and electron donor concentrations, is incorporated to a second Monod expression. The equations for substrates and electron acceptors (COD, NH4-N, dissolved oxygen [DO], NOx-N , phosphorus, and alkalinity) and the biomass forms (autotroph, heterotroph, and inerts yield) for the activated sludge model operate synchronously with the supplemental set of equations for the biofilm. Finally, it is worth noting that both approaches to modeling lend themselves to trickling filters and biological filters. Conceptually, this is shown in Figure 11.3. Within the models, the trickling filters are treated as MBBRs with biofilm (effective)
4
TF Segment #s modeled in series
5 6
“BOD” Removal Filter
“Nitrifying” Filter
TF/AS Tank
(a) Models allow trickling filter systems to be modeled as multiple segments per filter, filters in series, and filters in series with activated sludge.
Figure 11.3 Modeling biofilm in (a) trickling filter, (b) MBBRs, and (c) BAF.
Development and Application of Models
External Substrate Nitrate Recycle
Primary Effluent
Clarifier PreAnoxic
Aerobic
Post- Re Anoxic Air
Floc Tank
Optional Low RAS to avoid need for external flocculent MBBRs can use low RAS and at their operating MLSS, do not need coagulant or floc tank, MBBR do not use RAS; in most instances, they need a floc tank or static mixing to coagulate fine particulates TM
(b) Model allows MBBRs to be modeled with cells in series recycles, and external substrate feed. External Substrate (C)
4 5 6
Pre-anoxic Biological Filter
“Aerobic” Filter 1
Aerobic Filter 2
PostAnoxic Filter
Optional Nitrate Recycle
(c) Models allow biological filters (BAFs and SAFs) to be modeled with multiple segments per filter, filters in series, with recycles and external substrate feeds.
Figure 11.3 Continued. specific surface areas of 20 to 100 m2/m3, compared with 150 to 400 m2/m3 in MBBRs. The biological filters (aerated and anoxic) have specific surface areas of 800 to 1000 m2/m3 (Table 11.1). Both approaches have the versatility to handle filters in series (e.g., pre-anoxic, aerobic, and post-anoxic) and can model filters where the substrate and electron acceptor conditions vary along the height of the filter. Additionally, one
477
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Biofilm Reactors
Table 11.1 Biofilm specific areas and modeling trickling filter, TF/AS, MBBR, and BAFs within models. Biofilm specific surface area (m2/m3)
Direction of flow and modeling as multiple cells
Trickling filter (rock)
10 to 25
Vertical; if multiple filters in series, also horizontal
Trickling filter (plastic)
20 to 50
Vertical; if multiple filters in series, also horizontal
MBBRs
200 to 400
Horizontal
BAF
800 to 1200
Vertical; if multiple BAF in series, also horizontal
Process*
*All processes can have recycles.
can handle recycles of RAS through trickling filters, trickling filter/activated sludge (TF/AS) processes, and nitrate recycles.
2.0 MODELING In this section, the semi-empirical model for multicell IFAS and MBBR systems is presented first. Because of the inclusion of multiple cells and recycles, the model must be solved numerically. The biofilm one- and two-dimensional models are presented next.
2.1 Numerical Approach Using Semi-Empirical Equations for Biofilm (Steady-State and Dynamic Simulation) The basic premise of the semi-empirical model is to add equations for uptake, oxidation, and reduction of COD, NH4-N, oxidized nitrogen, and sludge production in biofilms in a format that is compatible with the IWA activated sludge model. The COD, NH4-N, and oxidized nitrogen uptake and removals are computed as the sum total of removals by the MLSS and the biofilm. The percentage of removal in the biofilm increases as the amount of biofilm surface area is increased, and the MLSS MCRT (and MLVSS) is decreased. The equations to compute the COD removal, NH4-N removal (biomass uptake for synthesis and nitrification), and denitrification are developed based on Monod process kinetics. Additional removal resulting from the biofilm in IFAS and MBBR configurations is computed based on removal rates per unit of surface area of biofilm multiplied by the surface area of biofilm in each reactor (cell).
Development and Application of Models
The model is set up to operate with up to as many cells in series (n cells) as required to simulate the plant. Influent and recycles can be fed to and removed from any cell. Biofilm support media can be installed in one or more cells. Each cell can be operated with or without aeration, as part of an anaerobic, pre- or postanoxic, aerobic, or reaeration zone. For the un-aerated cells, the model computes the dissolved oxygen and oxidized nitrogen (NOx-N ). This is compared with userspecified thresholds for aerobic, anoxic, and anaerobic conditions, to determine whether aerobic, anoxic, or anaerobic decay rates of biomass should be used for each cell. The semi-empirical version of the biofilm model incorporates Monod-like equations for the biofilm for substrate uptake and removal under aerobic and anoxic conditions. These coefficients of the equations are based on either (a) experimental measurements of biofilm flux rates observed in pilot studies or (b) calibrated by linking to and running a biofilm diffusional model to generate results that replicate those in the diffusional model under one or more conditions. The maximum nitrification rate per unit surface area of biofilm (qm,NH4N-Nitr,bf) and the maximum COD utilization rate per unit surface area of biofilm under aerobic and anoxic conditions (qmH, COD,bf,aer and qmH, COD,bf,anx) are quantified. The half-saturation constants for substrate and dissolved oxygen for the biofilm (KN,bf and KDO,bf) are determined through a separate model calibration. The methodology for quantifying rates was published in Sen (1995) and is summarized below. The biofilm yield in each cell is determined from a table that shows expected yields for heterotrophs and autotrophs under aerobic and anoxic conditions at different concentrations of soluble biodegradable COD (SCODbio) and NH4-N in the bulk liquid. This table is developed by running the biofilm one- and two-dimensional models for different conditions and determining the yields. Table 11.2 shows an example of such a table. The semi-empirical approach and its equations are used in a number of other models for IFAS and MBBR systems (e.g., those developed by Aqwise Wise Water Technologies [Herzliya, Israel], who developed the agar media, and Aquifas [Aquaregen, Mountain View, California]). Several media and tank manufacturers (Bioportz media [Entex Technologies, Chapel Hill, North Carolina], EEC USA [Newport Beach, California], and others) use this approach to model their systems. The biofilm-diffusion model is used in AQUASIM (Swiss Federal Institute of Aquatic Science and Technology (EAWAG, Dübendorf, Switzerland), Aquifas 4, BioWin (Envirosim Associates, Ltd., Hamilton, Ontario, Canada), and GPS-X (Hydromantis, Inc., Hamilton, Ontario, Canada.
479
480
Biofilm Reactors
Table 11.2 Typical biofilm yields. Yields are generated by running a biofilm onedimensional model or from measurements in pilot studies. Suggested yields for biofilm
Substrate concentration range
Mode of uptake
Based on Aquifas 4 runs or from experiments Lower
Upper
Anaerobic Anoxic
mg/L
mg/L
mg VSS/mg COD by biofilm or mg VSS/NH4-N uptake
20
200
0.25
0.35
5
19
0.2
0.25
1
4.9
0.18
0.22
0.1
0.99
0.16
0.2
5
20
0.5
1
4.9
0.045
0.1
0.9
0.04
SCODbio
NH4-N range
Aerobic
For the kinetics parameters, a comprehensive IFAS, MBBR, and activated sludge model should list default values for kinetic parameters. The default values are based on pilot studies and calibration of the model to full-scale plants. The user should have the option to modify the default values based on plant-specific data.
2.1.1 Ammonium-Nitrogen Uptake Rate Equations 11.1 to 11.16 show how NH4-N can be computed in a semi-empirical version of the model.
2.1.1.1 Ammonium-Nitrogen Uptake Rate by Nitrifiers in Biofilm Equation 11.1 shows the NH4-N uptake rate in the biofilm, BN,n (kg/d). The NH4-N uptake rate is the sum total of NH4-N uptake by nitrifiers for synthesis and nitrification. For cell n,
BN,n = q m,NH4 N − Nitr,bf
SO 2n SN n Vn Mn S O + K DO,bf 2n K N,bf + SN n
(11.1)
Where qm,NH4N-Nitr,bf = flux rate for NH4-N uptake by the nitrifiers (expressed as kg/1000m2 of biofilm surface/d, or mg/cm2/d).
Development and Application of Models
The value of qm,NH4N-Nitr,bf for the biofilm is adjusted for mixed-liquor temperature. This is done using the Arrhenius equation, with a temperature adjustment coefficient, θ, as follows:
qm,NH4N-Nitr,bf,T = qm,NH4N-Nitr,bf,25 θ
The literature on nitrification in activated sludge systems shows a range of values of the coefficient θ These vary from 1.03 to greater than 1.07 (Dold, 1991; Marais and Ekama, 1976; Randall et al., 1992; Wentzel et al., 1991). Research conducted by Weiss et al. (2005) using an MBBR showed a temperature coefficient of 1.047. A value of 1.05 +/− 0.02 is recommended for the biofilm. For the modified version of eq 11.1 (below), SN n and SO2 n = NH4-N and dissolved oxygen concentrations, respectively, measured in the liquid outside the biofilm in cell n (i.e., as measured in activated sludge systems); Vn = volume of cell n (m3); and Mn = m2 of biofilm surface area per m3 of cell volume in cell n. KN,bf = half-saturation constants for NH 4-N for nitrifier growth in the biofilm. Its default value is 2 mg/L NH4-N at 25°C. This is based on data presented by Hem et al. (1994) and Odegaard et al. (1994), which show that nitrification rates in biofilms at 11 to 15°C, when measured at high dissolved oxygen concentrations and low soluble COD levels, vary almost linearly from 0 to 4 mg/L NH4-N (the relationship is first-order over this range) and are fairly constant (zero order) above that. This relationship also can be represented using a Monod expression with a half-saturation constant of NH4-N of 2 mg/L. KN,DO = half-saturation constant for dissolved oxygen for nitrifier growth in the biofilm. Its default value is 4 mg/L at 25°C. The recommended temperature adjustment coefficient, θ, is 1.00 for KN,DO. Researchers also have used a first-order kinetics equation to simulate the variation in nitrification rates of the biofilm with dissolved oxygen. Equation 11.1 then would be modified as follows:
NH 4 -N u,bf,n BN,n = q m,NH4 N − Nitr,bf
Where KDO1,nitr,bf = 4.5 mg/L.
SO 2n K DO1,nitr,bf
SN n Vn Mn K N,nitr,bf + SN n
(11.1, modified)
481
482
Biofilm Reactors
Both forms of eq 11.1 replicate the variation in nitrification rates with dissolved oxygen and are consistent with the data presented by Huhtamäki (2007), Odegaard (2005b), and Weiss et al. (2005). In the data presented by Odegaard (2005b), the maximum nitrification rate in biofilms, measured at 11 to 15°C and operating at NH4-N levels above 3 mg/L and low soluble COD levels, varies linearly with dissolved oxygen from 3 to 6 mg/L. The rates increased from 0.75 kg/d·1000 m2 at 3 mg/L to 1.5 kg/d·1000 m2 at 6 mg/L dissolved oxygen, which shows a first-order relationship. The rate began to deviate gradually from first-order to half-order at dissolved oxygen levels above 6 mg/L. At a dissolved oxygen concentration of 9 mg/L, the rate was 2 kg/d·1000 m2. In the equation presented by Weiss et al. (2005), the nitrification rate in a MBBR varied linearly with dissolved oxygen over a range of 1 to 7 mg/L when the NH4-N levels were above 3 mg/L, as follows:
qm,NH4N-nit,bf = 0.214 (SO2 – 1.15) (1.047)(T-20)
2.1.1.2 Biofilm Nitrification Rates from Pilot Studies Bench-scale pilot systems must be operated to determine rate coefficients for the semi-empirical equations. The activated sludge, IFAS, and MBBR systems must be operated under identical wastewater loads, tank size and configuration, and nitrate recycle configuration, except for the biofilm support media in the aerobic zones of the IFAS and MBBR systems. An example of such a study is discussed by Sen and Randall (2005). The flowrate to the parallel IFAS/MBBR and activated sludge systems was 208 L/d, and the nominal HRT was 12 hours. Of this, 17% of the volume was anaerobic, 17% was anoxic, and the remaining 66% was aerobic (aerobic HRT of 8 hours). The operating temperature of 12°C was low enough to stress the nitrifiers in the MLSS as the MLSS MCRT was lowered. The maximum NH4-N uptake rates by nitrifiers in the biofilm (qm,NH4N-Nitr,bf) were measured for biofilm removed from the pilot systems. To measure the maximum rates of ammonium uptake, 2 L of the biofilm carrier particles and mixed liquor were removed from each of three aerobic cells of the IFAS and MBBR continuous-flow system. These were placed in flasks and spiked with ammonium chloride and bicarbonate, and aerated. The NH4-N uptake rate was measured over time. The tests were conducted at NH4-N concentrations of 50 to 10 mg/L. The rates were determined for the mixed liquor alone and for the mixed liquor with biofilm carrier particles. The tests were conducted when the systems were operated at 3.1-, 2.4-, 1.7-, 1.0-, and 0.3-day MLSS MCRTs. The combination of 5 MLSS MCRTs and three aerobic cells
Development and Application of Models
allowed one to measure rates for uptake of NH4-N by biofilm that developed under various SCODbio and NH4-N conditions. The batch tests for NH4-N uptake were operated at very low SCODbio concentrations. This limited the NH4-N uptake by the heterotrophs in the biofilm. Between each MLSS MCRT, the systems were operated for a minimum of 3 MLSS MCRTs or 8 weeks, whichever was longer, and monitored to ensure that steady-state was achieved following each change in MLSS MCRT. This was to ensure that the biofilm reached a reasonable degree of equilibrium with the new SCODbio profile. The data on qm,NH4N-Nitr,bf for the biofilm were graphed against SCODbio levels for those conditions where the SCODbio in the aerobic cell was above 10 mg/L (Figure 11.4a). Note that the concentration referenced is the concentration measured in the continuous-flow cell under steady-state operation at the various MLSS MCRTs. Above 10 mg/L, the studies showed that the nitrifiers in the biofilm were inhibited by the concentration of SCODbio and not influenced by the concentration of NH4-N in the cell. An empirical relationship (equation for the line) was developed from the data (points on the graph) using a statistical software package (SYSTAT [Chicago, Illinois]) to relate qm,NH4N-Nitr,,bf to SCODbio (Figure 11.4a), as follows: q m , NH4 N − N itr, bf =
AN K S, bfg , Ni tr K S, bfg , N itr + SCOD bio − 10
(11.3)
Where AN = 1.8 kg/1000 m2 biofilm surface/d, KS,bfg,Nitr = half-saturation constant for nitrifier growth in biofilm = 9.4 mg/L SCODbio, SCODbio = biodegradable soluble COD concentration in the mixed liquor—can be treated as equivalent to SS in the convention used by the IWA Biofilm Task Group in its presentation of models (Wanner et al., 2006). The SCODbio concentration is the soluble COD minus the non-biodegradable soluble COD. The non-biodegradable soluble COD is determined experimentally by aerating the mixed liquor or RAS in a flask for an extended period of time (24 hours) and measuring the residual soluble COD, or by operating the system at a long MLSS MCRT. For SCODbio < 10 mg/L, it was determined that the COD concentration did not inhibit the development of nitrifiers in the biofilm. The qm,NH4N-Nitr,,bf that could be
483
Biofilm Reactors
Nitrification Rate, kg/1000m2/d
(a)
3 NN4N Uptake Rate, Nitrifiers, kg/1000m2/d = 1.8 Ks / (Ks + SCODbio-10) Max NH4N Uptake Rate,Nitrifiers= 1.8 kg/1000 m2/d SCODbio = biodegradable Soluble COD Ks = 9.4 mg/L biodegradable Soluble SCOD
2.5 2 1.5 1 0.5 0
0
10
20
30
40
50
60
70
80
Soluble Biodegradable COD, mg/L (b) Nitrification Rate, kg/1000m2/d
484
NH4N Uptake Rate, Nitrif, kg/1000 m2/d = 0.42 (kg-L/mg-m2-d) x NH4N (mg/L) Maximum NH4N Uptake Rate by Nitrifiers = 1.8 kg/1000m2/d
3 2.5 2 1.5 1 0.5 0
0
1
2
3
4
5
NH4N, mg/L Note: Individual data points on the graphs are generated by operating a multicell IFAS pilot facility with three aerobic cells at different MLSS MCRTs and as an MBBR. The pilot facilities were operated in a continuous-flow mode with two anaerobic, two anoxic, and three aerobic cells and a clarifier. Media were removed from the cells and the rates were measured.
Figure 11.4 Nitrification rates for biofilm in IFAS and MBBR systems (Sen and Randall, 2008a; Sen et al., 2000): (a) biodegradable COD limiting nitrifiers in biofilm, ammonium-nitrogen is not limiting, liquid temperature = 12°C, soluble b iodegradable COD >10 mg/L, NH4-N > 3 mg/L, measured at a dissolved oxygen concentration of 8 to 9 mg/L; (b) liquid temperature = 12°C, soluble biodegradable COD < 10 mg/L, NH4-N < 5 mg/L, measured at dissolved oxygen concentration of 8 to 9 mg/L. sustained by the biofilm increased linearly with NH4-N in the continuous-flow cells (Figure 11.4b), as follows:
q m,NH4 N-Nitr,bf = (D) SN
(11.4)
Development and Application of Models
Where D = 0.47 kg-L/mg·m2 biofilm surface/d and SN = NH4-N concentration in the mixed liquor (mg/L). Sen and Randall (2007) used an alternate form of eq 11.4 to simulate conditions observed at SCODbio < 10 mg/L, as follows: q m,NH4 N− Nitr,bf =
AN SN K N,bfg − Nitr + SN
(11.5)
Where KN,bfg,Nitr = half-saturation constant term in the equation above. It is used to derive the value of qm,NH4N–Nitr,bf (maximum flux rate for NH4-N uptake by the nitrifiers) from AN. Its value is 2.1 mg/L NH4-N. Note that it is not the same as KN,bf in eq 11.1. The value of AN over the range observed is in general agreement with the values of 2 to 2.5 kg/1000 m2·d observed by Odegaard (2005a) and Odegaard et al. (1994) for Kaldnes Miljøteknologi (KMT) carriers at 10 to 15°C, low soluble biochemical oxygen demand (BOD) concentrations, and without oxygen limitation. The default value recommended for modeling is 2 kg/1000 m2·d. Note that, in an actual model construction, the coefficients of the semi-empirical equation would be modified during run time based measurements at the type of media or based on the results of the biofilm diffusional model (Section 2.2).
2.1.1.3 Ammonium-Nitrogen Uptake Rate by Nitrifiers in Mixed-Liquor Volatile Suspended Solids Equation 11.6 represents the NH 4-N uptake by nitrifiers in the suspended solids in cell n. The units for MLVSS are kilograms per cubic meter (kg/m3). The BVFn is the fraction of volume of liquid in cell n that is displaced by biofilm and its support media. It is characteristic of the type of media and media fill volume fraction (mf) in the cell. The mf is the fraction of empty tank volume that is filled with bare media. For example, if 50% of the empty tank volume is filled with bare media (media without biofilm on its surface), the mf = 0.5. The fnitr is the fraction of nitrifiers in the MLVSS. Ammonium-nitrogen uptake by nitrifiers in the suspended solids in cell n, AN,n (kg/d), is computed as shown in eq 11.6. The actual model implementation will expand the Monod function for additional parameters such as ortho-phosphorus starvation.
AN,n = q m,NH -N − Nitr,SS 4
SO 2n SN n V n(1 − BVFn ) fNitr Xn K DO,Nitr,SS+SO 2n K N,Nitr,SS + SN n
(11.6)
485
486
Biofilm Reactors
Where qm,NH4N–Nitr,SS = maximum NH4-N uptake rate for nitrifiers in the MLVSS at temperature, T; and KDO,Nitr,SS and KN,Nitr,SS = half-saturation constants for nitrifiers in the MLVSS at temperature, T; fNitr = fraction of nitrifiers in the MLVSS; and Xn = MLVSS in cell n (kg VSS/m3). The value of qm,NH4N–Nitr,SS measured by calibration of an IFAS model to the performance of the activated sludge system in the pilot studies was 8.51 mg NH4-N uptake by nitrifiers/mg nitrifier VSS·d, at 12° C (Table 11.3). This is in agreement with the value for qm,NH4N–Nitr,SS of 12.5 mg NH4-N uptake/mg nitrifier VSS·d at 25°C and a temperature adjustment coefficient of 1.03 (Randall et al., 1992; Wentzel et al., 1991).
Table 11.3 Values of coefficients measured in pilot studies and the default values for the semiempirical model (Sen and Randall, 2008a).a,b Heterotrophs in MLVSS
Value T (°C) measured
Default
Units / formula
Reference
Aerobic um,H,aer,SS
Max growth rate
12
3.57
3.57
mg VSS (mg COD/d)−1
qm,H,aer,SS
Max substrate util rate
12
8.72
8.72
mg COD (mg VSS/d)−1
YH,aer
Yield
12
0.41
0.41
mg VSS (mg COD)−1
kd,H,aer
Decay rate
12
0.042
0.042
day−1
KH,S,,aerSS
Half-saturation constant
12
48
48
mg/L SCOD
KH,DO,SS
Half-saturation constant
12
1
mg/L dissolved oxygen
Note b
qm,hydr,aer,COD,SS
Maximum hydrolysis rate
(0.1)(qm,H,aer,SS)
mg COD (mg VSS/d)–1
Note c (continued)
Table 11.3 Continued Heterotrophs in MLVSS qm,hydr,aer,PorgN,SS
qm,hydr,aer,PP,SS
Value T (°C) measured
Default
Units / formula
Reference
(qm,hydr,aer, ) (influent COD/TKN)
Maximum conversion rate (particulate OrN to Sol Org N)
COD,SS
(qm,hydr,aer, ) (influent COD/total phosphorus)
Maximum conversion rate (particulate OrgP to Sol P, not OP)
COD,SS
qm,hydr,aer,SorgN,SS
Maximum ammonification rate (soluble OrgN to NH4-N)
(0.8) (qm,hydr,aer,PorgN,SS)
qm,hydr,aer,SP,SS
Maximum ammonification rate (soluble P to OP)
(0.8) (qm,hydr,aer,PP,SS)
Anoxic um,H,anx,SS
Maximum growth rate
12
1.77
1.77
mg VSS (mg COD/d)–1
For pre-anx
qm,H,anx,SS
Maximum substrate 12 utilization rate
5.71
5.71
mg COD (mg VSS/d)–1
For pre-anx
YH,anx
Yield
12
0.31
0.31
mg VSS (mg COD)–1
kd,H,anx
Decay rate
12
0.022
0.022
day–1
KH,S,anx,SS
Half-saturation constant
12
56
48
mg/L SCODbio
KH,DO,I,SS
Half-inhibition constant
12
0.25
mg/L dissolved oxygen
b
KH,NO3N,SS
Half-saturation constant
12
1.0
mg/L NO3-N
b
(Aerobic rate) (qm,H,anx,SS/ qm,H,aer,SS)
c
Hydrolysis rate
(continued)
487
Table 11.3 Continued Heterotrophs in MLVSS
Value T (°C) measured
Default
Units / formula
12
0.005
day–1
Reference
Anaerobic kd,Hana
Decay rate
0.005
Hydrolysis rate
(0.5)(anoxic rate) c (0.5)
Autotrophs in MLVSS Nitrosomonas or Ammonia Oxidizing Bacteria
Values from literature calibrated to pilot studies
um,NH4N-Nitr,SS
Maximum growth rate
12
0.43
0.43
mg VSSNitr1 (mg NH4-N/d)–1
qm,NH4N-Nitr,SS
Maximum NH4N util rate
12
8.51
8.51
mg NH4N (mg VSSNitr1/d)–1
YN1
Yield
12
0.05
0.05
mg VSSNitr1 (mg NH4-N)
b
kd,N1
Aerobic decay rate
12
0.023
0.023
day–1
b
KDO,nitr, SS
Half-saturation 12 constant for nitrifiers
1
mg/L dissolved oxygen
b
KN,nitr,SS
Half-saturation 12 constant for nitrifiers
1
mg/L NH4-N
b
d
Biofilm in semi-empirical model These are concentrations in the bulk liquid, not inside biofilm KS,H,aer,bf
Half-saturation, 12 aerobic, heterotrophs
48
48
mg/L SCODbio
KS,H,anx,bf
Half-saturation, 12 anoxic, heterotrophs
56
48
mg/L SCODbio
KDO,H,aer,bf
Half-saturation, aerobic, heterotrophs
12
4.0
KDO,H,i,bf
Half-inhibition constant, anoxic
12
ASaer
Max aer COD 12 uptake rate, biofilm
mg/L dissolved oxygen
21
2.0
mg/L dissolved oxygen
21
kg COD/1000 m2·d
e
f (continued)
488
Table 11.3 Continued Heterotrophs in MLVSS
Value T (°C) measured
Default
Units / formula
Reference
ASanx
Max anx COD 12 uptake rate, biofilm
14
14
kg COD/ 1000m2·d
BSaer
Half-saturation, for AS
12
19.3
19.3
mg/L SCODbi o
BSanx
Half-saturation, for AS
12
19.3
19.3
mg/L SCODbio
KDO, nitr, bf
Half-saturation, aerobic, nitrifiers
12
4.0
mg/L dissolved oxygen
e
KN, nitr,bf
Half-saturation aerobic, nitrifiers
12
2.0
mg/L NH4-N
e
AN
Max NH4-N uptake 12 rate (nitr rate)
1.8
1.8
kg/1000 m2·d
KS,bfg,nitr
Half-inhibition, AN for nitrif
12
9.4
9.4
mg/L SCODbio
KN,bfg,nitr
Half-saturation, AN for nitrif
12
2.1
2.1
mg/L NH4-N
Aerobic COD hydrolysis rates by biofilm
20% of qmaer,COD,bf
kg COD/1000 m2·d
fN
Nitrogen in biomass
0.12
mg N/mg VSS
fCOD
COD in biomass
1.42
mg COD/mg VSS
g
General note: Mmdel is structured such that the modeler can use observed values instead of the default values for kinetic coefficients and Ө when observed values are available. a Unless mentioned, these values were determined from continuous-flow bench-scale pilot studies operated in activated sludge, IFAS, and MBBR modes (Sen, 1995). b Literature referenced: Dold, 1991; Marais and Ekama, 1976; Randall et al., 1992; Wentzel et al., 1991. c Maximum hydrolysis rate should be refined by calibrating model results to operating data. d The value of 0.43 day–1 was determined by calibrating the model to the performance of the bench-scale activated-sludge system operated at different MLSS MCRTs. e Hem et al. (1994); Huhtamäki (2007); Odegaard (2005b); Weiss et al. (2005). f Based on ratio of max substrate utilization rates under pre-anoxic and aerobic conditions. The value must be corrected for post-anoxic cells. g For aerobic hydrolysis by biofilm, substitute qmaer,COD,bf for qm,H,aer,SS in the equation for MLVSS; all other hydrolysis equations for the biofilm are related to this equation based on a form similar to the equations for hydrolysis with MLVSS.
489
490
Biofilm Reactors
The default value for KDO,nitr,SS is 1 mg/L dissolved oxygen at 25°C. Its temperature adjustment coefficient, θ, is 1.00. The default value for KN,nitr,SS is 1 mg/L NH4-N at 25°C. Its temperature adjustment coefficient, θ, is 1.06. This results in a value of 0.46 mg/L at 12°C. These values of the coefficients are based on the work on nitrification and denitrification in activated sludge systems by Marais and Ekama (1976) and Wentzel et al. (1991). For the purposes of design, one may adjust or inhibit the value of qm,NH4N−Nitr,SS (8.51 day−1) by a certain percentage to account for nitrifier inhibition experienced at a full-scale facility. A typical percentage for the inhibition recommended for design is 25%. The inhibition is because of the presence of certain chemicals in the wastewater, such as in a system that receives a combination of sewage and septage. Certain facilities have significantly higher inhibition, such as inhibition by cyanide recirculated from multiple hearth incinerators burning sludge. This has been observed at the Western Branch, Maryland, and the Vip, Norfolk, Virginia, plants, when operated with recycle flow from multiple hearth furnace exhaust gas scrubbers. The inhibition was as high as 50% (Solley, 2000). At other facilities, such as Broomfield, Colorado (discussed below), continuous simulation against extended periods of plant data showed no significant inhibition.
2.1.1.4 Mass Balance for Ammonium-Nitrogen in Each Reactor The NH 4-N concentration (mg/L) in cell n is calculated using a mass-balance approach, as follows: I N,n + N decay,n + N org-N,hydr,n = AN,n + BN,n + CN,n + DN,n + EN,n
(11.7)
Where IN,n = quantity of un-assimilated NH4-N entering aerobic cell n (kg/d); EN,n = quantity of un-assimilated NH4-N exiting cell n (kg/d); AN,n and BN,n = N H 4-N uptake by nitrifiers in suspended solids and biofilm, computed in eqs 11.6 and 11.1, respectively (kg/d); CN,n and DN,n = NH4-N utilization in cell n resulting from heterotrophic biomass production by MLVSS and biofilm, respectively (kg/d); Ndecay,n = NH4-N released through decay of VSS in cell n (kg/d); and Norg-N,hydr,n = hydrolysis of organic nitrogen (kg/d). CN,n is the NH4-N uptake by biomass in the MLSS using dissolved oxygen and NOx-N in cell n, as follows: CN,n (kg/d) = {[CODu, aer, SS,n][YAH,aer,SS] + [CODu, anx, SS,n][YAH,anxSS]} fN
(11.8)
Development and Application of Models
Where fN = f raction of nitrogen in the biomass (MLVSS) synthesized (default value of 0.12); and CODu,aer,SS,n and CODu,anx,SS,n = COD used by MLVSS under aerobic and anoxic conditions, respectively, in cell n. The variables Y AH,SS and Y AH,BF (eq 11.9) are the actual biomass yield for heterotrophs in the suspended solids and the biofilm. Their computation is discussed in the Biomass Production section. DN,n (kg/d) is the NH4-N uptake by heterotrophic yield of the biofilm in cell n, as follows:
DN,n = [CODu, bf, n][YAH,BF]} fN
(11.9)
Unlike NH4-N uptake rates by the MLVSS, AN,n and DN,n are the measured net of decay of autotrophic and heterotrophic biomass in the biofilms, respectively. Ndecay,n is the nitrogen released from biomass in MLVSS as a result of decay. When the dissolved oxygen is above the threshold that represents aerobic conditions, the form of the equation is as follows:
Ndecay,n (kg/d) = (fN) (kdH,aerT Xn Vn) (1 – BVFn)
(11.10)
When the dissolved oxygen in a particular cell drops below the threshold that represents aerobic conditions but is above the NOx-N threshold that represents anoxic conditions, the decay rate in eq 11.10 is switched from aerobic decay rate of kdH,aerT to the anoxic decay rate of kdH,anxT. If it is below the NOx-N threshold, the decay rate is switched to the anaerobic decay rate of kdH,anaT (Table 11.3). It should be noted that the Ndecay,n is not included if the value of the NH 4-N utilization rate term by the MLVSS for VSS production (Cn) is the measured net of biomass decay. The corresponding term for biofilm is not included, because the NH4-N uptake rate in eq 11.1 was the measured net of biofilm decay. Norg-N,hydr,n is the amount of un-assimilated organic nitrogen that is hydrolyzed in cell n. It is computed as follows (kg/d):
Norg-N,hydr,n = (IPorgN + ISorgN – EPorgN – ESorgN)n
(11.11)
The rate of hydrolysis can be computed based on a series of Monod equations. For the MLVSS, the maximum hydrolysis rate for un-assimilated particulate organic nitrogen (qm,hydr,EA,PorgN,SS) is expressed in units of kilograms of particulate organic nitrogen hydrolyzed per kilogram MLVSS per day; the half-saturation constant (Khydr,EA,PorgN,SS)
491
492
Biofilm Reactors
is expressed in milligrams per liter of particulate organic nitrogen. The subscript “EA” represents aerobic, anoxic, or anaerobic conditions for the electron acceptor (EA), based on the thresholds of dissolved oxygen and NOx-N , which represent aerobic and anoxic conditions, respectively. For the biofilm, the maximum hydrolysis rate for particulate organic nitrogen (qm,hydr,EA,PorgN,bf) is expressed in units of kilograms of particulate organic nitrogen hydrolyzed per day per 1000 m2 of biofilm surface. The half-saturation constant (Khydr,EA,PorgN,bf) is expressed in milligrams per liter of particulate organic nitrogen. The effluent particulate organic nitrogen (EPorgN) is computed from the influent particulate organic nitrogen (IPorgN) and the hydrolysis of particulate organic nitrogen by the MLVSS and the biofilm. This is simulated using eqs 11.12 to 11.15.
IPorgN,n = EPorgN,n + Porg Nhydr,SS,n + Porg Nhydr,bf,n Porg N hydr,SS,n = q m,hydr,EA,PorgN,SS
SPorgN,n K hydr,EA,PorgN, SS+SPorgN,n
Porg N hydr,bf,n = q m,hydr,EA,PorgN,bf
(11.12)
V n(1 − BVFn ) Xn
SPorgN,n K hydr,EA,PorgN, bf +SPorgN,n
(11.13) V n Mn
EPorgN = (SPorgN )(Qeff,n)
(11.14) (11.15)
Where SPorgN = concentration of particulate organic nitrogen in the effluent from cell n and Qeff,n = effluent flow from cell n. The particulate organic nitrogen gets hydrolyzed to soluble organic nitrogen. The soluble organic nitrogen then is converted (de-amminated) to NH4-N. The effluent soluble organic nitrogen (ESorgN) in cell n is computed as follows (kg/d): ISorgN + Porg Nhydr,SS,n + Porg Nhydr,bf,n = ESorgN + Sorg Nhydr,SS,n + Sorg Nhydr,bf,n
(11.16)
The equations to determine soluble organic nitrogen hydrolyzed by the biofilm and suspended solids (Sorg Nhydr,SS,n + Sorg Nhydr,bf,n) are structured in a format similar to eqs 11.12 to 11.15. In the subscript, SorgN is substituted for PorgN. The values of maximum hydrolysis rate for soluble organic nitrogen for MLVSS and biofilm (qm,hydr,SorgN,SS and qm,hydr,SorgN,bf) and the corresponding half-saturation constants for soluble organic nitrogen are used in eqs 11.13 and 11.14.
Table 11.4 Matrix showing kinetic coefficients for MLVSS in semi-empirical and biofilm models (X = MLVSS; AOB = ammonia-oxidizing bacteria; NOB = nitrite-oxidizing bacteria) (Sen and Randall, 2008a).* XH(X [1-fnitr]) Heterotrophs using dissolved oxygen
XN SS SN SO2 (X fnitr) (SCOD- (NH4-N) (DO) bio) −1 YHaer
1
− fN
SNO3N (NO3-N)
− (1 − fCODYHaer) YHaer
Kinetic expression for cell n (n = cell number in multicell reactor)
YHaer [q m,H,aer,SS
SO 2 n K H,DO,SS + SO 2 n
SCOD bio,n K H,S,aer,SS + SCOD bio,n V n(1 − BVFn ) (1 − fNitr ) Xn ]
493
Heterotrophs using NO3-N
−1 YHanx
1
− fN
SNOxN n − (1 − fCODYHanx) YHanx [qm,H,anx,SS 2.86 YHanx K H,NOx-N,SS+SNOxN n KH,DO,i,SS
SCOD bio,n
K H,S,anx,SS + SCOD bio,n K H,DO,i,SS + SO2 n Vn (1 − BVFn ) (1- fNitr )Xn ] AOB (N1) using dissolved oxygen
1
−1 YN1
−3.43(1 − fNYN1) YN1
YN1 [qm,NH4-N − Nitr,SS
SO 2 n K DO,nitr,SS+SO 2 n
SN n V n(1 − BVFn ) fNitr Xn ] K N,nitr,SS + SN n NOB (N2) using dissolved oxygen
1
− fN
−1.14 YN2
AOB+NOB (if YN1=YN2 & eff NO2-N=0
2
−1 − fN YN1
−4.57(1 − fNYN1) −1 YN1 YN1
−1 YN2 2 YN1 [qm,NH4 -N − Nitr,SS
SO 2 n K DO,nitr,SS+SO 2 n
SN n V n(1 − BVFn ) fNitr Xn ] K N,nitr,SS + SN n
*Note: except for BVF = 0, the same kinetic expressions also apply to each layer inside biofilm in the biofilm diffusional model. For each layer, n = bn.
494
Biofilm Reactors
A set of values of coefficients have been developed for use in the model (Table 11.3). These values are derived by calibrating the model to the data from the continuous-flow pilot studies (Sriwiryarat et al., 2005) and calibration of the model to several full-scale MBBRs and IFAS plants (Sen et al., 2006). In eqs 11.13 and 11.14, the half-saturation constants Khydr,EA,PorgN,SS and Khydr,EA,PorgN,SS are 1 mg/L for all electron acceptor conditions. It should be noted that the value of the hydrolysis rate selected has greater effect in winter compared with summer and in an MBBR compared with an IFAS. This is because of the slower hydrolysis rate in winter and the relatively low MLVSS in the MBBR, which requires most of the hydrolysis to be in the biofilm. The values computed from the pilot studies were revised based on comparison with MBBRs running in winter. Particular care should be taken in running models for MBBRs at low temperatures (2 to 12°C). The results should be checked against actual plants and the values of the coefficients adjusted if necessary. Particular care should be taken in identifying the correct hydrolysis rates for the biofilm (both for the semi-empirical version and the one- or two-dimensional version). Discrepancies in the hydrolysis rate can result in a significant difference in effluent organic nitrogen in cold weather between an actual plant data and the model data when operating as an MBBR. The assumption made in certain models that all of the hydrolysis occurs through enmeshment in the biological floc of MLVSS, while defensible in an IFAS and activated sludge, is not valid, because it results in a substantially higher value of effluent particulate nitrogen in MBBRs than observed in actual plants. The equations for nitrification, denitrification, and COD removal are summarized in a matrix format in Tables 11.4 (MLVSS) and 11.5 (biofilm).
2.1.2 Chemical Oxygen Demand Removal The COD removal rates are computed by a set of equations similar to those for nitrification. The COD can be removed aerobically and anoxically by the biofilm and the MLVSS in each cell. Equation 11.17 shows the aerobic COD uptake (CODu,aer,bf = Bn,1,S) by the biofilm.
COD u,aer,bf , Bn,1,S = q m,aer,COD,bf
SO 2 n
SCOD bio,n V n Mn (11.17) S O + K DO,H,aer,bf K S,H,aer,bf + SCOD bio,n 2n
Table 11.5 Semi-empirical modeling of the biofilm—equations for semi-empirical computation of biofilm flux (Sen and Randall, 2008a). XH(X [1-fnitr]) Heterotrophs using dissolved oxygen
1
Heterotrophs using NO3-N
1
XN SS (X fnitr) (SCODbio) −1 YHaer
SN (NH4-N)
SO2 (DO)
− fN
− (1 − fCODYHaer) YHaer
SNO3N (NO3-N)
Kinetic expression for cell n (n = cell number in multicell reactor)
[
YHaer qm,aer,COD,bf
SO 2 n K DO,H,aer,bf +SO 2 n
SCOD bio,n V n Mn K S,H,aer,bf + SCOD bio,n −1 YHanx
− fN
− (1 − fCODYHanx) 2.86 YHanx
[
YHanx qm,anx,COD,bf
]
K H,DO,i,bf K H,DO,i,bf +SO 2 n
495
SCOD bio,n K S,H,anx,bf + SCOD bio,n SNOxN n Vn Mn K NOx-N,bf + SNOxN n −1 YN1
−3.43(1 − fNYN1) YN1
AOB (N1) using dissolved oxygen
1
NOB (N2) using dissolved oxygen
1
− fN
−1.14 YN2
AOB+NOB (if YN1=YN2 & eff NO2-N=0
2
−1 − fN YN1
−4.57(1 − fNYN1) −1 YN1 YN1
YN1
[q
m,NH 4 -N − Nitr,bf
]
SO 2 n K DO,nitr,bf +SO 2 n
SN n Vn Mn K N,nitr,bf + SN n
]
−1 YN2
[
2 YN1 qm,NH4 -N − Nitr,bf
SO 2n K DO,nitr,bf +SO 2n
SN n Vn Mn K N,nitr,bf + SN n
]
Biofilm Reactors
Maximum COD removal flux rates (qm,aer,COD,bf) also increased with the SCODbio concentration. As mentioned earlier, SCODbio can be treated as equivalent to SS in the convention used by the IWA Biofilm Task Group in its presentation of models. Based on batch tests to measure rates on biofilm removed from the continuous-flow cells, an empirical relationship was developed to relate qm,aer,COD,bf to SCODbio (Figure 11.5), as follows: q m,aer,COD,bf =
AS,aer SCOD bio BS,aer + SCOD bio
(11.18)
Where AS,aer = 21 kg/1000 m2 of biofilm surface/d and BS,aer = 19.3 mg/L SCODbio. Equation 11.17 shows that the rate changes with biodegradable soluble COD (SCODbio,n) and dissolved oxygen (SO2,n ) in cell n. The default value of half-saturation constant for COD, KS,H,aer,,bf, is 48 mg/L COD at 12°C. This is equal to the value of KSHaer,SS measured for biomass in MLSS in the pilot studies (Table 11.3). This value increases to 70 mg/L COD at 25°C. The temperature adjustment coefficient, θ, in the Arrhenius equation, is 1.03. COD Uptake Rate, kg/1000m2/d = 21 SCOD_bio / (Ks + SCOD_bio) SCOD_bio = biodegradable Soluble COD Ks = 19.3 mg/L biodegradable Soluble SCOD
25 COD Uptake Rate, kg/1000 m2/d
496
20 15 10 5 0
NH4-N above 3 mg/L 0
20
40
60
80
100
SCOD_bio for biofilm development, mg/L Note: Individual data points on the graphs are generated by operating a multicell IFAS pilot facility with three aerobic cells at different MLSS MCRTs and as an MBBR. The pilot facilities were operated in a continuous-flow mode with two anaerobic, two anoxic, and three aerobic cells and a clarifier. Media were removed from the cells and the rates were measured.
Figure 11.5 COD uptake rates for biofilm in IFAS and MBBR systems (Sen and Randall, 2008a; Sen et al., 2000): liquid temperature = 12°C, measured at a dissolved oxygen concentration of 8 to 9 mg/L.
Development and Application of Models
The default values of half-saturation constant for dissolved oxygen, KDO,H,aer,bf , at 25°C, is 4 mg/L dissolved oxygen; the θ for KDO,H,bf,aer is 1.00. Equation 11.19 shows the anoxic COD uptake (COD u,anx,bf = Bn,2,S) by the biofilm. COD u,anx,bf
B2,S,n = q m,anx,COD,bf
KH,DO,i,bf
K H,DO,i,bf + SO2 n SNOxN n Vn Mn K NOx-N,bf + SNOxN n q m,anx,COD,bf =
SCOD bio,n K S,H,anx,bf + SCOD bio,n
AS,anx SCOD bio BS,anx + SCOD bio
(11.19)
(11.20)
Where the default values for A and B are as follows: AS,anx = 13.8 kg/1000 m2 of biofilm surface/d, if the media is in the pre-anoxic zone; and BS,anx = 19.3 mg/L SCODbio. The value for AS,anx is based on the ratio of maximum substrate utilization rates for heterotrophs in the MLVSS under pre-anoxic and aerobic conditions (qmHanx SS/qmHaer SS). The default value recommended for the half-saturation constant for dissolved oxygen inhibition, KH,DO,i,bf, at 12°C, is 2 mg/L dissolved oxygen (Table 11.3). The default value of the half-saturation constant for dissolved oxygen inhibition, KH,DO,i,bf, at 25°C, is 2 mg/L dissolved oxygen. Its temperature adjustment coefficient, θ, is 1.00. The default value for COD, KS,H,bf,anx, at 12°C, is 48 mg/L COD, which is the same as that recommended for aerobic conditions. Measurements in the pilot units indicated that the actual value may be slightly higher (Table 11.3, 56 mg/L). Values of 45 to 70 mg/L show reasonably good simulation of data from full-scale IFAS plants. The temperature adjustment coefficient, θ, for KS,H,bf,anx is 1.03. The default value half-saturation constant for NOx-N , K NOx-N ,bf , at 25°C, is 1 mg/L NOx-N. Equations 11.18 and 11.20 also can be used to compute the NH4-N uptake by the biofilm for cell synthesis and decay by multiplying the terms by fN (fraction nitrogen in the biomass). Equation 11.19 can be split into two separate equations based on COD uptake with NO2-N and NO3-N as the two separate forms of NOx-N. Also, it is important to recognize that COD can be taken up anoxically and aerobically in the biofilm, if located in an aerobic cell.
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498
Biofilm Reactors
Data from full-scale plants show that, even in conventional activated sludge, simultaneous aerobic and anoxic uptake of COD is possible in each cell. This forms the basis of eqs 11.21 and 11.22, which are used to calculate the aerobic (An,1,S) and anoxic (An,2,S) uptake of COD by the suspended solids in cell n. COD u,aer,ss A1,S,n = qm,H,aer,SS
SCOD bio,n SO 2 n V n(1 − BVFn ) (1 − fNitr ) Xn (11.21) O + S K H,S,aer,SS + SCOD bio,n K H,DO,SS 2n
KH,DO,i,SS SCOD bio,n SNO x N n K H,NOx N,SS+SNO x N n K H,S,anx,SS + SCOD bio,n K H,DO,i,SS + SO 2 n Vn (1 − BVFn ) (1- fNitr )Xn
COD uanx,ss A2,S,n = qm,H,anx,SS
(11.22)
Where qm,H,aer,SS and qm,H,anx,SS = maximum COD uptake rates for heterotrophs in the MLVSS at temperature, T; KH,DO,SS and KH,S,aer,SS = h alf-saturation constants for dissolved oxygen and COD for the heterotrophs in the MLVSS during aerobic uptake of COD at temperature, T; KH,NOx-N ,SS and KH,S,anx,SS = half-saturation constants for NOx-N and COD for the heterotrophs in the MLVSS during anoxic uptake of COD at temperature, T; and KH,DO,i,SS = half inhibition constant for denitrification, expressed as mg/L dissolved oxygen. The value for q m,H,aer,SS measured at 12°C was 8.72 mg COD taken up/mg heterotrophic VSS/d (Table 11.3). This is in agreement with the value 12.8 mg COD taken up/mg heterotrophic VSS/d at 25°C and a temperature adjustment coefficient of 1.03 (Dold, 1991; Randall et al., 1992, Wentzel et al., 1991). Values of θ for heterotrophs in the literature vary from less than 1.03 to 1.07 (Marais and Ekama, 1976; Wentzel et al., 1991). The value of KH,S,aer,SS measured in pilot studies was 48 mg/L (Table 11.3). This agrees with the KH,S,aer,SS value of 70 mg/L COD at 25°C observed by McClintock et al. (1988) and a temperature adjustment coefficient of 1.03. The default value for KH,DO,SS is 1 mg/L dissolved oxygen at 25°C. The value for qm,H,anx,SS measured at 12°C was 5.71 mg COD taken up/mg heterotrophic VSS (Table 11.3). This is for biomass that grows using a primary effluent with 25 to 33% of the COD available as SCODbio. A separate value of qm,H,anx,SS should
Development and Application of Models
be determined for post-anoxic cells with supplemental carbon. This value may be lower when using methanol as a substrate (Dold et al., 2007). The value of KH,S,anx,SS measured under anoxic conditions at 12°C was 56 mg/L (Table 11.3). This is slightly higher than the value of 48 mg/L under aerobic conditions. As a default, one may use the same value for modeling anoxic and aerobic zones. The default values for KH,NO3N,SS and KH,DO,i,SS are from the literature (Table 11.3). The value of KH,NO3N,SS is 1 mg/L NO3-N, and the value of KH,DO,i,SS is 0.25 mg/L dissolved oxygen. The SCODbio concentration (mg/L) in each aerobic cell n is calculated using a mass-balance approach similar to eq 11.7. The units of the terms in eq 11.23 are in kilograms per day (kg/d).
I n,S + S decay,n+ Shydr,n = An1,S + An2,S + Bn1,S +Bn2,S + En,S
(11.23)
Where In,S and En,s = influent and effluent of SCODbio for cell n (kg/d) and Sdecay,n = COD released through decay of MLVSS. A corresponding term for COD release through decay of the biofilm is not included, because the COD flux in eqs 11.18 and 11.20 are the net of decay. The COD flux observed in the one- and two-dimensional biofilm-diffusion model discussed below is also the net of decay.
Sdecay, n (kg/d) = (fCOD) (kdH,aerT Xn Vn) (1 – BVFn)
(11.24)
Where fCOD = COD content of biomass (with a default value of 1.42 mg COD per mg VSS).
2.1.3 Biomass Production The amount of heterotrophs and nitrifiers generated as a result of COD removal and nitrification is computed for the MLSS and biofilm. This computation is performed for each cell n. The sum total for all n cells is the total biomass production. The user specifies one or more MLSS MCRTs for the computation. In the advanced models, such as Aquifas, BioWin, and GPS-X, the model uses information on the plant configuration (e.g., anaerobic, anoxic, aerobic, and post-anoxic volume fractions; step-feed, modified Ludzack–Ettinger [MLE], University of Cape Town [Cape Town, South Africa] [UCT], or aerobic–anoxic–oxic [A2O] configurations) to compute the MLSS MCRT under aerobic, anoxic, and anaerobic conditions. Additionally, the IFAS and MBBR models have to compute or use a specified value for biofilm yield
499
500
Biofilm Reactors
in each aerobic cell. Unlike a biofilm-diffusion model (one- and two-dimensional models), the semi-empirical model has no means of computing the biofilm yield in each cell. Instead, it uses a reference table, in which the heterotrophic and autotrophic yields are specified for different conditions (anoxic or aerobic), bulk liquid COD, and NH4-N levels. This table must be generated from pilot studies or from multiple runs with a calibrated one- or two-dimensional biofilm model.
2.1.3.1 Mixed-Liquor Volatile Suspended Solids Computation of biomass production, MLVSS, and fraction nitrifiers in activated sludge and MBBR systems are simpler than in an IFAS system. The IFAS system must consider the interaction between the biofilm and the mixed liquor, with significant removals and biomass generation in both. This is done using an iterative multistep approach. The biomass yield for heterotrophs (kg/d) in the MLVSS is computed in four steps. (1) The first step computes the yield in cell n. Heterotrophic biomass yield by MLVSS in cell n:
AXHy,n = (CODu,aer,SS,n) (YHaer) + (CODu,anx,SS,n) (YHanx)
(11.25)
The CODu,aer,SS,n and CODu,anx,SS,n are computed in eqs 11.21 and 11.22. (2) The second step computes the decay of MLVSS biomass in reactor cell n (kg/d). As discussed earlier for eq 11.10, the decay rate, kdH,EA,n, in cell n is a function of dissolved oxygen and NOx-N in the cell. Biomass decay for heterotrophs in cell n:
An,XHd = −(kdH,EA,n) (MLVSS)(1 – fnitr)(Vn)(1 – BVF)
(11.26)
The overall heterotrophic biomass production in reactor cell n is the sum of biomass yield and decay, as computed in eqs 11.25 and 11.26. The value of fnitr must be computed iteratively by running through the entire set of eqs 11.1 to 11.44. (3) The nitrifier yield is computed in the third step using an equation that is similar to eq 11.25 for heterotrophs, as follows: Nitrifier biomass yield by MLVSS in cell n, An,XNy = (NH4-Nu,SS,n) (YN1) + (NO2-Nu,SS,n) (YN2) (11.27) The first term on the right side is for Nitrosomonas biomass yield = (NH4-Nu,SS,n) (YN1). The second term on the right side is for Nitrobacter biomass yield = (NO2-Nu,SS,n)(YN2). (4) The nitrifier decay is computed in the fourth step by eqs 11.28, 11.29, and 11.30. Equations 11.28 and 11.29 are similar to eq 11.26 for heterotrophs.
Development and Application of Models
Biomass decay for Nitrosomonas (N1) in cell n, An,XN1d = −(kdN1,EA,n) (fnitr1) (MLVSS) (Vn) (1 – BVF) (11.28) Biomass decay for Nitrobacter (N2) in cell n, An,XN2d = −(kdN1,EA,n) (fnitr2) (MLVSS) (Vn)(1 – BVF) (11.29) An,XNd = An,XN1d + An,XN1d
(11.30)
2.1.3.2 Biofilm The structure is similar to MLVSS. The heterotrophic biomass generated by biofilm in cell n, BXH,n, is computed by eq 11.31, as follows: Heterotrophic biomass generated by biofilm in cell n:
BXH,n = [CODu, bf,n][YH,bf,n]
(11.31)
This biomass is the quantity sloughed off the biofilm and released into the MLVSS. The COD used by the biofilm in reactor cell n is computed by eqs 11.17 (CODu,bf,aer) and 11.19 (CODu,bf,anx). For the semi-empirical model, the biofilm yield in cell n, YH,bf,n (flux of heterotroph out of the biofilm for each unit flux of COD into the biofilm per day), must be specified as an external input. Its value can be determined by running the mechanistic biofilm one-dimensional model in conjunction with the semi-empirical model. Alternatively, it may be based on measurements made during pilot studies, from a data table included as part of the semi-empirical model, or on data from the manufacturer. The data table is a table of yields observed at different SCOD bio and NH4-N concentrations. It is based on the results of several runs made with the biofilm one-dimensional model (Table 11.1). For nitrifiers, the computations are similar to heterotrophs. Equation 11.31 is modified for nitrifiers, as follows: Nitrifier biomass generated by biofilm in cell n:
BXN,n = [NH4-Nu, bf,n][YN1,bf,n] + [NO2-Nu, bf,n][YN2,bf,n]
(11.32)
Where
BXN,n = amount of nitrifiers sloughed off the biofilm and released into the mixed liquor; NH4-Nu, bf,n = computed by eq 11.1; YN1,bf,n = Nitrosomonas flux (as VSS out of biofilm) per unit flux of NH4-N into the biofilm per day; and YN1,bf,n = Nitrosomonas flux (as VSS out of biofilm) per unit flux of NH4-N into the biofilm per day.
501
502
Biofilm Reactors
The nitrifier biomass generated also can be based on a lumped yield of all nitrifying (ammonia-oxidizing and nitrite-oxidizing) bacteria (Tables 11.4 and 11.5). Nitrifier biomass generated by biofilm in reactor cell n:
BXN,n = [NH4-Nu, bf,n][YNtotal,bf,n]
(11.33)
Where YNtotal,bf,n = ammonia-oxidizing bacteria flux (as VSS out of biofilm) per unit flux of NH4-N into the biofilm per day, and YNtotal,bf,n = twice the values of YN1,bf,n in Table 11.1. Unlike heterotrophs, the effect of nitrifier decay rate within the biofilm is incorporated to the biofilm yields (i.e., the biofilm yields are the net of decay).
2.1.4 Fraction of Nitrifiers The fraction of nitrifiers in the MLVSS can be computed in the following three steps: (1) Quantify nitrifiers generated in the MLVSS and nitrifiers sloughed off the biofilm in each cell using the equations above. (2) Determine the sum total of nitrifiers and heterotrophs generated and lost through decay across n cells in the system; this is the biomass production per day. (3) Compute the fraction of nitrifiers based on eq 11.34. The fraction of nitrifiers (fnitr) can be computed as follows: n
fnitr =
∑ Nitrifier Biomass Yield and Decay for MLVSS and Biofilm 1
n
∑ Heterotroph and Nitrifier Biomass Yields and Decay for MLVSS and Biofilm 1
(11.34) Where n = number of reactor cells in operation. Biomass yield and decay are computed as shown in eqs 11.27 to 11.33. As mentioned earlier, the value of fnitr calculated during each iteration is fed back into the next model run. It is important to understand that, unlike activated sludge systems operating below the threshold (washout) MLSS MCRT, where the fnitr can be close to zero because of washout of the nitrifiers, the IFAS (and MBBR) system can have a significant fnitr, even when the MLSS MCRTs are below the washout MCRT of single-cell activated sludge systems. This is because of biofilm nitrification and nitrifiers sloughed off the biofilm. These nitrifiers become part of the MLVSS until it is wasted from the system or released in the effluent.
Development and Application of Models
Secondly, it is important to appreciate how the fnitr plays slightly different roles in IFAS and MBBR systems. In IFAS systems, the nitrifier population in the MLVSS can make a significant contribution to the overall nitrification, even when operated well below the washout MLSS MCRT. However, in the MBBR, which has a very low MLVSS, and MLSS MCRT = nominal HRT of the liquid, the MLVSS does not maintain a significant population of nitrifiers. It is for this reason that IFAS systems show a lower activated sludge tank volume requirement compared with MBBRs when designed with the same biofilm surface area. However, IFAS systems need sludge volume index control and, possibly, higher clarifier surface areas.
2.1.5 Denitrification The NOx-N denitrified by the biofilm (NOx-N u,anx,bf,n = BNOx-N ,u,n) and the MLVSS concentration (NOx-N u,anx,SS,n = ANOx-N ,u,n) are computed as follows:
NO x -N u,anx,bf,n ,
BNOx N,u,n =
NO x N u,anx,bf,n ,
ANOx N,u,n =
COD u,anx,bf,n DNCOD Factor
COD u,anx,SS,n DNCOD Factor
(11.35)
(11.36)
The CODu,anx,bf,n (B2,S,n) and CODu,anx,SS,n (A2,S,n) are computed as per eqs 11.19 and 11.22, respectively. As mentioned earlier, COD u,anx, bf, n and CODu,anx, SS, n can each be split into two terms to separately compute the COD used for denitrification of NO2-N and NO3-N. The DN COD factor is computed from the stoichiometry of denitrification, yield, and COD content of biomass (fCOD), as follows:
DN COD factor NO3N = 2.86 / (1 – Yh,anx * fCOD)
(11.37)
DN COD factor NO2N = 1.71 / (1 – Yh,anx * fCOD)
(11.38)
As mentioned earlier, the default value of fCOD is 1.42 mg COD/mg VSS. This method of computation of NO x-N uptake by the biofilm has some challenges. The user has to specify the half-saturation and inhibition constants in eq 11.19. Equation 11.20 is developed for anoxic conditions. Because COD may be taken up aerobically by the biofilm in the aerobic cells and denitrified in layers inside the biofilm, there can be some error in the estimation of denitrification in the aerobic zone. To overcome this, the user must calibrate the half-saturation constants and inhibition constants to a full- or pilot-scale facility.
503
504
Biofilm Reactors
The effluent NOx-N (ENOx-N ,n) load from reactor n (kg/d) is determined based on the influent NOx-N load (INOx-N ,,n), nitrification and denitrification in the MLVSS, and biofilm, as follows: I NOx N,n + B NOx N-Nitr,n+ ANOx N-Nitr,n = ANOx N,u,n + BNOx N,u,n + ENOx N,n
(11.39)
2.1.6 Oxygen The equations for oxygen requirement for the biofilm (BDO,n) and the MLVSS (ADO,n) are computed as follows:
BDO,n = B1,S,n + B2,S,n – fCOD BXH,n + 4.57 BNOx-N-Nitr,n – 2.86 BNOx-N,u,n
(11.40)
ADO,n = A1,S,n + A1,S,n – fCOD AXH,n + 4.57 ANOx-N-Nitr,n – 2.86 ANOx-N,u,n
(11.41)
A mass balance is conducted to determine the oxygen required (kg/d) in cell n, as follows:
I DO,n + D DO,n + TDO,n = ADO,n + BDO,n + EDO,n
(11.42)
Where, for cell n, DDO,n = dissolved oxygen diffusing in from the atmosphere that is not associated with an aeration device (occurs in un-aerated cells); TDO,n = dissolved oxygen supplied (transferred) by the aeration device, if present; and IDO,n and EDO,n = dissolved oxygen loadings in the influent and effluent. Equation 11.42 is used to compute the value DDO,n when the operating dissolved oxygen in the un-aerated cell is specified. This can be from measurements at the plant or an estimated dissolved oxygen in the un-aerated cell. For the aerated cells, the model computes the amount of dissolved oxygen that needs to be transferred, Tn,DO, to achieve the dissolved oxygen set-point. It checks the value of Tn,DO against the capacity of the aeration system in cell n. If the value is exceeded, the model guides the user to lower the set-point or raise the capacity.
2.2 Numerical Approach to Solve One- and Two-Dimensional Biofilm-Diffusion Models (Steady-State and Dynamic Simulation) In the biofilm-diffusion model, one determines the VSS generated in each layer of the biofilm based on the substrate and electron acceptor conditions in the layer. The model computes the biofilm fluxes (jF) of COD, dissolved oxygen, biomass (VSS and
Development and Application of Models
inerts), and NOx-N from one layer (dz) to the next and integrates the values over the thickness of the biofilm (LF). jF =
∫
LF 0
rF dz = DF
dS dZ LF
∂S ∂ 2S = D 2 + rF ∂t ∂z
(11.43) (11.44)
Where
DF = diffusion coefficient of substrate S inside the biofilm and rF = rate of transformation of the substrate in a biofilm layer inside the biofilm. Equation 11.44 shows a dynamic simulation where the substrate conditions vary with time. For a steady-state simulation, the left side of eq 11.44 = 0. The flux from one concentric layer to the next deeper layer is the net of uptake and release in the layer and the flux from the concentric outer layer to this layer. For the VSS, the yield in each layer decreases with dissolved oxygen and NOx-N levels inside the biofilm. Therefore, the biomass yield for the biofilm decreases as its thickness increases. Based on this principle, the innermost layer of biofilm is the layer at which the total thickness of the biofilm reaches a point where the VSS flux computed by the model begins to drop below the biomass yield for the biofilm that is observed in full-scale applications. This condition defines the thickness of the biofilm. One should note that the VSS flux for the biofilm is the sum total of biofilm flux from the surface layer (surface shearing) and from the breakage (sloughing) of “chunks” of the biofilm. The breakage occurs periodically from the inner layer. When the shearing from the surface layer is set to zero, all of the biomass yield (generation) from the biofilm is from the periodic breakage from the inner layers. For the biofilm to be stable (not increase or decrease in thickness over time), the rate of breakage off the innermost layer equals the biofilm flux generated. It is important to understand the differences between the one-dimensional model (eqs 11.43 and 11.44) and the two-dimensional approach. The two-dimensional approach allows the modeler to vary the coverage of the media by the biofilm as a function of substrate conditions. Evidence from the field shows that biofilm coverage and thickness on a plastic media changes from the first cell to the last cell in a reactor. Also, they are different for high-strength wastewaters versus primary effluent from a municipal wastewater plant (e.g., MBBR operating with wastewater with 1500 to 2500 mg/L COD at a potato chip manufacturing plant versus municipal primary
505
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Biofilm Reactors
effluent). Most one-dimensional models have the ability to model the change in thickness (Reichart, 1998; Sen and Randall, 2008b; Wanner et al., 2006). The intelligence regarding the second dimension is a function of the media design and needs to be incorporated by upgrading the one-dimensional model. For the applications discussed in this chapter, the models have an additional (pseudo) dimension (as defined in by the IWA Biofilm Task Group in their discussion of AQUASIM [Wanner et al., 2006]), because the conditions vary along the length of the reactor. This is simulated using multiple cells in series (Figure 11.2). The following section discusses a finite difference technique to solve eqs 11.43 and 11.44 in the biofilm-diffusion model. As with the semi-empirical model, the equations are presented for NH4-N first and then related to other substrates. The fluxes computed by the diffusion model replace the fluxes computed using semi-empirical eqs 11.3, 11.5, and 11.18. The rest of the equations can be applied to compute the profile of substrates and electron acceptors in a multicell reactor using the finite difference model.
2.2.1 Ammonium-Nitrogen Equation 11.45 shows the NH4-N flux rate across a layer of biofilm. The NH4-N flux into the outer surface of a layer (left side of the equation) is equal to the NH 4-N flux out into the deeper layer plus the transformation of NH4-N in the layer (the units of the equation can be expressed in mg/cm2 biofilm surface/d [or other equivalent units]).
DN F DN F [SNbn− 1 − SNbn ] = [SNbn − SNbn+ 1 ]+ (N uptake − N releaase)bn L bn L bn+ 1
(11.45)
Where DNF = rate of diffusion of NH4-N in the biofilm (cm2/d); Lbn = thickness of biofilm layer, bn (cm); SN bn = NH4-N concentration in biofilm layer, bn (mg/cm3); Nuptake,bn is from nitrification and heterotrophic synthesis in biofilm layer, bn (mg/ 2 cm biofilm surface/d); and N release,bn is from decay of biomass in biofilm layer, bn (mg/cm 2 biofilm surface/d). The units can be converted to kg/1000 m2·d, which is a typical method of expressing flux rates for IFAS and MBBRs, by multiplying the flux rates in mg/cm2·d by 10. For the stagnant liquid layer (boundary layer liquid film) outside the biofilm, the equation takes the following form:
Development and Application of Models
DN L DN F [SNn − SNsL ] = [SNsL − SNb1 ] L sL L b1
(11.46)
Where DNL = rate of diffusion of NH4-N in bulk liquid and stagnant liquid layer (cm2/d). DNF = rate of diffusion of NH4-N in the biofilm. Because of resistance by the biomass, its value is 75 to 80% of the rate in the bulk liquid. LsL and Lb1 = t hickness of the stagnant liquid layer and biofilm layer 1, respectively. NsL = NH4-N level at the interface of the stagnant liquid layer and biofilm layer 1 (b1). SNn = NH4-N concentration in cell n of the multicell reactor. For biofilm layer bn =1, the equation takes the following form:
DN F DN F [SNsL − SNb1 ] = [SNb1 − SNb 2 ] + (N uptake − N release)b1 for bn= 1 (11.47) L b1 L b2
For biofilm layers bn = 2 to bnd–1 , where bnd is the deepest layer of biofilm, the equations take the following form: DN F DN F [SNb n − 1 − SNb n ] = [SNb n − SNb n+ 1 ]+ (N uptake − N release)bn for bn = 2 to bnd− 1 L L b b n+1 n
(11.48)
For the deepest layer (assuming a non-porous media surface, as observed with plastic and cord media), the equation takes the following form:
DN F [S Nbnd− 1 − S Nnd ] = (N uptake − N release)bnd for bn = bnd (bnd is the L bnd
deepest layer of the biofilm)
(11.49)
The NH4-N uptake (mg/d) by biofilm in cell n (BFn) of the IFAS or MBBR system is computed as follows:
BFn =
D NL [SNn − SNsL ] [biofilm surface area in cell n] L sL
(11.50)
507
508
Biofilm Reactors
Where SN n = NH4-N level in the bulk liquid in cell n (mg/cm3) and Biofilm surface area in cell n is in cm2.
2.2.2 Linkage to Equations 11.1 to 11.42 Presented Earlier The value of BFN,n can be converted to kilograms per day (kg/d) and substituted for BN,n and DN,n in eq 11.7. Therefore, the one- or two-dimensional biofilm-diffusion model substitutes for the semi-empirical equations and operates within the multicell model. BFN,n = BN,n + DN,n
(11.51)
Equation 11.7 had the following form: I N,n + N decay,n + N org-N,hydr,n = AN,n + BN,n + CN,n + DN,n + EN,n
Substituting equation 11.51 in 11.7 gives it the following form: I N,n + N decay,n + N org-N,hydr,n = AN,n + BFN,n + CN,n + EN,n
To run the one- or two-dimensional biofilm-diffusion model, the user must specify the following:
(1) (2) (3) (4)
Diffusion coefficients for the liquid and in the biofilm at temperature, T; Thickness of the stagnant liquid layer; Total thickness of the biofilm in each cell; and MLSS of the biofilm.
It is recommended that the outermost layers of the biofilm, where the transformation is faster, are simulated as thinner layers compared with the inner layers. Also, it is important that the IFAS and MBBR model be designed so that it guides the user through the selection of biofilm thickness for each cell. The thickness is a function of bulk liquid substrate and electron acceptor concentrations, media design, and mixing (turbulence). The bulk liquid concentration in a cell is a function of influent strength, recycle rates, and size and number of cells. For the same media and mixing intensities, the thickness will vary from cell 1 to cell n in a multicell reactor. The best source of thickness of biofilm on the media is data from similar wastewater treatment plants (WWTPs) operating with the same media. In general, a system should be designed with a level of turbulence that controls the thickness and maintains a biofilm MCRT of 10 to 80 days.
Development and Application of Models
The SCODbio, dissolved oxygen, and NOx-N concentrations in the bulk liquid are required for the flux computations. These are determined by the link to the multicell model. The fraction of nitrifiers in the first biofilm layer (fnitr, bn=1) must be specified based on measurements (Downing and Nerenberg, 2007) or computed. One method of computing fnitr, n=1 is discussed below. The NH4-N uptake by the biomass in a biofilm layer bn (Nuptake,bn) is the sum of uptake in cellular synthesis (DN,bn) and nitrification in the layer. Nuptake,bn = DN,bn + NH4-Nu,Nitr,bn
(11.52)
(11.53)
DN,bn = fN [(CODu,aer,bn)(YH,aer) + (CODu,anx, bn)(YHanx)]
Where Nuptake,bn = mg uptake/cm2 biofilm surface·d; and CODu,aer,bn and CODu,anx,bn = COD used aerobically and anoxically, respectively, in layer bn. The structures of the equations for COD removal in layer bn are similar to eqs 11.21 and 11.22 for the MLVSS. The aerobic COD uptake in layer bn is as follows (mg/cm2 biofilm surface·d): COD u,aer,bn = q m,H,aer,bfd
SCOD bio,bn DO bn L bn (1− fNitr,bn ) XF,bn K H,DO,bfd + DO K H,S,aer,bfd + SCOD bio,bn
(11.54)
Where XF,bn = MLVSS of layer bn (mg/cm3), Lbn = thickness of layer bn (cm), and Subscript “bfd” signifies coefficients for the biofilm-diffusion model. This is important because measurements in pilot studies (Sen, 1995) showed that certain parameters, such as the maximum COD utilization rate, can be different for biomass in biofilm and the MLVSS. The nitrification in layer bn is computed by eq 11.55 Its structure is similar to eq 11.6. NH4-N u,Nitr,bn = q m,NH4 -N-Nitr,bfd
DO bn NH 4 N bn L bn fNitr,bn XF,bn K N,DO,bfd + DO K N,bfd + NH 4 -N bn (11.55)
The MLVSS is computed based on the value of biofilm MLSS selected by the user and the fraction of VSS in layer bn computed by the model. The MLSS can be determined by measuring the thickness of the biofilm and the weight of growth per carrier particle.
509
510
Biofilm Reactors
The values reported in the literature range from 5000 to 50 000 mg/L MLSS (Boltz et al., 2009; Wanner et al., 2006). The lower end of the range has not been observed in MBBR and IFAS systems treating primary effluent. The biofilm in IFAS and MBBRs with plastic carrier particles (as at Broomfield, Colorado, Kaldnes K1 [AnoxKaldnes, Inc., Lund, Sweden]) is between 10 000 and 15 000 mg/L. Values between 25 000 and 50 000 mg/L are observed in Captor (Fairfield WWTP, Connecticut) and Linpor (Lotepro, Mixing and Mass Transfer Technologies, State College, Pennsylvania) sponge systems, where the growth is protected from shear inside the sponge (Sen, 1995; observations at Fairfield WWTP, Connecticut [Sen et al., 2000]). Because of the wide range of numbers, it is recommended that the biofilm MLVSS be measured from full-scale MBBRs and IFAS systems. The computation of percent of VSS and inerts and their fluxes is performed using a series of equations similar to those for NH4-N. The interdependencies between the equations for biomass, COD, NH4-N, dissolved oxygen, and NOx-N are similar to those shown for the semi-empirical version of the IFAS model discussed above. The fnitr,bn is computed for each layer in the biofilm. Unlike the complete mixed liquor, where fnitr does not change much along the series of reactors (cells) in series, fnitr,bn can change substantially from one layer to the next within the biofilm. Nrelease,bn (mg/cm2 biofilm surface/d) is computed similar to eq 11.10.
Nrelease,bn = (fN) (kdH,EA,T XF,bn Lbn)
(11.56)
where kdH,EA,T is selected based on the dissolved oxygen and NOx-N level in layer bn compared with aerobic and anoxic thresholds for dissolved oxygen and NOx-N that are specified by the user.
2.2.3 Chemical Oxygen Demand, Biomass (Volatile and Total Suspended Solids), Dissolved Oxygen, and NOx-N The equations for COD, biomass (VSS and inerts), dissolved oxygen, and NO x-N fluxes are similar to eqs 11.45 to 11.50 for NH4-N.
2.2.4 Biofilm Thickness, Growth, and Fraction Nitrifiers Biomass is lost from the biofilm through shearing and sloughing. Shearing is the loss of biomass from the surface layer. Sloughing is the periodic breakage of biomass of an inner layer. If the shear from the surface layer is assumed to be zero, all of the biomass flux is assumed to be from periodic breakage (sloughing) of biofilm off an inner layer. The rate of sloughing and the biofilm yield for heterotrophs and nitrifiers in each cell are dependent on the following:
Development and Application of Models
(1) Shear forces imparted by mixing and roll pattern; (2) Design of the biofilm carrier particle (media); and (3) Substrate and electron acceptor levels (i.e., COD, dissolved oxygen, NH4-N, and NOx-N levels in each cell).
The biofilm thickness is a function of the substrate concentrations in the bulk liquid (mixed liquor), average MLVSS of the biofilm, and shear forces. It is computed based on biofilm detachment rate and thickness of the biofilm. The detachment rate coefficient (kde), as modified by hydrodynamic shear factor, G; and media shape factor, Mn. The equation for biofilm detachment rate is as follows (Sen, Boltz, Copithorn, Morgenroth, 2009):
rde,Xk = (kde,n G/Mn) (XF,k)[(LF)n ]
(11.57)
Where rde,Xk = the detachment rate of biofilm biomass component Xk (g/m2/d). This value is computed from the substrate flux rate into the biofilm. G = factor to increase the shear for the type of aeration. Its value increases from fine bubble roll to coarse bubble to jet mixing. Mn = modifier for the type of media surface. Its value increases with roughness of the surface, increases from smooth plastic surface to cord. XF,k = the density of biomass component k in the biofilm (g/m3). LF = the thickness of the biofilm (m). For n = 1, the units of kde,1 are in d−1. For n = 2, the units of kde,2 are in m−1d−1. The biofilm one-dimensional model computes the biomass flux (heterotroph + nitrifier flux) in each cell for the biofilm thickness. The thickness and biomass flux then are used to compute a “Biofilm MCRT,” as follows:
Biofilm MCRT = (LF) (XF)/Biomass flux per day
(11.58)
In the diffusion model, the biofilm develops up to a thickness at which the biofilm yield computed for the thickness is equal to the observed biomass yield. The expected biomass yield (for each type of media at different substrate and electron acceptor levels and for specified mixing conditions) is obtained from pilot studies and from plants that are in operation. For the same substrate and electron acceptor levels, the biofilm yield decreases as its thickness increases. Therefore, the biofilm cannot develop to a thickness where its yield is below the observed yield. This can be used to check the computations.
511
512
Biofilm Reactors
One additional challenge in modeling the biofilm is the initial estimate of fnitr, in biofilm layer 1 (fnitr,b1). The fnitr,b1 can vary with the concentrations of SCODbio, dissolved oxygen, and NH4-N in the cell bulk liquid. The fnitr,b1 influences the fluxes of nitrifiers and substrates into the deeper layers and, therefore, influences the fnitr,bn computed for each of the deeper layers. The user can select the value of fnitr,b1 based on the relative rates of growth and yield of nitrifiers and heterotrophs in each layer. Monod kinetics are used to adjust the rates of growth that are influenced by the electron acceptor and substrate levels in each layer. For example, the growth rate for ammonia oxidizers in a layer bn would be as follows:
N g,Nitr1,bn = q m,NH4 -N-Nitr,bfd
DO bn NH 4 -N bn K N,DO,bfd+DO K N,bfd + NH 4 -N bn
(11.59)
The growth rate for nitrite oxidizers can be set equal to ammonia oxidizers or computed using the NO2-N concentration and the maximum growth rate and halfsaturation constants for nitrite oxidizers. The growth rates for heterotrophs are computed as follows: H g,bn = q m,H,aer,bfd
q m,H,anx,bfd
SCOD bio,bn DO bn + +DO K H,S,aer,bfd + SCOD bio,bn K H,DO,bfd SCOD bio,bn NO x -N bn K H,DO,bfd+NO x -N bn K H,S,aer,bfd + SCOD bio,bn
K H,DOi,bfd K H,DOi,bfd+DO
(11.60)
The fraction of nitrifiers in layer bn is as follows:
fNitr,bn = q m,NH4 -N-Nitr,bfd
N g,Nitr1,bn + N g,Nitr2,bn N g,Nitr1,bn + N g,Nitr2,bn +H g,bn
(11.61)
3.0 MODEL APPLICATIONS TO FULL-SCALE FACILITIES This section applies the models to full-scale IFAS and MBBR facilities. The purpose is to show how these models were applied to facilities and summarize the results. The authors have done a reasonable amount of calibration. Users can calibrate the models further. As mentioned earlier, some of the models can be modified to MBBRs, trickling filter, and biological aerated filters (BAFs), by varying the specific surface area of media. The characterization of the influent in all models discussed in this chapter can be based on the Water Environment Research Foundation (Alexandria, Virginia) report, Methods for Wastewater Characterization in Activated Sludge Modeling (Melcer et al., 2003).
Development and Application of Models
The influent fractions do not refer to specific chemical compounds; they merely divide COD, total Kjeldahl nitrogen (TKN), and phosphorus into groups that describe their behavior and fate in the model. The 5-day BOD (BOD5) is converted to COD using a ratio of COD/BOD5. Specialized sampling can be conducted to quantify some of the influent fractions. For example, soluble COD can be determined by measuring the filtered COD of the plant effluent and comparing it with the influent total COD. The SCODbio can be determined by subtracting the soluble COD that would remain after an extended period of aeration of the mixed liquor (the latter is the non-biodegradable soluble COD). The soluble readily biodegradable (SCODrbio or Fbs) can be determined by filtering and flocculating raw influent samples using a procedure described by Mamais et al. (1993) and then measuring the COD and subtracting the soluble non-biodegradable COD. The standard filtration technique is to use 0.45-µm filters to measure the soluble COD. Other levels of filtration use 0.1- to 1.4-µm filters. The soluble non-biodegradable TKN (SKNnbio or FNus) can be measured by taking an effluent sample, filtering it, and measuring the TKN, but the plant must be nitrifying fully when the sample is taken. With adequate characterization of the influent fractions, one should not have to adjust kinetic or yield parameters in the models, assuming that the plant has typical domestic wastewater without any inhibitory substances. If there is a potential for nitrification inhibition, rate tests should be conducted to properly calibrate the model. It also is important to conduct nitrification and denitrification rate tests during IFAS and MBBR pilot work or at similar full-scale plants. This can provide information on the thickness of the biofilm that develops for a certain level of turbulence and can help determine the aeration system and media placement for full-scale designs. The Water Environment Research Foundation report, Methods for Wastewater Characterization in Activated Sludge Modeling (Melcer et al., 2003) outlines ways to enhance model calibration efforts with additional sampling and testing. These must be adapted for IFAS and MBBR systems.
3.1 Integrated Fixed-Film Activated Sludge Plant Description and Modeling For the purposes of IFAS modeling, a plant in Broomfield, Colorado (Rutt et al., 2006; Sen et al., 2006) was selected and modeled in steady-state and dynamic modes. After analyzing the data from 2004 to 2006, the authors selected December 2005 and 2006 to run several models and demonstrate how they may be used. The results from December 2006 are presented below.
513
514
Biofilm Reactors
3.1.1 Integrated Fixed-Film Activated Sludge Plant Description The Broomfield, Colorado, plant is designed with influent screening and partial flow equalization in the headworks. It has primary clarifiers. The primary effluent is sent to the secondary treatment system that is operated in the A2O or modified Johannesburg (South Africa) configuration (Figure 11.6). The secondary treatment system is operated with Kaldnes K1 media in the aerobic cells at a 30% fill. This results in a biofilm specific surface area of 150 m2/m3 (when the media is sufficiently loaded in the winter season to create a biofilm on it). The mixed liquor is operated at approximately a 5-day MLSS MCRT, of which 3.25 days are aerobic. The plant has secondary clarifiers. Within the IFAS system (Figure 11.6), the RAS is sent to a cell that is designated as “pre-anoxic.” This designation was based on elevated levels of NOx-N in the primary effluent observed in samples before the IFAS process went online. The analysis of the data for 2006 showed that NOx-N levels are less than 1 mg/L in the primary effluent. Therefore, the first cell may be designated as “anaerobic,” which makes the modeled configuration equivalent to the A2O configuration. The total volume of the activated sludge tanks is 7015 m3. Of this, 15% is in the two anaerobic cells (also designated as one pre-anoxic and one anaerobic) that receive the primary effluent and the RAS; 20% is in the two pre-anoxic cells that also receive the mixed-liquor recycle; and 65% (4550 m3) is in the two aerobic cells. Both aerobic cells have the Kaldnes K1 media that is currently at 30% fill (mf = 30%). This media has a bare media surface area of 750 m2/m3 on the inside of the cylinder and on the two cross-vanes. The biofilm surface area is 500 m2/m3 based on 1-mm thickness and growth on the inner surface. Nitrate Recycle (0 to 200%) Partially Equalized Primary Effluent
Anaerobic
Anoxic Aerobic (with media) RAS (30 to 50%)
Anox Kaldnes K1 Media at 30% Media Fill Fraction Biofilm (Effective) specific surface area of 150 m2/m3
Figure 11.6 Layout of Broomfield WWTP, Colorado.
Secondary Clarifiers
Development and Application of Models
The aerobic cells are 4.5 m deep. They are aerated with coarse-bubble diffusers that are installed 0.25 m above the floor and concentrated at certain points within a grid on the floor (approximately 10% of the floor is covered with diffusers). This results in a roll pattern that is similar to a quarter-point arrangement (Rooney and Huibregtse, 1980; Water Pollution Control Federation, 1988). The media is retained by screens. The plant has two parallel trains that are similar.
3.1.2 Integrated Fixed-Film Activated Sludge Plant Operation The plant is operated similar to an activated sludge system, with additional monitoring for the media. The plant takes samples of the mixed liquor, strains it through a metal sieve to separate the media from the MLSS, and measures the MLSS, MLVSS, and amount of growth on the media (referred to as fixed growth below). To measure the growth, the media with the biofilm is dried and weighed. Its weight is compared with the weight of bare media without the biofilm. Alternatively, the oven-dried biomass (at 105°C) can be removed by physical scraping or chemical rinsing and the dry media weighed again to obtain the weight of the biomass. While the plant does not measure the biofilm thickness on a weekly basis, visual observations show that the biofilm is thinner in summer compared with winter. It is thicker in the first aerobic cell (approximately 1 +/-0.2 mm) compared with the second cell (approximately 0.6 mm +/-0.2 mm). To determine VSS in the biofilm, a sufficient number of carrier particles should be sampled to scrape (or squeeze out, in the case of sponge media) sufficient biomass and measure its total suspended solids (TSS) and VSS.
3.1.2.1 Data from December 2006 The operating data for December 2006 are shown in Table 11.6. 3.1.2.2 Flow and Recycle The primary effluent flow averaged 20 000 m3/d (5.3 mgd). The RAS was 40% of this flow (8000 m3/d). The operator estimated the nitrate recycle to be 150% (30 000 m3/d).
3.1.2.2.1 Primary Effluent The plant measures primary effluent flow and temperature daily; it measures the BOD5, TSS, NH4-N, NO2-N, and NO3-N five times per week and alkalinity once per week. The plant measures the COD/BOD5 ratio once per month. This ratio averaged 2.7 for 3 years of data (2004 to 2006) and was 2.67 for a single sample in December. While there was no measurement of the TKN in the primary effluent, discussions with AnoxKaldnes, Inc., indicated that their samples showed an average TKN of 40 mg/L for the typical primary effluent (110 to 120 mg/L BOD5). A TKN/NH4-N ratio and a TSS/particulate organic nitrogen ratio were
515
Table 11.6 Actual plant data used for IFAS process model (courtesy of Black & Veatch). Primary effluent and mixed liquor data from plant used as input to the dynamic simulation model Flow (m3/d)
Flow (mgd)
TSS (mg/L)
18 889
4.99
103
23 583
6.23
103
20 252
5.35
87
19 343
5.11
19 116
5.05
18 283 20 593
BOD5 (mg/L)
COD (mg/L)
NH4-N (mg/L)
122
325.74
27.8
32.0
32.0
0
122
325.74
27.8
32.0
32.0
0
122
325.74
25.8
29.7
29.7
<0.00
<0.01
0
88
85
226.95
24.3
27.9
27.9
<0.00
<0.01
0
92
109
291.03
28.3
32.5
32.5
0.01
0.01
0
4.83
93
114
304.38
27.8
32.0
32.0
0.02
<0.01
236
5.44
88
105
280.35
28.8
33.1
33.1
0.02
<0.01
0
19 343
5.11
103
122
325.74
27.8
32.0
32.0
17 829
4.71
103
122
325.74
27.8
32.0
32.0
20 100
5.31
85
107
285.69
25
28.8
28.8
0.02
<0.01
0
19 192
5.07
93
98
261.66
27
31.1
31.1
0.02
0.01
0
19 116
5.05
100
136
363.12
27
31.1
31.1
0.02
0.02
0
18 548
4.9
97
125
333.75
32.3
37.1
37.1
0.23
<0.01
267
17 564
4.64
101
149
397.83
33.8
38.9
38.9
0.02
0.03
0
18 019
4.76
103
122
325.74
27.8
32.0
32.0
0
19 835
5.24
103
122
325.74
27.8
32.0
32.0
0
18 586
4.91
91
144
384.48
27.3
31.4
31.4
0.03
<0.01
0
21 350
5.64
96
110
293.7
25.3
29.1
29.1
0.02
0.02
0
18 473
4.88
99
126
336.42
28.3
32.5
32.5
0.01
0.03
0
19 495
5.15
94
91
242.97
30.3
34.8
34.8
0.02
<0.01
249
19 495
5.15
122
154
411.18
30.5
35.1
35.1
0.02
<0.01
0
19 306
5.1
103
122
325.74
27.8
32.0
32.0
0
19 911
5.26
103
122
325.74
27.8
32.0
32.0
0
20 441
5.4
164
189
504.63
28
32.2
32.2
0.02
0.02
0
17 224
4.55
116
129
344.43
28
32.2
32.2
0.01
<0.01
0
18 208
4.81
142
133
355.11
28
32.2
32.2
0.02
0.01
0
20 744
5.48
105
94
250.98
24.5
28.2
28.2
0.03
<0.01
239
20 214
5.34
105
111
296.37
25.25
29.0
29.0
0.02
0.02
0
20 100
5.31
103
122
325.74
27.8
32.0
32.0
0
20 744
5.48
103
122
325.74
27.8
32.0
32.0
0
20 971
5.54
97
137
365.79
28
32.2
32.2
0
19 512
5.15
103
27.791 94
32.0
32.0
122.1935 326.256 8
516
SKN (mg/L)
TKN (mg/L)
NO2-N (mg/L)
NO3-N (mg/L)
Alkalini (mg/L)
0 0
0.031 111
0.018 889
248
NO3-N mg/L)
Effluent data (this is the reported data, not model output) Alkalinity (mg/L)
MLSS temperature (°C)
NH4-N (mg/L)
NO3-N (mg/L)
NO2-N (mg/L)
Temperature (°F)
MLSS (mg/L)
VSS (%)
MLVSS Alkali(mg/L) nity (mg/L)
0
19.4
67
1560
0
0
19.4
67
1660
0
.01
0
19.4
0.15
15.07
0.03
67
1580
0
.01
0
15.6
0.11
14.24
0.06
60
1651
0
.01
0
16.1
0.07
15.37
0.03
61
1793
0
.01
236
16.1
0.11
16.38
0.24
61
1785
.01
0
16.1
0.08
16.38
0.02
61
1726
0
0
16.1
61
1732
0
0
16.1
61
1712
0
.01
0
16.1
0.09
16.09
0.01
61
1584
0
.01
0
15.6
0.16
14.79
0.02
60
1616
0
.02
0
15.6
0.11
16.29
0.01
60
1656
0
.01
267
17.2
0.87
19.83
0.17
63
1298
0
.03
0
16.1
0.17
18.07
0.03
61
1720
0
0
16.1
61
1766
0
0.77
0
0
15.6
60
1694
0
.01
0
15.6
0.23
13.49
0.01
60
1646
0
.02
0
15.6
0.09
11.38
0.02
60
1633
0
.03
0
15.6
0.18
11.03
0.07
60
1633
.01
249
15.6
0.18
11.38
0.02
60
1848
.01
0
13.3
0.18
17.49
0.02
56
1633
0
0
0
13.9
57
1617
0
13.9
57
1651
0
.02
0
13.9
10.77
0.03
57
1680
0
.01
0
13.3
12.68
0.02
56
1633
0
.01
0
15.0
12.28
0.02
59
1588
0
.01
239
15.0
12.78
0.02
59
1568
.02
0
15.0
12.77
0.03
59
1504
0
0
13.9
57
1476
0
0
14.4
58
1490
0
0
13.3
56
1500
0
248
15.6
.018 889
0.19
0.03
14.4
1633
517
86
0 0.76
0
13.47
78
0.81
0.78
0
0
100
97
90
518
Biofilm Reactors
used to generate the TKN levels for December 2006 (Table 11.6). Additionally, there were a few measurements that showed that 80% of the COD measured as filtered COD. There were no data on the VSS in the primary effluent. Based on typical data, the VSS were estimated to be 80 to 85% of the TSS.
3.1.2.2.2 Aerobic Cells The plant measures MLSS and MLVSS in the aerobic cells and the amount of biomass in the biofilm. In December 2006, the MLSS averaged 1630 mg/L, and the MLVSS averaged 1270 mg/L, resulting in a percent VSS of 78%. The “fixed biomass” averaged 2050 mg/L in the first aerobic cell and 1104 mg/L in the second aerobic cell. For a cell volume of 2271 m3 and a specific surface area of 150 m2/m3, this quantity of fixed biomass equates to 13.7 and 7.4 kg/1000 m2 of biofilm surface. At a biofilm density of 12.5 kg/m3 (12 500 mg/L), the biofilm thickness would be 1.1 mm in the first aerobic cell and 0.6 mm in the second cell. The mixed-liquor temperature dropped from 19°C at the beginning of the month to 13°C at the end of the month. The dissolved oxygen averaged 4.2 mg/L in the first aerobic cell and 5.6 mg/L in the second aerobic cell. The NH 4-N and NOx-N levels were not measured in the anaerobic, anoxic, and aerobic cells.
3.1.2.2.3 Secondary/Plant Effluent The plant achieved complete nitrification in December 2006, with effluent averages less than 0.2 mg/L for NH4-N. The NOx-N level was 14.4 mg/L, of which less than 0.1 mg/L was as NO2-N. The NO3-N was higher than expected, the reasons for which are discussed later. The effluent TSS and VSS in the plant effluent were less than 5 mg/L. 3.1.2.2.4 Discussion of the Data All three models presented below indicated substantial bleed-through at 150% nitrate recycle (2 to 4 mg/L). In the absence of measurements of profiles of NH4-N and NOx-N along the IFAS tank (anaerobic, anoxic, and aerobic cells), it was not possible to verify whether the actual nitrate recycle was closer to 100% instead of 150%, or if there was significant bleed-through of NOx-N from the anoxic to the aerobic cells. If the profiles show significant bleed-through, this would indicate a need for media in the anoxic cells to supplement the denitrification achieved. This is an example of a significant finding from the modeling that can help improve the design and operation.
3.1.3 Modeling Integrated Fixed-Film Activated Sludge in Aquifas Aquifas can model the biofilm in the IFAS system by two different approaches.
(1) Semi-empirical equations and
Development and Application of Models
(2) One- and two-dimensional biofilm modeling.
The advantage of the semi-empirical equations is that computations are much faster, especially for dynamic simulation of 31 days of data. The advantage of the one- and two-dimensional biofilm modeling, where Aquifas breaks the biofilm up into 12 layers, is that it provides additional data, such as the biofilm thickness for each cell and biofilm yield. Also, it provides a second method of computation of flux rates. Because the semi-empirical equations are calibrated to real-life conditions, they provide the requisite accuracy. The section below presents the results of both models and compares them with each other and with the plant data. Aquifas can be run with different levels of primary effluent characterization. For this run, the influent was characterized as described above, which is similar to requirements for BioWin and GPS-X. The denitrification kinetics for the MLVSS (in the absence of media in the anoxic cells) were based on the estimated value of flocculated filtered biodegradable COD (40% of the COD and half the filtered COD). The phosphorusremoval mechanisms in the anaerobic cells were driven by the volatile fatty acids in the primary effluent and those generated through fermentation in the anaerobic cells.
3.1.3.1 Results from Aquifas Table 11.7 shows the average values for December 2006 from the plant and those predicted by a 31-day dynamic simulation using the semi-empirical and biofilm onedimensional models. The input to the model is the data in Table 11.6. The results of the 31-day dynamic simulation showed that both the semi-empirical model and the biofilm one-dimensional model were able to accurately predict the effluent NH4-N, NOx-N , alkalinity, and phosphorus (Table 11.7 and Figure 11.7). However, there was a higher degree of granularity in the NH4-N prediction on a dayto-day basis with the biofilm one-dimensional model. Additionally, the models were accurate in their prediction of the average MLSS, MLVSS, percent VSS, and waste activated sludge (WAS) production. The COD, NH4-N, and NO3-N profiles from the semi-empirical model and biofilm one-dimensional models are shown in Figure 11.8. Both models predict similar profiles. Also, they show significant bleed-through of NO3-N from the anoxic to the aerobic cells; there were no data on NO3-N in the anoxic cell to verify this. The high NO3-N concentration indicates one of two possibilities—either the nitrate recycle pump was not operating at the rate applied for the modeling, or there is a need for media in the anoxic cells.
519
520
Biofilm Reactors
The removals in the biofilm and the mixed liquor are shown in Figure 11.9. The semi-empirical model shows a slightly higher nitrification in the biofilm, while the biofilm one-dimensional model predicts a slightly higher denitrification in the biofilm. This explains the slightly better nitrification predicted by the semi-empirical model (closer to the level measured at the plant) and the slightly higher denitrification Table 11.7 Comparison of 31-day average from dynamic simulation against plant data Aquifas semi-empirical and biofilm one-dimensional models. Semiempirical model
Biofilm onedimensional model Plant data
Primary effluent (input) Flowrate (m /d)
19 512
19 512
19 512
Flowrate (mgd)
5.33
5.33
5.33
BOD (mg/L)
122
122
122
TSS (mg/L)
103
103
103
VSS
82
82
COD (mg/L)
326
326
SCOD (mg/L) (flocculated)
130
130
TKN (mg/L)
41.7
41.7
NH3-N (mg/L)
27.8
27.8
Total phosphorus (mg/L)
5
5
PO4-P (mg/L)
3
3
3
27.8
MLSS (mg/L)
1650
1660
1673
MLVSS (mg/L)
1310
1315
1271
Percent VSS
79
79
78
Dissolved oxygen (first cell) (mg/L)
4.2
4.2
4.2
Dissolved oxygen (second cell) (mg/L)
5.6
5.6
5.6
3.3
3.3
3.3
It was assumed that the filtered flocculated COD was 50% of the measured soluble COD
326
IFAS process (output)
MLSS SRT (oxic) (days)
Notes
Fixed biomass results are for the average condition at 14°C. The rest of the data is the average of 31-day dynamic simulation.
(continued)
Development and Application of Models
Table 11.7 Continued Semiempirical model
Biofilm onedimensional model Plant data
Fixed biomass (first cell) (g/m2)
11.7
13.7
Fixed biomass (second cell) (g/m2)
8.4
7.4
2180 (4807)
2177 (4800)
WAS total solids (kg/d) (lb/d)
2173 (4790)
Secondary effluent (output) BOD (mg/L)
5.0
5.4
2.6
TSS (mg/L)
5.0
5.0
4.5
TKN (mg/L)
1.3
1.4
NH3-N (mg/L)
0.26
0.33
0.19
NO3-N (mg/L)
14.5
13.2
14.4
Total phosphorus (mg/L)
0.1
0.1
0.1
PO4-P (mg/L)
0.1
0.1
0.1
Alkalinity
94
106
90
Notes
Effluent TKN is based on an assumed value of 0.5 mg/L non-biodegradable SKN.
in the biofilm one-dimensional model (lower effluent NO3-N compared with the plant data). While additional calibration of the models to this month’s data would eliminate this difference for December 2006, the calibration parameters were selected because they also were appropriate for other months modeled (e.g., December 2005, when the plant’s effluent NO3-N was lower, at 10.5 mg/L). The analysis of the data for the semi-empirical model showed that 32% of the TKN taken up across the reactor (for cell synthesis and nitrification) was taken up by the biofilm (this number was 25% when computed using the biofilm one-dimensional model) compared with just 4% for COD (4% computed using the biofilm one-dimensional model). Therefore, 68% of the TKN and 96% of the COD was converted/consumed by biomass in the mixed-liquor VSS. The semi-empirical model showed that 47% of the
521
IFAS, Full Scale Plant B Model Eff NH4N
Plant NH4N
1.50 1.00 0.50 0.00
2.00
Effluent NH4N, Model & Plant, mg/L
Effluent NH4N, Model & Plant, mg/L
2.00
IFAS, Full Scale Plant B Model Eff NH4N Plant NH4N
1.50 1.00 0.50 0.00
1 2 3 4 5 6 7 8 9 10 11 1213 14 15 16 17 18 1920 21 22 23 24 25 26 2728 29 30 31
1 2 3 4 5 6 7 8 9 10 11 1213 14 15 16 17 18 1920 21 22 23 24 25 26 2728 29 30 31 Date of Month, Dec 2006
Date of Month, Dec 2006
IFAS, Full Scale Plant B
IFAS, Full Scale Plant B Model Eff NO3N
Plant NO3N
20.0 15.0 NO3N
10.0
25.0 Effluent NO3N, Model & Plant, mg/L
Effluent NO3N, Model & Plant, mg/L
25.0
5.0 0.0
Model Eff NO3N
20.0 15.0 NO3N
10.0 5.0 0.0
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31
522
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31
Date of Month, Dec 2006
Date of Month, Dec 2006
Plant MLSS
Model MLVSS
Plant MLVSS
2000 1500 1000 500 0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 Date of Month, Dec 2006
MLSS & MLVSS, Model & Plant, mg/L
MLSS & MLVSS, Model & Plant, mg/L
IFAS, Full Scale Plant B Model MLSS
Plant NO3N
Model MLSS
IFAS, Full Scale Plant B Plant MLSS Model MLVSS
Plant MLVSS
2000 1500 1000 500 0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 Date of Month, Dec 2006
Figure 11.7 Aquifas output: The figures in the left column are from the semi-empirical model; the figures in the right column are from the biofilm one-dimensional model. The diffusion model can offer a higher degree of precision in its ability to predict day-to-day variations in a dynamic simulation, but takes substantially longer to run. Both models are able to predict the diurnal and 31-day average.
SCODbio SCOD COD, mg/L
COD, mg/L
COD Profile
COD Profile
90.0 80.0 70.0 60.0 50.0 40.0 30.0 20.0 10.0 0.0 Anaerobic
Anaerobic
Anoxic
Anoxic
Aerobic
Aerobic
90.0 80.0 70.0 60.0 50.0 40.0 30.0 20.0 10.0 0.0
SCODbio SCOD
Anaerobic
Clarifier Feedwell
Anaerobic
Anoxic
Cells in System
30.0 NH4N SKN N, mg/L
N, mg/L
15.0
Aerobic
Clarifier Feedwell
NH4N SKN
25.0
20.0
20.0 15.0
10.0
10.0
5.0
5.0 0.0
523
0.0 Anaerobic
Anaerobic
Anoxic
Anoxic
Aerobic
Aerobic
Anaerobic
Clarifier Feedwell
Anaerobic
Anoxic
Oxidized N Profile
16.0
Aerobic
Aerobic
Clarifier Feedwell
Aerobic
Aerobic
Clarifier Feedwell
Oxidized N Profile
14.0
NO3N
14.0
Anoxic Cells in System
Cells in System
12.0
12.0
10.0
10.0
N, mg/L
N, mg/L
Aerobic
NH4N and SKN Profile
NH4N and SKN Profile 30.0 25.0
Anoxic Cells in System
8.0 6.0
NO3N
8.0 6.0 4.0
4.0
2.0
2.0
0.0
0.0 Anaerobic
Anaerobic
Anoxic
Anoxic Cells in System
Aerobic
Aerobic
Clarifier Feedwell
Anaerobic
Anaerobic
Anoxic
Anoxic Cells in System
Figure 11.8 Substrate profiles for COD, NH4-N, and NO3N. The left column is from the semi-empirical model; the right column is from the biofilm one-dimensional model.
COD Uptake in Biofilm and MLVSS (kg/d)
COD Uptake in Biofilm and MLVSS (kg/d) 3000.000
Biofilm Mixed Liquor VSS
2500.000
COD Uptake, kg/d
COD Uptake, kg/d
3000.000
2000.000 1500.000 1000.000 500.000
Biofilm Mixed Liquor VSS
2500.000 2000.000 1500.000 1000.000 500.000 0.000
0.000 Anaerobic
Anaerobic
Anoxic
Anoxic
Aerobic
Aerobic
Anaerobic
Clarifier Feedwell
Anaerobic
Anoxic
Aerobic
Aerobic
Clarifier Feedwell
NH4N Uptake in Biofilm and MLVSS (kg/d)
NH4N Uptake in Biofilm and MLVSS (kg/d) 350.000
350.000 NH4N Uptake, kg/d
Mixed Liquor VSS
250.000
300.000
Biofilm
250.000
Mixed Liquor VSS
NH4N Uptake, kg/d
Biofilm
300.000
200.000
200.000
150.000
150.000
100.000
100.000
50.000
50.000
524
0.000
0.000 Anaerobic
Anaerobic
Anoxic
Anoxic
Aerobic
Aerobic
Anaerobic
Clarifier Feedwell
Anaerobic
Oxidized N Denitrified, kg/d
Biofilm Mixed Liquor VSS
100.000
Anoxic
Anoxic
Aerobic
Aerobic
Clarifier Feedwell
Denitrification in Biofilm and MLVSS
Denitrification in Biofilm and MLVSS 120.000 Oxidized N Denitrified, kg/d
Anoxic
80.000 60.000 40.000 20.000
120.000
Biofilm Mixed Liquor VSS
100.000 80.000 60.000 40.000 20.000 0.000
0.000 Anaerobic
Anaerobic
Anoxic
Anoxic
Aerobic
Aerobic
Clarifier Feedwell
Anaerobic
Anaerobic
Anoxic
Anoxic
Aerobic
Aerobic
Clarifier Feedwell
Figure 11.9 COD uptake, NH4-N uptake, and denitrification in the biofilm and MLVSS. The left column is from the semi-empirical model; the right column is from the biofilm one-dimensional model.
Development and Application of Models
nitrification is in the biofilm (35% for the biofilm one-dimensional model); the remaining 53% was in the mixed-liquor VSS. In the aerobic zone, the contribution of the biofilm toward denitrification was not significant. Only 2.5% of the overall denitrification was in the biofilm (12% for the biofilm one-dimensional model). This is because of the high dissolved oxygen concentration in the aerobic cells and a fairly high turbulence associated with coarse-bubble diffusers. The turbulence can reduce the depth of the stagnant liquid layer and increase the dissolved oxygen levels within the biofilm (Figure 11.10). The quantity of growth (fixed biomass) measured on media in the two aerobic cells in series compared satisfactorily with the computed values from the biofilm one-dimensional model (Table 11.8). The substrate, electron acceptor, and fraction
DO NH4N NOxN SCODbio %VSS /10
DO 14 NH4N 12 NOxN SCODbio 10 % VSS / 10 8
14 12 10 mg/L
8
mg/L
6 4
6 4
2
2
0 –2
–0.5
0
0
0.5
1
–0.4
Distance from Surface of biofilm, mm
–0.2
–2
0
0.2
0.4
0.6
0.8
Distance from Surface of biofilm, mm
Figure 11.10 Substrate profiles (dissolved oxygen, NH4-N, NOx-N, SCODbio, and percent VSS) inside the biofilm in aerobic cells 1 and 2.
Table 11.8 Measured and computed values of biofilm growth (Aquifas biofilm onedimensional model). Biofilm growth
kg/1000 m2·d
11.7
8.5
Observed biofilm growth
kg/1000 m ·d
13.7
7.4
Thickness as computed
µm*
932
679
Heterotrophic yield
0.290
0.249
Autotrophic yield
0.038
0.007
Fnitr in layer 1 of biofilm as used in the model
46.29%
41.13%
24.3
51.1
Biofilm MCRT *µm = micrometers (1000 µm = 1 mm).
2
days
525
526
Biofilm Reactors
VSS profiles within the 12 layers inside the biofilm in aerobic cells 1 and 2 are shown in Figure 11.10.
3.1.3.2 Key Inputs to Aquifas Biofilm One-Dimensional Model The user must provide the model with information on the MLSS in the biofilm. This should be based on the measurement of fixed solids and thickness of the biofilm on the carrier particle in each cell. In this plant, the growth on the carrier particle is measured once per week in each aerobic cell. The Aquifas biofilm two-dimensional model requires information on biofilm surface area changes on each carrier particle based on substrate conditions and biofilm thickness.
3.1.3.3 Discussion of Aquifas Model and Accuracy of Results In general, the Aquifas semi-empirical model was able to predict the effluent satisfactorily and is significantly faster to run than all three biofilm one-dimensional models discussed in this chapter (Aquifas biofilm one-dimensional model, BioWin, and GPS-X). In the Aquifas biofilm one-dimensional model (also called Aquifas 4), the biofilm is broken up into 12 layers. The layers that are close to the surface are kept thinner than the deeper layers, because there are substantial changes in the concentrations of the substrates (Figure 11.10). The total number of layers is higher than other models; the variation in thickness across the depth of the biofilm as a function of the steepness of the substrate profile is an additional feature. The fraction of nitrifiers and VSS are allowed to vary from layer to layer. The model computes the thickness of the biofilm based on the substrate and electron acceptor conditions (eq 11.57). The user specifies a value for the shear factor (range 1 to 5, with a higher number for higher turbulence). This adjusts the thickness of the biofilm. For example, at this facility, which uses coarse-bubble aeration and a relatively vigorous roll pattern, the value of this parameter is between 3 and 4. For a gentler roll pattern, as may be encountered with ceramic fine-bubble diffusers (Port Neuf, Quebec, Bioportz media), or a media with a different structure (e.g., cord media), this factor is expected to be between 1 and 3. For anoxic cells, the values may vary from 0.25 to 2.5. The user is able to incorporate information on the changes in biofilm specific surface area (biofilm two-dimensional modeling) as a function of substrate concentration and biofilm thickness. The additional level of detail in the Aquifas model relative to the other biofilm one-dimensional models requires a numerical procedure with a level of stiffness, as discussed by Wanner et al., 2006. In the Aquifas biofilm one-dimensional model, this is provided by manually entering the initial estimates of the NH4-N and
Development and Application of Models
NO3-N concentrations in each cell of a multicell reactor (to initiate the iterative computations). The initial estimates can be based on the semi-empirical model run or an observed substrate profile from the plant. The model is run until the iterations converge to a condition where the flux computed by the next iteration is not significantly different from the previous iteration.
3.1.4 Modeling in BioWin 3.1.4.1 Framework The BioWin version 3.0 (released on February 1, 2007) contains a one-dimensional biofilm model. The underlying principles of the biofilm model are based on the one-dimensional model framework of Reichert and Wanner (1997) and Wanner and Reichert (1996), in which a multilayer biofilm grows perpendicular to the media surface. The model includes diffusion of soluble and particulate components, a boundary layer that serves as diffusion resistance for solutes, and the exchange of particulates resulting from detachment (erosion) and attachment (impingement) of solids. The biofilm thickness changes dynamically based on substrate loading and solids exchange, though the user specifies a minimum and maximum biofilm thickness. The model also forms “streamers” when the biofilm solids concentration (gTSS/m2 surface area) exceeds a user-specified value. These streamers increase the diffusion into the biofilm, by increasing the available surface area (for modeling porous films or high-turbulence situations). There also are provisions in the model to adjust the “stickiness” of parameters, including active microorganisms, to control their overall densities within the biofilm. For example, inert solids are less “sticky” than organic variables, which explains the high VSS/TSS ratios measured in many biofilm systems. The BioWin biofilm model is integrated to the full general activated sludge/ anaerobic digestion model (ASDM), which also includes a chemical equilibrium, precipitation, and pH module (Envirosim Associates, Ltd., 2006). In BioWin, the user specifies influent or primary effluent COD (or BOD, though the model will convert BOD to COD to use as the basis for its computations). Influent COD fractions also must be specified to define the quantities of COD that are soluble, particulate, biodegradable, and non-biodegradable. The fraction of readily biodegradable COD (Fbs) in the model is important, because it is used by heterotrophic organisms for denitrification and is fermented into acetate and propionate, which is used by phosphorus-accumulating organisms. A portion of the influent Fbs is already acetate or propionate, and this is defined by the fraction Fac. A small amount of COD is soluble and non-biodegradable (Fus), and this material passes through the
527
528
Biofilm Reactors
model unchanged. Non-biodegradable particulate COD defined by the fraction Fup is not transformed biologically in the model, but can be removed physically in the clarifier or solids-separation devices. The remaining COD is slowly biodegradable and particulate, and the user must specify how much of this is not colloidal or biomass (Fxsp). Other parameters describe nutrient fractions, such as FNus—the fraction of TKN that is non-biodegradable and soluble (and passes through the model unchanged). Figure 11.11 and Table 11.9 summarize the COD fractions defined within BioWin. Other information that is essential for model calibration is solids inventory data— MLSS, MLVSS, WAS flowrate, TSS, and attached biomass TSS and VSS (in grams per square meter of surface area, or equivalent MLSS or MLVSS). Actual biofilm thicknesses, attachment rates, and detachment rates are much more difficult to quantify, so it is recommended to start with the default values, which are listed in Table 11.10. Readily Biodegradable COD, Essential for BNR, Acetate and Propionate 85 + 15 = 100 mg/L (VFA as COD), Fbs= 100/500 = 0.20 Essential for biological phosphorus removal Fac= 15/100 = 0.15
Slowly Biodegradable & Particulate COD, must hydrolyze first, then degrades biologically 232 + 78 = 310 mg/L Fxsp= 232/310 = 0.75
85 mg/L 15 mg/L
25 mg/L 232 mg/L
65 mg/L
Total COD = 500 mg/L
Colloidal and/or Biomass, 78 mg/L
Unbiodegradable ,& Soluble COD cannot degrade; passes through unchanged Fus= 25/500 = 0.05
Unbiodegradable & Particulate COD, cannot degrade, but can be removed by settling or filtering Fup= 65/500 = 0.13
Figure 11.11 Typical raw influent COD fractions in BioWin (Envirosim Associates, Ltd., 2006).
Table 11.9 Summary of the BioWin influent fractions (Envirosim Associates, Ltd., 2006). Fraction Description
Fate in the model
BioWin default
Typical range Typical range for raw for primary influent effluent
0.05
0.04 to 0.10
0.05 to 0.20
Fus
Passes through Fraction of unchanged total COD that is nonbiodegradable and soluble
Fup
Fraction of total COD that is nonbiodegradable and particulate
0.13 Remains unchanged, but can be settled or filtered
0.07 to 0.22
0.13 to 0.20
Fbs
Fraction of total COD that is biodegradable and soluble
0.16 Readily degrades; important for biological nutrient removal
0.05 to 0.25
0.08 to 0.35
Fxsp
Must hydrolyze Fraction first, then of slowly degrades degradable COD that is not biomass or colloids
0.75
0.70 to 0.80
Varies
Fac
Fraction of biodegradable soluble COD that is acetate or propionate
0.15
0 to 0.50
0 to 0.50
FNus
Fraction of Passes through TKN that the plant is nonunchanged biodegradable and soluble
0.02
0.01 to 0.05
0.03 to 0.07
Readily degrades; important for biological phosphorus removal
529
Table 11.10 Steady-state BioWin calibration results for a full-scale IFAS facility.* BioWin result
Plant data Notes
Primary effluent (input) Flowrate (m /d)
19 512
19 512
Flowrate (mgd)
5.33
5.33
BOD (mg/L)
126
126
TSS (mg/L)
105
105
VSS (mg/L)
89
-
TKN (mg/L)
33
-
NH3-N (mg/L)
27
27
Total phosphorus (mg/L)
9.6
PO4-P (mg/L)
7.0
3
IFAS process (output) MLSS (mg/L)
1580
1633
MLVSS (mg/L)
1040
1334
Percent VSS
66
78
Dissolved oxygen (first cell) (mg/L)
4.2
4.2
Dissolved oxygen (second cell) (mg/L)
5.6
5.6
MLSS SRT (oxic) (days)
3.3
3.3
13.7
13.7
Fixed biomass (second cell) (g/m )
7.5
7.4
WAS total solids (lb/d)
4400
4800
Fixed biomass (first cell) (g/m ) 2
2
Secondary effluent (output) BOD (mg/L)
2.6
1.7
TSS (mg/L)
4.5
4.5
TKN (mg/L)
2.6
NH3-N (mg/L)
0.40
0.19
NO3-N (mg/L)
14.7
14.4
Total phosphorus (mg/L)
<1
0.1
PO4-P (mg/L)
0.1
VSS, TKN, and total phosphorus data not available and were assumed.
Influent VSS data not available, though the model predicted a low MLVSS ratio (0.66) compared with the plant data.
The model predicted conservatively. High dissolved oxyen carry-over into the clarifier likely contributed to lower BOD and NH3-N concentrations measured at the plant.
*Note: This is not from a 31-day dynamic simulation. It is a steady-state run on average conditions observed in December 2006.
530
Development and Application of Models
Figure 11.12 Screen shot of IFAS model in BioWin.
3.1.4.2 Results from BioWin The IFAS model was set up as shown in Figure 11.12. It calibrated reasonably well to historical data, with no adjustment of biofilm parameters. Two ordinary heterotrophic organism (OHO) kinetic parameters were varied from the default values to mimic the level of denitrification observed at the plant. The anoxic growth factor was increased from 0.5 to 1.0, and the anoxic hydrolysis factor was increased from 0.28 to 0.50. Biofilm thickness limits were varied in individual reactors to mimic field observations; the “thin” film limit was increased in the first reactor from 0.5 to 0.75 mm, and the “thick” film limit was decreased in the second reactor from 3 to 1 mm. All other kinetic, stoichiometric, and biofilm parameters were left at the default values. The steady-state model results are summarized in Table 11.10. This was run on the average conditions for the month. The results from the dynamic model runs are shown in Figures 11.13, 11.14, and 11.15. 3.1.4.3 Discussion of Results from BioWin Overall, the model predicted accurate results with a reasonable amount of effort. One parameter, the MLVSS concentration, did not calibrate perfectly, but reached
531
Effluent NH3N and NO3N (mg/L)
25.0
20.0 Nitrate
15.0
10.0
5.0 Ammonia 0.0 12/1/06
12/7/06
12/13/06
12/19/06
BioWin Predicted Effluent NH3N BioWin Predicted Effluent NO3N
12/25/06
1/1/07
Measured Effluent NH3N Measured Effluent NO3N
Figure 11.13 BioWin IFAS calibration results for effluent ammonia and nitrate.
Mixed Liquor Suspended Solids (mg/L)
3,000 2,500 2,000 1,500 1,000 500 0 12/1/06
12/7/06
12/13/06
12/19/06
BioWin Predicted MLSS
Measured MLSS
Figure 11.14 BioWin IFAS calibration results for MLSS. 532
12/25/06
1/1/07
Development and Application of Models
Effluent BOD and TSS (mg/L)
25.0
20.0
15.0
10.0
5.0
0.0 12/1/06
12/7/06
12/13/06
12/19/06
BioWin Predicted Effluent BOD BioWin Predicted Effluent TSS
12/25/06
1/1/07
Measured Effluent BOD Measured Effluent TSS
Figure 11.15 BioWin IFAS calibration results for effluent BOD and TSS.
sufficient accuracy for this exercise. Despite the primary influent VSS/TSS ratio being set to 0.85, the model predicted a MLVSS/MLSS ratio of 0.66. Because the intent of this modeling was to calibrate to effluent quality and MLSS solids retention time, the MLVSS was not investigated further. The biomass inventory (g/m 2) calibrated well, but required slight adjustments to the film thickness limits to mimic the specific biomass inventories measured in individual IFAS zones at the full-scale facility. It is common for the first IFAS zone in series to have more biomass, as a result of the higher organic loading rate; therefore, it is imperative to collect as much calibration data as possible, with specific attention to nutrient and dissolved oxygen concentrations, and biomass inventories within each zone of the process. Overall, the BioWin biofilm model is a reliable tool for modeling IFAS systems. Steady-state solutions are achieved quickly, although modeling long periods of dynamic data requires more time. One way to expedite dynamic simulations is to switch “off” the pH modeling, especially if chemical phosphorus removal is not part of the design.
533
534
Biofilm Reactors
3.2 Moving-Bed Biofilm Reactor Plant Description and Modeling The South Adams Plant in Adams County, Colorado, was selected for application of the GPS-X and Aquifas models to an operating MBBR. The plant has two parallel trains, each with two anoxic and two aerobic basins (tanks or cells) in series. The MBBR has a nitrate recycle (MBBR recycle) from the end of the second aerobic cell to the first anoxic cell (Figure 11.16). There are various ways to introduce the concept of an MBBR model. An MBBR may be conceived as a flooded trickling filter. It also may be conceived as an IFAS without the RAS recycle. The absence of RAS brings the MLSS MCRT down to the nominal HRT of the basin. The nominal HRT of a South Adams MBBR is equal to the volume of cells available for mixed liquor divided by the primary effluent flowrate. The cell volume available for mixed liquor is the available volume of anoxic and aerobic cells minus the volume displaced by the media and the biofilm. At the South Adams plant, the volume of the cells is 4802 m3 (1.27 mil. gal). At the flowrate of 12 300 m3/d (3.2 mgd) observed in January 2007, the nominal HRT was 9.4 hours before correcting for the volume displaced by the media and the biofilm. After correcting for the media and biofilm, the nominal HRT was 7.7 hours. The MBBR was modeled for January 2007. The influent characteristics and operating temperature are shown in Table 11.11. The average primary effluent COD was 427 mg/L. The NH4-N and TKN were 34 and 44 mg/L, respectively. The majority of the sampling that is conducted on the MBBR is based on grab samples three or four times per week (which complicates dynamic modeling efforts); however, the plant recently expanded its sampling routine to include more frequent composite samples, in preparation for future plant upgrades. MBBR Recycle
Primary Effluent
Anoxic Basin 1 Media
Anoxic Basin 2 media
Aerobic Basin 1 With media
Sec Effluent
Aerobic Basin 2 With media
WAS
Figure 11.16 Schematic of one of two MBBR trains at the South Adams plant, Adams County, Colorado.
Table 11.11 SA-T1: MBBR influent data and cell data for dissolved oxygen, January 2007* (courtesy of Black & Veatch).
535
Flow (m3/d)
Flow Tempera- COD COD TKN TKN NH3-N NH3-N Dissolved Dissolved (mgd) ture (°C) (mg/L) (ppd) (mg/L) (ppd) (mg/L) (ppd) oxygen, ox1 oxygen, ox2 (mg/L) (mg/L)
1/1/2007
11 940
3.155
14.1
486
12 788 50.1
1319
38.1
1003
3.235
6.105
1/2/2007
12 250
3.236
14.25
375
10 108 52.7
1423
40.1
1081
3.67
6.36
1/3/2007
12 700
3.356
14.4
361
10 108 54.5
1526
41.44
1160
4.105
6.615
1/4/2007
12 000
3.172
13.7
382
10 108 49.6
1311
37.7
997
4.405
6.81
1/5/2007
11 430
3.019
13
295
7428 52.1
1311
39.6
997
4.705
7.005
1/6/2007
11 840
3.128
13.6
372
9708 50.3
1311
38.2
997
4.5375
6.8375
1/7/2007
12 130
3.204
13.6
363
9708 49.1
1311
37.3
997
4.5375
6.8375
1/8/2007
12 200
3.225
14.2
361
9708 48.8
1311
37.1
997
4.37
6.67
1/9/2007
11 900
3.15
14.35
370
9708 49.9
1311
37.9
997
4.5325
6.5575
1/10/2007
11 950
3.158
14.5
369
9708 41.6
1096
31.64
833
4.695
6.445
1/11/2007
11 600
3.066
13.95
380
9708 41.2
1054
31.3
801
5.1725
6.9225
1/12/2007
11 450
3.026
13.4
475
11 987 41.8
1054
31.7
801
5.65
7.4
1/13/2007
12 310
3.253
13.6
445
12 074 38.8
1054
29.5
801
5.0475
7.1
1/14/2007
12 510
3.306
13.6
438
12 074 38.2
1054
29.0
801
5.0475
7.1
1/15/2007
12 680
3.351
13.8
432
12 074 37.7
1054
28.7
801
4.445
6.8
1/16/2007
11 810
3.121
13.9
464
12 074 40.5
1054
30.8
801
4.6875
6.7075
1/17/2007
11 750
3.105
14
466
12 074 39.1
1011
29.68
769
4.93
6.615
1/18/2007
12 130
3.204
13.5
452
12 074 38.9
1039
29.6
790
5.0075
6.8975 (continued)
Table 11.11 Continued
536
Flow (m3/d)
Flow Tempera- COD COD TKN TKN NH3-N NH3-N Dissolved Dissolved (mgd) ture (°C) (mg/L) (ppd) (mg/L) (ppd) (mg/L) (ppd) oxygen, ox1 oxygen, ox2 (mg/L) (mg/L)
1/19/2007
11 690
3.089
13
472
12 160 40.3
1039
30.7
790
5.085
7.18
1/20/2007
12 220
3.229
13.1
456
12 275 38.6
1039
29.3
790
4.9375
7.115
1/21/2007
12 410
3.278
13.1
449
12 275 38.0
1039
28.9
790
4.9375
7.115
1/22/2007
12 000
3.171
13.2
464
12 275 39.3
1039
29.9
790
4.79
7.05
1/23/2007
11 850
3.132
13.5
470
12 275 39.8
1039
30.2
790
4.8375
7.09
1/24/2007
11 950
3.157
13.8
466
12 275 40.5
1067
30.8
811
4.885
7.13
1/25/2007
12 240
3.235
13.8
455
12 275 43.1
1164
32.8
884
4.8175
6.98
1/26/2007
11 960
3.161
13.8
470
12 390 44.1
1164
33.5
884
4.75
6.83
1/27/2007
12 020
3.176
13.3
451
11 955 43.9
1164
33.4
884
4.6675
6.8825
1/28/2007
12 760
3.371
13.3
425
11 955 41.4
1164
31.5
884
4.6675
6.8825
1/29/2007
12 120
3.203
12.8
448
11 955 43.6
1164
33.1
884
4.585
6.935
1/30/2007
11 560
3.055
12.4
469
11 955 45.7
1164
34.7
884
4.9925
7.1175
1/31/2007
11 500
3.038
12
472
11 955 49.7
1260
37.8
958
5.4
7.3
Average, Jan 2007
12 040
3.18
13.57
427
11 329 44.0
1165
33.4
885
4.71
6.88
*mgd × 3.785 = ML/d and ppd = lb/d; lb/d × 0.4536 = kg/d.
Development and Application of Models
The MBBR cells were operated at fairly high dissolved oxygen levels. The average dissolved oxygen concentration in the first aerobic cell (average of two trains) was 4.7 mg/L. The average dissolved oxygen concentration in the second cell was 6.9 mg/L. The MBBR effluent data is shown in Table 11.12, SA-T2; the TSS is the MLSS of the MBBR effluent before settling. During this month, the plant was at the threshold of nitrification, with an effluent NH4-N concentration of 2.5 mg/L. The amount of media was just sufficient to maintain nitrification. The plant currently does not have an ammonia limit for the month of January; ammonia permit limits are enforced only 6 months of the year, with concentrations ranging from 10 to 24 mg/L. The effluent nitrate-nitrogen concentration for the calibration period was 10 mg/L, but the plant currently does not have a nitrate discharge limit. The following sections discuss modeling the MBBR with GPS-X and Aquifas models. The Aquifas model was applied in both the semi-empirical and biofilmdiffusion modes. The electronic copy of the semi-empirical model, as set up for South Adams, is available as an attachment to this chapter. The model can be downloaded from the website that hosts the model (http://www.aquifas.com).
3.2.1 Moving-Bed Biofilm Reactor Modeling with GPS-X 3.2.1.1 Introduction The past several versions of GPS-X have contained attached-growth models for trickling filters, rotating biological contactors, submerged biological contactors, BAFs, and a “hybrid” model for IFAS and MBBR systems (Hydromantis, Inc., 2007). The current version (5.0) has several new features, including a more sophisticated aeration model, the ability to create controllers and graphs with a simple “drop and drag,” and the ability to create custom reports. The “hybrid” model was developed for an MBBR system, though it can be applied easily to an IFAS application. The GPS-X simulator contains several models, including the activated sludge models (ASM) 1, 2, 2d, and 3; the Mantis; the two-step Mantis; and the newer versions of the model. However, the “hybrid” model currently is only available using the Mantis model (for modeling carbon and nitrogen) and the ASM2d model (for modeling carbon, nitrogen, and phosphorus). Unless it is critical to model phosphorus, it is recommended to use the Mantis model for modeling IFAS and MBBR systems with GPS-X, because it tracks 16 state variables (as opposed to 27 with ASM2d) and converges more rapidly, with less manipulation. The biofilm model in GPS-X is based on the work done by Spengel and Dzombak (1992), but was adapted to the hybrid system (Hydromantis, Inc., 2007). The model
537
538
Biofilm Reactors
includes soluble material diffusion, biofilm growth, and particulate attachment and detachment in a one-dimensional plane perpendicular to the media surface. Suspended-growth and substrate use occur based on the Mantis (or ASM2d) model. Each layer of the biofilm (6 is the default) is modeled as a continuously stirred tank reactor, with the same reactions that occur in the bulk liquid. Attachment and detachment coefficients provide a means for particulate variables to transfer between the bulk liquid and the biofilm layers. The concentrations of the particulate variables are Table 11.12 SA-T2: MBBR effluent data (courtesy of Black & Veatch).* Effluent BOD mg/L
Effluent TSS mg/L
1/1/07
9
10.8
1/2/07
10
11.6
1/3/07
10
12.2
1/4/07
9
11.4
Effluent Effluent MBBR TSS NH3-N mg/L NO3-N mg/L mg/L
4.76
1/5/07
8.145
118.5
8.78
119
9.815
129
8.065
116
9.25
116
1/6/07 1/7/07
9
14.8
1/8/07
9
12.4
1/9/07
9
11.2
1/10/07
8
10.6
1/11/07
7
11
1.79
1/12/07
12.3
135
1/13/07 1/14/07
11
15.4
1/15/07
11
13.4
1/16/07
9
11.2
1/17/07
10
13
1/18/07
11
15.3
1/19/07
8.48 1.46
124
12.1
123
13.6
129
1/20/07 (continued)
Development and Application of Models
Table 11.12 Continued Effluent BOD mg/L
Effluent TSS mg/L
1/21/07
10
18.67
1/22/07
9
13.33
1/23/07
8
11.67
1/24/07
8
12.2
1/25/07
10
12.6
Effluent Effluent MBBR TSS NH3-N mg/L NO3-N mg/L mg/L
1.74
1/26/07
9.42
126.5
9.95
127.5
9.25
129
8.9
137
1/27/07 1/28/07
11
17.57
1/29/07
11
13
1/30/07
11
14
1/31/07
12
13.67
2.74
11.6
130
Averages
10
13.1
2.50
10.0
126
*Note: MBBR TSS represents effluent before settling.
converted to a volume based on the user-specified “dry material content of biofilm” and “density of the biofilm.” The user also specifies the “maximum biofilm thickness” and the “number of biofilm layers,” which limit the amount of growth that occurs before detachment increases. A summary of the biofilm parameters is provided in Table 11.13. The influent parameters are presented in Table 11.14. The Mantis model within GPS-X is identical to ASM1 (Henze et al., 2000), with the following exceptions: • Additional growth processes that account for the observed growth during conditions of low ammonia and high nitrate concentrations. Under these conditions, microorganisms can take up nitrate as a nutrient source. • Temperature-dependent kinetics, described by an Arrhenius equation. • Aerobic denitrification was added to model the increased level of denitrification that often is observed as a result of incompletely mixed tanks or dissolved oxygen gradients within flocs. The modification includes an anoxic oxygen half-saturation coefficient, which allows anoxic growth rates to be adjusted independently from aerobic growth.
539
Table 11.13 SA-T3: Principal GPS-X biofilm parameters (Hydromantis, Inc., 2007). Name
Units
Default value
Reactor portion filled by media
m3/m3
0.5
Typical use
Number of biofilm layers (plus one) –
6
Specific surface of media
1/m
500
Obtain from media manufacturer
Water displaced by media
m3/m3
0.18
Obtain from media manufacturer
Specific density of media
kg/m3
940
Obtain from media manufacturer
Attached biofilm thickness
mm
0.05
Maximum biofilm thickness
mm
Density of biofilm
g/cm
1.02
Dry material content of biofilm
–
0.1
Reduction in diffusion in biofilm
–
0.5
Attachment rate
m/d
Detachment rate
kg/m ·d
0.07
Anoxic shear reduction factor
–
1
Internal solids exchange rate
L/(m ·d) 0.02
0.5 3
Affects inert fraction of biofilm
0.5 2
2
Table 11.14 SA-T4: Summary of GPS-X influent fractions (Mantis and ASM2d) (Hydromantis, Inc., 2007). Mantis model parameters
Unit
Default value
XCOD/VSS
–
2.2
BOD5/BOD ultimate
–
0.66
VSS/TSS
–
0.6
Soluble fraction of total COD
–
0.35
Inert fraction of soluble COD
–
0.35
Substrate fraction of particulate COD
–
0.75
Non-biodegradable fraction of particulate COD
–
0 (continued)
540
Table 11.14 Continued Mantis model parameters
Unit
Default value
Heterotrophic biomass fraction of particulate COD
–
0
Autotrophic biomass fraction of particulate COD
–
0
Ammonium fraction of soluble TKN
–
0.9
Inert fraction of soluble TKN
–
0
Metal-hydroxide fraction of inorganic suspended solids
–
0
Metal-phosphate fraction of inorganic suspended solids
–
0
Nitrogen content of active biomass (mg N/mg COD biomass)
–
0.068
Nitrogen content of endogenous/inert mass
–
0.068
Phosphorus content of active biomass
–
0.021
Phosphorus content of endogenous/inert mass
–
0.021
ASM2d model parameters
–
XCOD/VSS
–
2.2
BOD5/BOD ultimate
–
0.66
VSS/TSS
–
0.6
Soluble fraction of total COD
–
0.35
Inert fraction of soluble COD
–
0.35
VFA fraction of soluble COD
–
0
Substrate fraction of particulate COD
–
0.75
Heterotrophic biomass fraction of particulate COD
–
0
Autotrophic biomass fraction of particulate COD
–
0
Poly-phosphate biomass fraction of particulate COD
–
0
Poly-hydroxy alkanoate fraction of particulate COD
–
0
Ortho-phosphate fraction of soluble phosphorus
–
0.9
Particulate xpp fraction of particulate phosphorus
–
0
Ammonium fraction of soluble TKN
–
0.9
Inert fraction of soluble TKN
–
0
Metal-hydroxide fraction of inorganic suspended solids
–
0
Metal-phosphate fraction of inorganic suspended solids
–
0 (continued)
541
542
Biofilm Reactors
Table 11.14 Continued Mantis model parameters
Unit
Default value
Nitrogen content of active biomass
–
0.07
Nitrogen content of particulate inert material
–
0.02
Nitrogen content of particulate substrate
–
0.04
Nitrogen content of soluble inert material
–
0.01
Nitrogen content of fermentable substrate
–
0.03
Phosphorus content of active biomass
–
0.02
Phosphorus content of particulate inert material
–
0.01
Phosphorus content of particulate substrate
–
0.01
The ASM2d model within GPS-X is identical to that described by Henze et al. (2000). The ASM2d is an extension of the ASM1 model and includes equations for biological phosphorus removal. Within GPS-X, the modeler has the ability to enter the influent characteristics based on BOD, COD, state variables, sludge only, or a TSS/COD combination. Regardless of how the inputs are entered, GPS-X performs all computations based on COD; thus, as with all models, it is important to have a good understanding of the influent stoichiometry. The Mantis and ASM2d influent parameters are summarized in Table 11.14. Please refer to www.hydromantis.com for additional guidance in selecting appropriate influent parameters.
3.2.1.2 Example The layout of the plant, as applied in the GPS-X model, is shown in Figure 11.17. The GPS-X model calibrated reasonably well to the plant operational data. The steadystate model results are summarized in Table 11.15; and dynamic model results are shown in Figure 11.18 (effluent NH4-N and NOx-N ), Figure 11.19 (MBBR effluent BOD5 and TSS, after settling in secondary clarifier), and Figure 11.20 (MBBR effluent MLSS before settling). The results show very good calibration and simulation of the plant biomass profiles and reasonable effluent . The facility recently has expanded its sampling regime to include more composite sampling at the MBBRs, to prepare for more stringent ammonia-discharge limits and a future nitrate limit.
3.2.2 Moving-Bed Biofilm Reactor Modeling with Aquifas The Aquifas model released in June 2008 was applied in both the semi-empirical and biofilm-diffusion modes. The model was run in both the steady-state and dynamic
Development and Application of Models
Figure 11.17 Layout of MBBR as modeled in GPS-X.
simulation modes. The steady-state was based on the average values for the month of January 2007 (Table 11.11). The dynamic simulation was based on the 31 days of data in January 2007 (Table 11.11). The reactor layout in Aquifas is shown in Figure 11.21. The steady-state model results are shown in Table 11.16. The numbers compared well with the actual plant data. The results of the dynamic simulation are shown in Figure 11.22 (effluent NH4-N and NOx-N ), Figure 11.23 (MBBR effluent BOD5), and Figure 11.24 (MBBR effluent MLSS before settling). The results showed good calibration and simulation to the plant data. The parameters for the biofilm modeling are presented in Table 11.16.
3.2.3 Moving-Bed Biofilm Reactor Modeling—General Comments The calibration of an MBBR to actual plant data in a dynamic simulation mode requires significant effort over and above that required to set up an MBBR model in
543
Table 11.15 South Adams County (Colorado) steady-state results for January 2007 with GPS-X. Data
GPS-X
% Variance Notes
Flowrate, mgd
3.18
3.18
0%
COD, mg/L
437
437
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
sCOD, mg/L
249
249
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
BOD, mg/L
no data
213
–
TSS, mg/L
128
128
0%
VSS, mg/L
No data
108
–
TKN, mg/L
No data
44
–
NH3-N, mg/L
34
34
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
Suspended TSS, mg/L
126
143
14%
Average of grab samples (measured in the morning, 3 to 4 times per week)
Suspended VSS, mg/L
No data
124
–
Dissolved oxygen (first cell), mg/L
4.65
4.65
0%
MBBR influent (includes sidestreams)
Average of grab samples (measured in the morning, 3 to 4 times per week)
544
MBBR Process
Average of grab samples (measured in the morning, 3 to 4 times per week)
Dissolved oxygen (second cell), mg/L
6.90
MLSS SRT (oxic), days
No data
0.2
–
MLSS SRT (total), days
No data
0.4
–
Temperature, °C Fixed biomass (first anoxic cell), g/m
2
Fixed biomass (second anoxic cell), g/m
2
Fixed biomass (first oxic cell), g/m
2
Fixed biomass (second oxic cell), g/m
2
WAS TS, lb/d
6.90
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
14
14
0%
10.0
10.3
3%
Measured once per month
7.8
7.7
–1%
Measured once per month
31.5
32.0
2%
Measured once per month
9.5
9.2
–3%
Measured once per month
No data
3447
–
Plant effluent
545
BOD, mg/L
9.7
8.3
TSS, mg/L
13.1
13.9
TKN, mg/L
No data
–15%
Composite samples measured 5 times per week
6%
Composite samples measured 5 times per week
3.5
–
NH3-N, mg/L
2.5
2.3
–8%
Composite samples measured once per week
NO3-N, mg/L
10.0
8.8
–12%
Composite samples measured twice per month
Effluent NH3N and NO3N (mg/L)
25.0
20.0
15.0
10.0
5.0
0.0 1/1/07
1/7/07
1/13/07
1/19/07
GPS-X Predicted Effluent NH3N GPS-X Predicted Effluent NOxN
1/25/07
2/1/07
Measured Effluent NH3N Measured Effluent NOxN
Figure 11.18 Comparison of MBBR effluent NH3-N and NOx-N from GPS-X model, with actual data. The NH3-N represents ammonium-nitrogen + ammonia-nitrogen; at pH ~7, >99% is as ammonium-nitrogen.
Effluent BOD and TSS (mg/L)
25.0
20.0
15.0
10.0
5.0
0.0 1/1/07
1/7/07
1/13/07
1/19/07
GPS-X Predicted Effluent BOD GPS-X Predicted Effluent TSS
1/25/07
2/1/07
Measured Effluent BOD Measured Effluent TSS
Figure 11.19 Comparison of MBBR secondary effluent BOD5 and TSS model results from GPSX model with actual data. 546
Mixed Liquor Suspended Solids (mg/L)
300 250 200 150 100 50 0 1/1/07
1/7/07
1/13/07
1/19/07
GPS-X Predicted MLSS
1/25/07
2/1/07
Measured MLSS
Figure 11.20 Comparison of MBBR effluent MLSS (before settling in secondary clarifier) from GPSX model with actual data.
Figure 11.21 Layout of South Adams MBBR as modeled in AquaNET (Windows. NET) version of Aquifas. 547
Table 11.16 South Adams County—Steady State Results for January 2007 with Aquifas Data
Aquifas
% Variance
Notes
MBBR influent (includes sidestreams) Flowrate, mgd
3.18
3.18
0%
548
COD (mg/L)
427
427
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
SCOD (mg/L)
249
249
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
BOD (mg/L)
No data
213
–
TSS (mg/L)
128
128
0%
VSS (mg/L)
No data
120
–
TKN (mg/L)
No data
44.0
–
NH3-N (mg/L)
34.0
34.0
0%
Average of samples, January
Suspended TSS (mg/L)
126
144
14%
Average of grab samples (measured in the morning, 3 to 4 times per week)
Suspended VSS (mg/L)
No data
134
–
Average of grab samples (measured in the morning, 3 to 4 times per week)
MBBR process
Dissolved oxygen (first cell) (mg/L)
4.65
4.65
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
Dissolved oxygen (second cell) (mg/L)
6.9
6.9
0%
Average of grab samples (measured in the morning, 3 to 4 times per week)
MLSS SRT (oxic), days
No data
0.2
–
MLSS SRT (total), days
No data
0.4
–
Temperature, °C
13.6
13.5
–1%
Monthly average
10.0
10.1
1%
Plant data measured once per month
7.8
0.8
–90%
Plant data measured once per month; model showed biofilm was nitrate-N limited
31.5
31.3
–1%
Measured once per month
9.5
13.2
39%
Measured once per month
3,204
–
From one-dimensional biofilm-diffusion model Fixed biomass (first anoxic cell), (g/m2) Fixed biomass (second anoxic cell), (g/m2) Fixed biomass (first oxic cell), (g/m2)
549
Fixed biomass (second oxic cell), (g/m2) WAS TS, (lb/d)
No data
Plant effluent BOD, (mg/L)
9.7
TSS, (mg/L)
13.1
TKN, (mg/L)
No data
9.0 14
–7%
Composite samples measured 5 times per week
7%
Composite samples measured 5 times per week
3.9
–
NH3-N, (mg/L)
2.5
2.3
–8%
Composite samples measured once per week
NO3-N, (mg/L)
10.0
9.5
–5%
Average of MLR grab samples (measured in the morning, 3 to 4 times per week)
Biofilm Reactors
Actual Plant Data versus Model
Effluent NH4N, Model & Plant, mg/L
Model Eff NH4N
NH4N
5.00 4.50 4.00 3.50 3.00 2.50 2.00 1.50 1.00 0.50 0.00
Date of Month
Actual Plant Data versus Model
Effluent NO3N, Model & Plant, mg/L
550
16.0 14.0 12.0 10.0 8.0 6.0 4.0 2.0 0.0
Model Eff NO3N
NO3N
Date of Month
Figure 11.22 Comparison of MBBR effluent NH3-N and NOx-N from Aquifas model with actual data.
steady-state or for a hypothetical plant that has no effluent data. The modeler must collect, analyze, and understand the plant data in greater detail to ensure that the data accurately reflects the operating conditions over the period of the dynamic simulation. Particular attention must be paid to the accuracy of the daily measurements, such as BOD5 and COD, NH4-N, and TKN in the influent and effluent; dissolved
Effluent BOD5, Model & Plant, mg/L
Development and Application of Models
Actual Plant Data versus Model Model Effluent BOD5
BOD5
14.00 12.00 10.00 8.00 6.00 4.00 2.00 0.00
Date of Month
MLSS & MLVSS, Model & Plant, mg/L
Figure 11.23 Comparison of MBBR secondary effluent BOD5 from Aquifas model with actual data.
Actual Plant Data versus Model Model MLSS
Plant MLSS
200 150 100 50 0
Date of Month
Figure 11.24 Comparison of MBBR effluent MLSS (before settling in secondary clarifier) from Aquifas model with actual data.
oxygen measurements in the aerobic cells; recycles; and MLSS. One must examine several months of data to understand the operation of the plant; also, one must identify if there are some unusual events during the period that need to be accounted for. Unusual events may include periods of heavy loadings from solids dewatering or a
551
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Biofilm Reactors
malfunction of equipment. These events would be reflected in the plant effluent data and the inputs to the model.
3.2.4 Integrated Fixed-Film Activated Sludge and Moving-Bed Biofilm Reactor Modeling—General Observations Modeling IFAS and MBBR systems is more complex than activated sludge systems. This is because of the need to calibrate the model to not only the activated sludge parameters, but also the biofilm parameters. Biofilm models are inherently more complex than activated sludge models and require more processor resources and simulation time. While good modeling practice typically recommends a plant-wide model, including biosolids processing, it may not always be practical to model all processes at a plant simultaneously, because dynamic biofilm modeling alone may take several hours, depending on the model and length of simulation. Therefore, the engineer must weigh model complexity with the amount of data that is available and the objective of the modeling exercise. More time, effort, and sampling should be spent on modeling that will be used for full-scale design or operations than on modeling used for facility planning or technology comparison. Modeling IFAS and MBBR systems in a dynamic simulation mode is much more difficult and time-intensive than for steady-state. The calibration of the model to actual plant data requires an additional level of effort when one is examining the trends (day-to-day variation) in addition to average levels. The calibration to actual plant data in a dynamic simulation mode allows the user to calibrate the biofilm parameters in greater detail than would be possible in a steady-state mode. This may be important in IFAS and MBBR plants, because differences in the media and aeration design can result in differences in the biofilm specific surface area, thickness of the biofilm and boundary layer, and performance of the plants. One observation that can be made based on the experiences with the modeling undertaken as part of this chapter is that the engineer and owner will benefit if the selected model has been calibrated to full-scale plant(s) that are representative of the plant being designed. Unlike a full-scale plant, a bench-scale or small pilot-scale study of a plant may not be able to duplicate the mixing and biofilm hydrodynamic properties observed in a full-scale system. A system operating with a different media or aeration system may not be representative of the plant being designed. The hydrodynamic properties are important, because they affect biofilm thickness, shear, and specific surface area, which, in turn, affect nitrification.
Development and Application of Models
A second observation that can be made is that there are differences in the approaches to biofilm modeling in different process models. These result in differences in the accuracy with which the models are able to represent different parameters. Each model has its strengths and weaknesses. This may be why some engineers prefer to use multiple models to simulate a plant and improve the level of confidence in the analysis or the design.
4.0 REFERENCES Boltz, J. P.; Johnson, B. R.; Daigger, G. T.; Sandino, J.(2009) Modeling Integrated Fixed-Film Activated Sludge (IFAS) and Moving Bed Biofilm Reactors— Mathematical Treatment and Model Development. Water Environ. Res., 81, 555–575. Dold, P.; Takacs, I.; Mokhayeri, Y.; Nichols, A.; Hinojosa, J.; Riffat, R.; Bailey, W.; Murthy, S. (2007) Denitrification with Carbon Addition—Kinetic Considerations. Proceedings of Nutrient Removal 2007: The State-of-the-Art, Baltimore, Maryland, March 4–7; Water Environment Federation: Alexandria, Virginia, 218–238. Dold, P. L. (1991) Incorporation of Biological Excess Phosphorus Removal in a General Activated Sludge Model. Department of Civil Engineering and Engineering Mechanics, McMaster University: Hamilton, Ontario, Canada. Downing, L. S.; Nerenberg, R. (2007) Evaluation of the Hybrid Membrane Biofilm Process (HMBP) for Nitrification and Denitrification of Wastewater. Proceedings of Nutrient Removal 2007: The State-of-the-Art, Baltimore, Maryland, March 4–7; Water Environment Federation: Alexandria, Virginia, 1127–1138. Envirosim Associates, Ltd. (2006) BioWinTM User Manual; Envirosim Associates, Ltd.: Hamilton, Ontario, Canada. Hem, L.; Rusten, B.; Odegaard, H. (1994) Nitrification in a Moving Bed Reactor. Water Res., 28, 1425–1433. Henze, M.; Gujer, W.; Mino, T.; Loosdrecht, M. V. (2000) Activated Sludge Models ASM1, ASM2, ASM2d and ASM3; International Water Association: London, United Kingdom. Huhtamäki, M. (2007) IFAS Process for Nutrient Removal. Practical Experiences and Guidelines for Design, Modeling and Operation. Ripesca-projecti LIFE03
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ENV/FIN000237; Raiseo Group Project Report. http://www.raisio.fi/ripesca/; Raisio Finland. Hydromantis, Inc. (2007) GPS-X Model. http://www.hydromantis.com (accessed June 2008). Johnson, T. L.; Shaw, A.; Landi, A.; Lauro, T.; Butler, R.; Radko, L. (2007) A PilotScale Comparison of IFAS and MBBR to Achieve Very Low Total Nitrogen Concentrations. Proceedings of Nutrient Removal 2007: The State-of-the-Art, Baltimore, Maryland, March 4–7; Water Environment Federation: Alexandria, Virginia, 521–535. Mamais, D.; Jenkins, D.; Pitt, P. (1993) A Rapid Physical–Chemical Method for the Determination of Readily Biodegradable Soluble COD in Municipal Wastewater. Water Res., 27, 195–197. Marais, G. v. R.; Ekama, G. A. (1976) The Activated Sludge Process: Part I—Steady State Behaviour. Water SA, 2, 164–200. McClintock, S. A.; Sherrard, J. H.; Novak, J. T.; Randall, C. W. (1988) Nitrate versus Oxygen Respiration in the Activated Sludge Process. J. Water Pollut. Control Fed., 60, 342–350. Melcer, H.; Dold, P. L.; Jones, R. M.; Bye, C. M.; Takacs, I.; Stensel, H. D.; Wilson, A. W.; Sun, P.; Bury, S. (2003) Methods for Wastewater Characterization in Activated Sludge Modelling; Water Environment Research Foundation: Alexandria, Virginia. Odegaard, H. (2005a) Combining CEPT and Biofilm Systems. Proceedings of the IWA Specialized Conference on Nutrient Management in Wastewater Treatment Processes and Recycle Streams, Krakow, Poland, Sept 19–21; International Water Association: London, United Kingdom. Odegaard, H. (2005b) Innovations in Wastewater Treatment—The Moving Bed Biofilm Process. Norwegian University of Science and Technology, Department of Hydraulic and Environmental Engineering: Trondheim, Norway. Odegaard, H.; Rusten, B.; Westrum, T. (1994) A New Moving Bed Biofilm Reactor—Application and Results. Proceedings of the 2nd International Specialized Conference on Biofilm Reactors, Paris, France, Sept 29–Oct 1; International Association on Water Quality: London, United Kingdom, 221–229.
Development and Application of Models
Randall, C. W.; Barnard, J. L.; Stensel, D. H. (1992) Design and Retrofit of Wastewater Treatment Plants for Biological Nutrient Removal; Technomic Publishing Company: Lancaster, Pennsylvania. Reichert, P. (1998) AQUASIM 2.0, User Manual; Swiss Federal Institute for Environmental Science and Technology (EAWAG): Dübendorf, Switzerland. Reichert, P.; Wanner, O. (1997) Movement of Solids in Biofilms: Significance of Liquid Phase Transport. Water Sci. Technol., 36 (1), 321–328. Rooney, T. C.; Huibregtse, G. L. (1980) Increased Oxygen Transfer Efficiency with Coarse Bubble Diffusers. J. Water Pollut. Control Fed., 52, 2315–2326. Rutt, K.; Seda, J.; Johnson, C. H. (2006) Two Year Case Study of Integrated Fixed Film Activated Sludge (IFAS) at Broomfield, CO, WWTP. Proceedings of the 79th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 21–25; Water Environment Federation: Alexandria, Virginia, 225–239. Sen, D. (1995) COD Removal, Nitrification and Denitrification Kinetics and Mathematical Modeling of Integrated Fixed Film Activated Sludge Systems. Ph.D. Dissertation, Virginia Polytechnic Institute and State University, Blacksburg, Virginia. Sen, D.; Boltz, J. P.; Copithorn, R. R.; Morgenroth, E. (2009) Proceedings of the Biofilm Reactors Workshop, 82nd Annual Water Environment Federation Technical Exposition and Conference, Orlando, Florida, Oct 10–14; Water Environment Federation: Alexandria, Virginia. Sen, D.; Copithorn, R.; Randall, C. W. (2006) Successful Evaluation of Ten IFAS and MBBR facilities by Applying the Unified Model to Quantify Biofilm Surface Area Requirements for Nitrification, Determine Its Accuracy in Predicting Effluent Characteristics, and Understand the Contribution of Media Towards Organics Removal and Nitrification. Proceedings of the 79th Annual Water Environment Federation Technical Exposition and Conference, Dallas, Texas, Oct 21–25; Water Environment Federation, Alexandria, Virginia, 185–199. Sen, D.; Copithorn, R.; Randall, C.; Jones, R.; Phago, D.; Rusten, B. (2000) Investigation of Hybrid Systems for Enhanced Nutrient Control, Project 96-CTS-4; Water Environment Research Foundation: Alexandria, Virginia.
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Sen, D.; Randall, C. W. (2007) Improving the Aquifas (Unified) Computational Model for Activated Sludge, IFAS and MBBR Systems by Embedding a MultiLayer Biofilm Diffusion Model within a Multi-Cell Activated Sludge System. Proceedings of Nutrient Removal 2007: The State-of-the-Art, Baltimore, Maryland, March 4–7; Water Environment Federation: Alexandria, Virginia, 1270–1299. Sen, D.; Randall, C. W. (2008a) Improved Computational Model (AQUIFAS) for Activated Sludge, IFAS and MBBR Systems, Part I: Semi-Empirical Model Development. Water Environ. Res., 80, 439–453. Sen, D.; Randall, C. W. (2008b) Improved Computational Model (AQUIFAS) for Activated Sludge, IFAS and MBBR Systems, Part II: Biofilm Diffusional Model. Water Environ. Res., 80, 624–632. Sen, D.; Randall, C. W. (2008c) Improved Computational Model (AQUIFAS) for Activated Sludge, IFAS and MBBR Systems, Part III: Analysis and Verification. Water Environ. Res., 80, 633–645. Sen, D.; Randall, C. W. (2005) Unified Computational Model for Activated Sludge, IFAS and MBBR Systems. Proceedings of the 78th Annual Water Environment Federation Technical Exposition and Conference, Washington, D.C., Oct 29–Nov 2; Water Environment Federation: Alexandria, Virginia, 3889–3904. Sen, D.; Randall, C. W.; Copithorn, R. R.; Huhtamaki, M.; Farren, G.; Flournoy, W. (2007) Understanding the Importance of Aerobic Mixing, Biofilm Thickness Control and Modeling on the Success or Failure of IFAS Systems for Biological Nutrient Removal. Water Practice, 1 (5), 1–18. Spengel, D. B.; Dzombak, D. A. (1992) Biokinetic Modeling and Scaleup Considerations of Rotating Biological Contactors. Water Environ. Res., 64, 223–235. Solley, D. (2000) Upgrading of Large Wastewater Treatment Plants for Nutrient Removal; The Winston Churchill Memorial Trust of Australia: Canberra, Australia. Sriwiriyarat, T.; Randall, C. W.; Sen, D. (2005) Computer Program Development for the Design of Integrated Fixed Film Activated Sludge Processes. J. Environ. Eng., 131, 1540–1549. Wanner, O.; Eberl, H.; Morgenroth, E.; Noguera, D.; Picioreanu, C.; Rittman, B.; Loosdrecht, M. V. (2006) Mathematical Modeling of Biofilms. IWA Task Group
Development and Application of Models
on Biofilm Modeling, Scientific and Technical Report 18. IWA Publishing: London, United Kingdom. Wanner, O.; Reichert, P. (1996) Mathematical Modeling of Mixed-Culture Biofilms. Biotechnol. Bioeng., 49, 172–184. Water Pollution Control Federation (1988) Aeration, Manual of Practice No. FD-13; Water Pollution Control Federation: Alexandria, Virginia. Weiss, J. S.; Alvarez, M.; Tang, C.; Horvath, R. W.; Stahl, J. F. (2005) Evaluation of Moving Bed Biofilm Reactor Technology for Enhancing Nitrogen Removal in a Stabilization Pond Treatment Plant. Proceedings of the 78th Annual Water Environment Federation Technical Exposition and Conference, Washington, D.C., Oct 29–Nov 2; Water Environment Federation: Alexandria, Virginia, 3889–3904. Wentzel, M. C.; Ekama, G. A.; Marais, G. v. R. (1991) Kinetics of Nitrification Denitrification Biological Excess Phosphorus Removal Systems—A Review. Water Sci. Technol., 23, 555–565.
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Index A Activated sludge, 260 Algae, 22 Anammox biofilm reactors, 405 Aquifas, integrated fixed-film activated sludge modeling, 518 moving-bed biofilm reactor modeling, 542 B Bacteria, 26 aerobic growth and respiration, 42 anaerobic respiration and fermentative metabolism, 44 anoxic conditions and respiration, 43 calculation of electrode reduction potentials and ΔG for halfreactions, 48 cell membrane, 30 cell wall, 30 chemical composition of cells, 32 chemoautotrophic metabolism, 38 chemoheterotrophic metabolism, 36 chromosome and plasmids, 28 co-metabolism, 50 cytoplasm, 30 energetics of respiration, 46 energy source, 36 Escherichia coli, 27 extracellular polymeric substances, 31 flagella, 31 nutrient requirements, 40 photosynthetic metabolism, 39 pili, 31 plasmids and chromosome, 28 theoretical demand of bacterial cells, 32 Bacterial, energy metabolism, 41
growth, 51 curve, 51, 52 effects of environmental variables, 54 enrichment cultures, 53 mixed cultures, 53 stability of mixed cultures, 54 growth kinetics in biofilms, 55 effect of temperature on organic removal, 61 inhibition and toxicity, 63 mass-transfer-rate limitations, 64 physical and chemical changes, 59 rate of bacterial processes, 56 structured models, 59 temperature effects, 60 metabolism, nutrition, and respiration, 33 Biofilm, airlift suspension reactor, 402 flux rate, 278 microbial communities, 72 modeling, 285 rate model, 287 reactors with suspended carriers or granules, 402 transformations, 65 aerobic nitrification, 68 anoxic nitrification/denitrification, 69 biological phosphorus removal, 70 biology of nitrogen transformations, 67 chemoautotrophic processes, 67 chemoheterotrophic processes, 66 denitrification, 67 hydrogen oxidation, 71 sulfide and sulfur oxidation, 71 Biological aerated filters, models, 471 Biological filters, 7, 325 backwashing and air-scouring, 345 biologically active filter process design, 349
559
560
Index
combined nitrification and denitrification, 362 design example, 353, 360 nitrification, 355 phosphorus removal considerations, 373 secondary treatment, 351 tertiary denitrification, 365 design considerations, 374 backwash handling facilities, 374 biologically active filter process aeration, 375 oxygen-transfer efficiency, 375 process air control, 377 process air distribution systems, 377 preliminary and primary treatment, 374 supplemental carbon feed requirements, 378 downflow biologically active filter with sunken media, 331 history, 329 media for use in biologically active filters, 341 mineral media, 341 modular plastic media, 345 random plastic media, 343 moving-bed, continuous backwash filters, 337 non-backwashing, open-structure media filters, 339 upflow biologically active filter with floating media, 335 upflow biologically active filter with sunken media, 334 Biologically active filter case studies, 379 Biologically enhanced phosphorus removal (BEPR), 261 Biology of fixed-growth process, 15 Bioreactor configuration (see Bioreactors), 93 Bioreactors, classification, 96, 97 configuration, 93
hydraulic application, 98 standard process flow diagrams, 93 BioWin, integrated fixed-film activated sludge modeling, 527 C Calculation, design of a submerged, upflow biological aerated filter system for nitrification following secondary treatment, 360 design of a submerged, upflow biological aerated filter system for secondary treatment (no nitrification), 353 effect of temperature on organic removal, 61 electrode reduction potentials and ΔG for half-reactions, 48 rotating biological contactors, 202 advanced secondary treatment, 203 secondary treatment, 202 theoretical demand of bacterial cells, 32 Captor process, 267 Case studies, biologically active filter, 379 hybrid process, Annapolis Water Reclamation Facility, Anne Arundel County, Maryland, 288 Broomfield Wastewater Treatment Plant, Broomfield, Colorado, 307 Colony Wastewater Treatment Plant, Colony, Texas, 308 Westerly Wastewater Treatment Plant, Westerly, Rhode Island, 299 moving-bed biofilm reactors, 235 Gardemoen Wastewater Treatment Plant, Gardemoen, Norway, 250 Harrisburg Wastewater Treatment Plant, Harrisburg, Pennsylvania, 238 Klagsham Wastewater Treatment Plant, Malmö, Sweden, 246
Moa Point Wastewater Treatment Plant, Wellington, New Zealand, 235 Moorhead Wastewater Treatment Facility, Moorhead, Minnesota, 240 Williams Monaco Wastewater Treatment Plant, Henderson, Colorado, 241 Chartered Institution of Water and Environmental Management model, 110 Clarification, 413 design approach, 417 clarifier enhancements, 420 considerations for nutrient removal sludges, 419 flocculation criteria, 424 settling regimes, 418 wastewater flocculation, 422 design details, 429 effluent weir and launder, 440 energy-dissipating inlet, 431 feed well (flocculating type), 433 floor slope, 438 influent column, 431 McKinney baffle, 440 side water depth, clear water zone, and overflow rate, 435 sludge collectors, 443 sludge hopper, 445 Stamford baffle, 440 design example, 446 interaction with other facilities, 458 international practices, 458 modeling, 458 process performance, 455 rectangular versus circular clarifiers, 445 secondary (integrated fixed-film activated sludge) clarifiers, 451 sludge hopper, 454 solids-separation choices, 416 Clarifier following moving-bed biofilm reactor, trickling filter, rotating biological contactor, and biotower, 448
Index
Co-metabolism, bacteria, 50 Contact beds, 4 D Denitrification, biofilm transformations, 67 modeling, 503 moving-bed biofilm reactors, 229 rotating biological contactors, 197 Denitrification and nitrification combined, biologically active filter process design, 362 Diseases caused by viruses, 25 E Eckenfelder model, 109 Effluent filtration, 463 process performance, 465 Emerging technologies, anammox biofilm reactors, 405 biofilm airlift suspension reactor, 402 biofilm reactors with suspended carriers or granules, 402 expanded granular sludge blanket, 404 internal circulation reactor, 404 membrane biofilm reactors, 406 upflow anaerobic sludge blanket, 404 Empirical model, 187 Energy source, bacteria, 36 Environmental variables, bacterial growth, 54 Equivalent-sludge-age, 285 Escherichia coli, 27 Example, calculation of electrode reduction potentials and ΔG for halfreactions, 48 clarification, 446 design of a submerged, upflow biological aerated filter system for nitrification following secondary treatment, 360 design of a submerged, upflow biological aerated filter system for secondary treatment (no nitrification), 353
561
562
Index
effect of temperature on organic removal, 61 rotating biological contactors, 202 advanced secondary treatment, 203 secondary treatment, 202 theoretical demand of bacterial cells, 32 Expanded granular sludge blanket, 404 F Fixed-film, 260 Fungi, 22 G Germain model, 108 GPS-X, moving-bed biofilm reactor modeling, 537 Gujer and Boller nitrifying trickling filter model, 122, 126 H History, 3 Hybrid processes, 7, 260 Annapolis Water Reclamation Facility, Anne Arundel County, Maryland, 288 Broomfield Wastewater Treatment Plant, Broomfield, Colorado, 307 Colony Wastewater Treatment Plant, Colony, Texas, 308 Westerly Wastewater Treatment Plant, Westerly, Rhode Island, 299 I Integrated fixed-film activated sludge (IFAS), 260 application, 267 biofilm flux rate, 278 cord media, 283 empirical methods, 285 kinetics, 269 mixed-liquor suspended solids, 281 models, 471 organics removal, 278 process design, 277
process-kinetics-based, 284 sponge-media, 283 web fixed-bed-media, 283 Internal circulation reactor, 404 International practices, clarification, 458 Invertebrates, multicellular, 23 K Kincannon and Stover model, 106 KMT Moving Bed Biofilm Reactor, 266 L Logan Trickling Filter Model, 111 M McKinney baffle, 440 Media, biomass growth, 273 breakage, 270 characteristics, 280 fixed, 267 free-floating, 264 growth, 268 looped-cord, 263 plastic-carrier elements, 265 rope, 263 sponge, 265 strand, 263 Membrane biofilm reactors, 406 Microbial communities in biofilms, 72 Model, biofilm rate, 287 Modeling applications to full-scale facilities, 512 clarification, 458 integrated fixed-film activated sludge, 513 Aquifas, 518 BioWin, 527 moving-bed biofilm reactor, 534 Aquifas, 542 GPS-X, 537 one- and two-dimensional biofilmdiffusion (steady-state and dynamic simulation), 504 ammonium-nitrogen, 506
biofilm thickness, growth, and fraction nitrifiers, 510 chemical oxygen demand, biomass, dissolved oxygen, and NOx-N, 510 semi-empirical equations for biofilm (steady-state and dynamic simulation), 478 ammonium-nitrogen uptake rate, 480 biomass production, 499 chemical oxygen demand removal, 494 denitrification, 503 fraction of nitrifiers, 502 oxygen, 504 Models bacterial growth kinetics in biofilms, 59 biological aerated filters, 471 Chartered Institution of Water and Environmental Management, 110 combining models, 113 comparison of model predictions, 190 Eckenfelder, 109 empirical model, 187 Galler and Gotaas, 105 Germain, 108 Gujer and Boller nitrifying trickling filter model, 122. 126 integrated fixed-film activated sludge, 471, 513, 552 Kincannon and Stover, 106 Logan Trickling Filter Model, 111 Monod kinetic model, 184 moving-bed biofilm reactors, 471, 534, 543, 552 nitrification, 194 Okey and Albertson nitrifying trickling filter model, 125, 128 Schulze, 107 second-order model, 186 selecting a model, 112 trickling filters, 103, 471 trickling filters, National Research Council, 103 Velz, 107
Index
Modified Ludzack-Ettinger (MLE), 261 Monod kinetic model, 184 Moving-bed biofilm reactors (MBBRs), 211 case studies, 235 Gardemoen Wastewater Treatment Plant, Gardemoen, Norway, 250 Harrisburg Wastewater Treatment Plant, Harrisburg, Pennsylvania, 238 Klagsham Wastewater Treatment Plant, Malmö, Sweden, 246 Moa Point Wastewater Treatment Plant, Wellington, New Zealand, 235 Moorhead Wastewater Treatment Facility, Moorhead, Minnesota, 240 Williams Monaco Wastewater Treatment Plant, Henderson, Colorado, 241 design considerations, 217 carrier biofilms, 218 carbonaceous matter removal, 220 combined pre- and postdenitrification, 230 denitrification, 229 high-rate designs, 220 low-rate designs, 221 nitrification, 224 normal-rate designs, 221 post-denitrification, 230 pre-denitrification, 229 mixers, 231 pretreatment, 232 general considerations, 234 approach velocity, 234 foaming, 234 media transfer and inventory management, 234 models, 471 plastic biofilm carriers, 215 solids separation, 233 Multicellular invertebrates, 23
563
564
Index
N National Research Council model, 103 Nitrification biologically active filter process design, 355 models, 194 moving-bed biofilm reactors, 224 Nitrification and denitrification combined, biologically active filter process design, 362 O Okey and Albertson nitrifying trickling filter model, 125, 128 Organism types in biological treatment, 18 P pH, rotating biological contactors, 180 Phosphorus removal, biofilm transformations, 70 Pilot-plant studies, rotating biological contactors, 207 Protozoa, 23 Pumping stations, trickling filters, 91, 144 R Rotating biological contactor (RBC), 5, 173, 262 denitrification application, 197 design examples, 202 advanced secondary treatment, 203 secondary treatment, 202 design methods, 184 comparison of model predictions, 190 empirical model, 187 manufacturers’ design curves, 189 Monod kinetic model, 184 predicted performance versus fullscale data, 191 second-order model, 186 temperature correction, 192 nitrification models, 194 physical design features, 198 biomass control, 201 covers, 201 drive systems, 200
hydraulics and flow control, 198 media, 199 physical layout, 198 tank volume, 198 pilot-plant studies, 207 problems and corrective actions, 204 inadequate treatment capacity, 205 excessive first-stage loadings, 205 excessive biomass growth, 206 loping of air-drive systems, 206 high clarifier effluent suspended solids, 207 corrosion of media supports, 207 process design considerations, 178 flow and loading variability, 182 media surface area, 179 operating temperature, 183 oxygen transfer, 180 pH and nutrient balance, 180 solids production, 183 toxic and inhibitory substances, 184 S Schulze model, 107 Second-order model, 186 Sludge collectors, clarification, 443 Sludge hopper, 454 Sludge hopper, clarification, 445 Solids production, 233 clarification, 416 rotating biological contactors, 183 Stamford baffle, 440 Startup, trickling filters, 153 Submerged biological contactor (SBC), 263 Systems fixed-media, 263 integrated fixed-film activated sludge, 260 T Temperature correction, rotating biological contactors, 192 Temperature rotating biological contactors, 183 trickling filter process design, 131
Temperature effects, bacterial growth kinetics in biofilms, 60 biologically active filter process design for nitrification, 360 Toxic substances, rotating biological contactors, 184 Transformations in biofilms (see Biofilm transformations), 65 Trickling filter/activated sludge process, 6 Trickling filters, 4, 81 biofilm carriers, 87 containment structure, 90 design considerations, 132 combined trickling filter and suspended-growth processes, 155 activated biofilter, 156 trickling filter/solids contact, 158 roughing filter/activated sludge, 162 biofilter/activated sludge, 162 trickling filter/activated sludge, 162 control mechanisms for macro fauna, 144 chemical treatment, 149 flooding, 148 operational strategies and facility improvements, 145 physical control, 152 Spülkraft, 147 distribution system, 133 electronic or mechanical drive rotary distributors, 136 hydraulic drive rotary distributors, 134, 137 optimizing rotary distributor operation, 138 pumping station or dosing siphon, 144 startup, 153 trickling filter media selection, 139 depth, 141 structural integrity, 142
Index
distribution system, 84 hydraulic and containment loading, 93 media, 88 models, 471 process design, 115 combined carbon oxidation and nitrification, 115 Gujer and Boller nitrifying trickling filter model, 122, 126 nitrifying trickling filters, 120 Okey and Albertson nitrifying trickling filter model, 125, 128 temperature and hydraulic application effects, 131 process flow sheets and bioreactor configuration (see Bioreactors), 93 process models, 103 Chartered Institution of Water and Environmental Management, 110 combining models, 113 Eckenfelder, 109 Galler and Gotaas, 105 Germain, 108 Kincannon and Stover, 106 Logan Trickling Filter Model, 111 National Research Council, 103 Schulze, 107 selecting a model, 112 Velz, 107 pumping stations, 91 underdrain system and ventilation, 90 ventilation and air supply alternatives, 100 mechanical ventilation, 101 natural draft, 101 U Upflow anaerobic sludge blanket, 404 V Velz model, 107 Viruses, 24 Viruses that cause disease, 25
565
